MetPlant 2011 - Metallurgical Plant Design and Operating Strategies

Published by:The Australasian Institute of Mining and MetallurgyGround Floor, 204 Lygon Street, Carlton Victoria 3053, Australia

METPLANT 2011Metallurgical Plant Design and Operating Strategies

8 - 9 AUGUST 2011

PERTH, WA

The Australasian Institute of Mining and MetallurgyPublication Series No 7/2011

© The Australasian Institute of Mining and Metallurgy 2011

All papers published in this volume were refereed prior to publication.

The conference convenors and the Institute as a body are not responsible for the facts and opinions advanced in this publication.

ISBN 978 1 921522 44 4

Compiled on CD ROM by:Visual Image Processing Pty LtdPO Box 3180 Doncaster East VIC 3109

Keynote Speakers

Guidelines for Economic Evaluation of Projects P Card 1

Keeping Projects on the Rails J Canterford 12

Predicting Variations in Mill Feed P McCarthy 19

Setting Processing Plant Projects up for Success in Australia R Young and A Hollonds 26

Building Skills

Reducing Mining and Mineral Processing Plant Fatality Rates N Cann, S Casey, R Mills and J Ross

27

Professional Development for Metallurgists – Improving Technical Skills

D Drinkwater, D Bradshaw, P Tilyard and P Munro

39

Differences between the Engineering Cultures of Australia and Brazil C Fountain, P Libânio and G Lane

49

Comminution

Crushers – An Essential Part of Energy Effi cient Comminution Circuits R Bearman, S Munro and C M Evertsson

66

Reducing the Energy Required in Grinding Clinker to Cement – Some Case Studies

H Benzer, N Aydogan, H Dundar and A J Lynch

86

The Importance of Evaluating Grinding Performance R L Koenig and K T Broekman 100

Increasing Capacity and Effi ciency of Grinding Circuits with High Frequency Screens

J Wheeler and B Packer 107

Comminution and Gravity

Recent Improvements to the Gravity Gold Circuit at Marvel Loch A Bird and M Briggs 115

In-Line Pressure Jig Preconcentration Plant at Pirquitas Mine A H Gray, G Delemontex, N Grigg and T Yeomans

138

Energy Effi ciency Opportunities in Milling – Improving Comminution Circuit Effi ciency

F Musa, M Stewart and G Weiss

154

Non-Contact Acoustic Measurement of Dynamic In-Mill Processes for SAG/AG Mills

R A Pax 163

IsaMill™ Design Improvements and Operational Performance at Anglo Platinum

C Rule and H de Waal 176

The Infl uence of Liner Wear on Milling Effi ciency P Toor, M Bird, T Perkins, M Powell and J Franke

193

CONTENTS

Design

Energy Effi ciency Assessments in Design – Seminal Decisions and Effective Processes

M Allen, N Rosaguti, B Innes and M Stewart

213

Filtration Test Work – Extracting the Whole Story for Studies and Design

G Bickert and B Länger 228

Sensible Cost Cutting for Resource Projects D Connelly 250

Integrating Sustainability Principles into Mineral Processing Plant Design

G Corder and S Green 264

Chirano Gold Mines Expansion – A Case Study S Ellis and I Dunlop 280

Engineering and Science in Flotation Cell Design J Euston 292

The Wonawinta Silver Project – Value Engineering at Work P G Greenhill 312

An Approach to High Solids Slurry Pipeline Design M Griffi ths and N Steward 328

The Integration of Geometallurgy with Plant Design G Harbort, G Cordingley and M Phillips

339

Cost Effective Concentrator Design G Lane, P Dakin and D Elwin 364

‘We’re Metallurgists, not Magicians!’ E McLean 374

In-Pit Tailings – World’s Best Practice for Long-Term Management of Tailings

G M Mudd, H D Smith, G Kyle and A Thompson

391

Flotation and Pyrometallurgy

Flotation Mechanism Design for Improved Metallurgical and Energy Performance

R Coleman and A Rinne 405

The Great Oil Debate – Does Quenching Oil have a Deleterious Effect on Flotation?

C J Greet and J Kinal 419

Carbon Prefl oat Circuit Improvements at Century Mine D G Rantucci, T J Akroyd and L J Grattan

430

Quantifying Plant Flow Availability for an Alumina Refi nery Expansion using Dynamic Simulation

B Reynolds and S Collins 442

Secondary Copper Processing using Outotec Ausmelt TSL Technology J Wood, S Creedy, R Matusewicz and M Reuter

460

Hydrometallurgy

Solvent Extraction of Uranium – Towards Good Practice in Design, Operation and Management

P Bartsch and S Hall 468

Fosterville Gold Mine Heated Leach Process M Binks and P Wemyss 480

Leach Residue and Pregnant Liquor Separation – Process and Capital Comparison of Counter-Current Decantation and Counter Current Washing with Vacuum Filtration

R Klepper and P McCurdie 488

Agglomeration – The Key to Success for the Murrin Murrin Ni Laterite Heap Leach

D Readett and J Fox 506

Scale Suppression using Swirl Flow Technology J Wu, D Nairn, B Nguyen, J Farrow and D Stegink

515

Plant Case Studies

On-Belt Analysis at Sepon Copper Operation T Arena and J McTiernan 527

The Casposo Gold-Silver Project – Process Selection and Design D Connelly and K Nilsson 536

Optimising Western Australian Magnetite Circuit Design D David, M Larson and M Li 552

Contract Commissioning and Operation of Western Areas’ Cosmic Boy Concentrator

C Dick, D Boska and C Fitzmaurice

563

Float it, Clean it, Depress it – Consolidating the Separation Steps at Clarabelle Mill

V Lawson and M Xu 589

Flotation Process Control Optimisation at Prominent Hill J Lombardi, N Muhamad and M Weidenbach

602

Ore Ageing Test Work for the Ok Tedi Skarns M Morey and R Cantrell 615

Author Index 630

Metallurgical Plant Design and Operating Strategies (MetPlant 2011) 1 8 - 9 August 2011 Perth, WA

Guidelines for Economic Evaluation of Projects

P Card1

MAusIMM, Peter Card Mining Evaluations, 48 Station Street, Aspendale Vic 3195. Email: [email protected]

ABSTRACT

“Let’s get discipline and quality into easy to understand evaluations!”

Project Managers and MetPlant Operators accept poor quality evaluations because they are generally unaware of what they should expect.

The 2010 draft AusIMM ‘Guidelines for Techno-Economic Evaluations’ were collated by a subcommittee of practitioners and are regarded as best practice.

They apply directly to projects and metallurgical plant evaluations, including those with a heavy technical basis. They can be applied to technical problem solving where there are no monetary computations.

Project Managers and MetPlant Operators should demand that economic models and evaluations follow the six principles. They should become working tools which are easy to understand, fit for purpose, consistent, rigorous, record sources of input data, have key graphs and are rapid to audit.

INTRODUCTION

Poor workmanship is common but usually accepted If industry made an award for the worst performance in MetPlant project design and plant operation then the odds-on favourite would be the economic evaluation!

If an economic evaluation is sophisticated, complex, and terribly clever with Excel so that only one or two experts can use it, then it probably is worst practice. But if anyone with only basic knowledge of evaluation can readily follow it, sees the correct data being employed and feels it is easy to understand then it is on the way to best practice. Project Managers and Plant Operators need to take control and demand economic evaluations they can quickly follow during their busy working day.

The world’s best practices are readily available on the Internet and are in active use in the mining industry. But for historical reasons most project managers and plant operators will accept poor quality work in this arena. What is accepted for economic evaluation would not be tolerated in its sister disciplines of geology, mining, metallurgy, engineering and accounting. Fortunately this is rarely due to sloppy management by project managers and plant operators, but due to lack of awareness of what they should expect and demand.

As a horrible start, most professionals do not even call it by its correct name of ‘economic evaluation’ but talk of ‘financial modelling’ or even more incorrectly ‘financial analysis’ (more later).


Project managers and plant operators would insist on metallurgists doing the processing study work, insist on engineers doing the design and epcm, would want accountants to do the books, but probably would accept almost anyone willing and able to perform the economic evaluation. This person probably would be allowed to do the evaluation however they believed was best. This is because the discipline of economic evaluation has evolved over recent decades to bridge between operations/engineering and accounting without an academic or professional foundation.

Older style project managers and plant operators see economic evaluation as a backroom activity to be hurriedly completed, typically when the last of the cost estimates are finished the night before the submission is due to management. They see the activity as a mathematical computation to fill in the paperwork with NPV, IRR, payback etc. Fortunately these are a dying breed.

These older style managers do not really understand they are designing or operating a business, but live in a closed world of professional engineering or hands-on plant operating. They are very confident that they are ‘working on the important stuff’ and ‘getting things done!’ They do not realise that contemporary managers demand an economic evaluation up and running from Day 1 as a tool to steer the project or operating plant through the study process into the optimum state of business. The economic evaluation specialist should be the second best role in the team; after the leader.

Mantras of best practice A few mantras of best practice in economic evaluation are: -

“If you do not readily understand and comprehend my evaluation then you do not have a problem, I do!”

“Every worksheet should be as easy to read as a school text book!”

“Do not try to impress with sophisticated Excel functions, but use your intelligence to convert complex interactions into simple steps on the worksheet.”

And definitely not “Trust me! I am the expert in evaluation modelling!”

Financial analysis versus economic evaluation Completion of a Project requires two separate money-focussed activities:

1. Economic Evaluation, and

2. Financial Modelling

The first, Economic Evaluation, is all about understanding the business health of the Project. What cash will be required to establish, operate and pay taxes versus the cash generated or saved by the project. This is the simple economics of cash-in and cash-out over time. Economic Evaluation does not worry where the cash comes from (i.e. financing) but rather it wants to understand the cash generating power of the underlying project including what sort of prices and operating environment is required to cover costs and generate an economic return. (The discount rate is usually before financing)

Everything is computed in cash in the year it is actually spent or received. There is no accounting depreciation, no accounting charges for future closure and no other non-cash items.


There is no equity raisings nor company borrowings, and so no financing charges on borrowings during construction. Taxes are included and computed before the apparent benefit of debt.

The second, Financial Modelling, is all about sourcing the cash to establish and run the project until it becomes self-sustaining. What mix of company internal cash flow, debt, new equity, convertible notes, derivatives and hedging will the company use to progressively pay for owner’s costs, epcm, first fills, commissioning, ramp-up to commercial operations and interest on these borrowings? How will each type of capital raising impact company profits, its balance sheet and share price?

Economic Evaluation might be best lead by a person with a technical-operating background whereas Financing might be best lead by an Accountant. Both perform spreadsheet modelling of the future business but they are very different in purpose and process. Both should be presented in simple language and easy-to-understand concepts.

The two activities should not be woven together because:

a) Each is a stand-alone decision.

b) Combining their mathematics is very tricky, especially adjusting the discount rate as debt is introduced. (Do people still get fooled by false claims of improving project returns by using debt?)

c) The spreadsheets will become unnecessarily complex to use and audit and so alienate all in the Team except specialists in finance.

The financing spreadsheets could be appended to the end of the economic evaluation workbook providing the flow is one-way and nothing feeds back to the economic evaluation worksheets.

Evaluation, valuation, modelling ‘Modelling’ is a component of, but not all of ‘Valuation’ which is a component of, but not all of ‘Evaluation’. They form a hierarchy with ‘modelling’ at the bottom providing the hands-on computations that feed results for various cases and scenarios up into ‘valuation’. This ‘valuation’ quantifies and characterises the value of the project or metplant. In turn this ‘valuation’ feeds up as one element in the intellectual activity of ‘evaluation’ by fully understanding the project or metplant as a business.

Anyone thinking that economic evaluation is all about pouring numbers into a spreadsheet model to get NPV and IRR is living in the past. Today it is all about having a working knowledge of the whole project or metplant from ore in the ground through all activities and influences to the market. It encompasses nearly everything from engineering to paying taxes. It is about understanding the key drivers and key interactions of the business. It is getting a helicopter view of the total entity, deciding how it fits the existing business and helping to test ideas and create better projects and metplants. It is all about putting the ‘E’ into Evaluation. Yes, the economic evaluation specialist should have the second best job in the team: after the leader.

Best practice in economic evaluation Best Economic Evaluation Practice in the mining industry is described in the draft guidelines released last year by the AusIMM: 2010 AusIMM Draft Guidelines for Technical Economic Evaluation of Minerals Industry Projects. They are awaiting feedback.


www.ausimm.com.au/content/docs/guidelines_tech_economic_evaluation.pdf - 2010-05-11

The first part defines what represents best practice in ‘spreadsheet modelling’, and has a worked example in pdf form in an appendix. The second part deals with key aspects of ‘valuation’ and ‘evaluation’. Together they represent a collation of sound practices that competent practitioners in economic evaluation have developed over the years.

Part 1 – Guidelines for Spreadsheet Modelling These Guidelines were developed for economic evaluation but apply equally to geologists, mining engineers, metallurgists, engineers, project manager and plant operators using a spreadsheet such as Excel.

There are six Principles at the top of the hierarchy.

KEY PRINCIPLES The key principles to which one should adhere when performing spreadsheet modelling are simple but extremely effective. They are:

• Easy to follow

• Tailored‐to‐Purpose

• Transparent

• Disciplined, Rigorous, and Consistent

• Recording Sources of All Data

• Rapid to Audit

CHARACTERISING EACH PRINCIPLE Each of these Principles has a profile. For example ‘Easy-to-follow’ is characterised as: -

• Intuitive and Visual

• Time to do it properly

• Architecture: Workbook Layout

• Worksheet Layout

• Modular Construction

DESCRIBING EACH PROFILE

Each of these profiles is described. For example ‘Worksheet Layout’ is defined as:

• A non‐expert should readily understand the function of each worksheet, how it is arranged into component parts, how the data is entered, the computations, and the relative importance of the parts.

• The visual flow down and across each worksheet should be intuitive and logical.


• Each worksheet should have a bold heading followed by a brief outline of its purpose, and where helpful, its important links with other worksheets.

• Sections within each worksheet should be in discrete work blocks, with obvious subsections and subheadings using a cascading layout for subtitles.

• Visually each work block should be self contained with an obvious step‐by‐step development toward a bold subtotal for that work block.

• The separate work blocks should be in a logical sequence down and across the worksheet and their aggregation is obvious.

• Complex and extended computations should be shown in a series of small steps so that the logic is visible and the input parameters are obvious. There is no need to interrogate the algorithm. Explain the logic of complex algorithms in a note.

• Usually if a row of data that has already been presented above is needed again in a work block then the entire row should be repeated so that there is visual flow of the logic. If referenced from another worksheet then it should be coloured coded.

• Key inputs and results should be shown in graphs as a self‐check and for rapid understanding.

• The Data Group and Outline facility can be used to define worksheet structure, collapse related groups of rows or columns to reduce visual clutter, and aid navigation.

WORKED EXAMPLE OF MODELLING This profile of ‘worksheet layout’ is illustrated by the worked example in the AusIMM draft Guidelines. One of its worksheets is extracted and reproduced in Appendix A below.

READY MADE EVALUATION MODELS There was unanimity amongst the AusIMM Members who generated the Guidelines that ready-made economic evaluation models, where users fill in the blanks, were too dangerous to use. Experience is that these black-box models transgress the six principles of Best Practice, but more importantly have a bad history. Their computations cannot be audited, the models must be exceedingly complex (or deficient) to accept a wide variety of scenarios and they simply cannot be trusted.

Part 2 – Best Practice Valuation and Evaluation The AusIMM Guidelines describe key issues in valuation and evaluation but are not intended to be a comprehensive ‘how to do” manual. Key topics covered are:

Theory and Computation of NPV, IRR and Risk Economic Evaluation and Financial Modelling Evaluation and Valuation Stand Alone or Incremental Value Stand Alone, Synergies, Using Tax Losses Gearing (Debt) Expected Values and Results Alternatives


Ranges and Likelihoods / Sensitivity Analysis Input Data Quality Assessment Real and Nominal Life of the Evaluation After Tax Risk and Uncertainty Materiality Evaluation Framework Process Product Prices, Treatment and Refining Charges, and Foreign Exchange

Forecasts Sales: Volumes, qualities, working stocks/capital Production: Volumes, head grades, recoveries, product grade, delivery Capital Costs Operating Costs Closure Costs Taxes (direct company level taxes) and Carbon Costs Discount Rate Auditing Graphs Helicopter View Avoid

These issues are summarised in the Guidelines and will not be repeated here.

WHAT TO EXPECT FROM SPREADSHEET MODELLING A Project Manager or MetPlant Operator should demand that every economic evaluation model under his/her management is: -

• Absolutely rigorous in its construction (example: follows AusIMM Guidelines 100%).

• In discrete blocks of simple steps with clear headings and obvious end results

• Input data is coloured (example: blue) so that it is immediately visible and the project manager can skim across it to quickly check validity of all inputs.

• Input data has its source (when, who and what) typed in the row above so the project manager can immediately see if it is the correct version.

• Data referenced across from another sheet is coloured (example: green) so it is immediately recognised

• Absolutely consistent across rows, so that algorithms do not have hidden changes

• Every item of input data exposed in a row before being used and absolutely no fresh data entered as data hidden in algorithms.

• Starts with a brief overview of results including multiple graphs of the four cash flows (see below) and all important inputs and outputs


• Audited by the specialist, by experts in their areas (example: metallurgist audited the processing section) and usually an external specialist

• Intuitive

• Trust in what you see without drilling down into algorithms.

USING EVALUATIONS IN FEASIBILITY STUDIES Project Managers and MetPlant Operators appear to use three phases of study for projects:

1. Concept or Scoping Studies to assess if the project fits company strategy and at least one alternative has a reasonable likelihood of being economically viable.

• For these brief studies the modelling and evaluation usually can be relatively coarse and simple. It needs to explore the range of outcomes and key drivers of success and failure. It needs to weed out pet projects and support only quality concepts.

• The evaluation person needs to work cooperatively between the project manager and the many experts inside and outside the company who have the knowledge. The specialist may be a ‘backroom’ type from any background, but better if an active specialist with an operating/engineering background.

• The project manager is likely to work closely with the evaluation specialist interacting every day or two.

2. Pre-Feasibility Studies that assess the complete range of project alternatives. The divergent thinking in this phase generates the most value.

For these studies the modelling and valuation needs to be detailed enough to differentiate the economics and character of each alternative. This phase excites those who are creative but objective. It needs to confirm the attractiveness of the selected alternative and define its business character.

The evaluation person needs to frequently interact with specialists inside the project team and outside, drawing out the complete information, going back and confirming it has been correctly modelled (audits) and being absolutely objective. The evaluation specialist must resist the temptation to be too clever with Excel functions but keep the model simple so everyone can readily understand, audit and feel it represents the alternative truly. The specialist should not be a ‘backroom’ type but probably from an operating/technical/engineering background with a broad understanding of the business and a work ethic that is energetic, collaborative, creative and accurate.

The project manager is likely to see the evaluation model as an engine to assess alternatives and generate ideas. He/she is likely to rely heavily on the evaluation specialist to understand where each of the alternatives is heading and to help steer the study activities to best effect. Daily contact, with creative thinking, would be common.

3. Final Feasibility Studies that thoroughly assess and define the alternative to be taken into execution or construction. This is convergent thinking with lots of detail.


For these studies the modelling and evaluation will become quite detailed. The whole business needs to be thoroughly defined. The modelling will evaluate different methods and different equipment within the selected alternative. It may need to work in nominal terms in some areas and integrate with company’s accounting and financing activities.

The evaluation person needs to work in detail as part of the enlarged study team. By this time the Project Manager should know if the existing evaluation person is the right fit for this phase. Contact may become periodic but the evaluation model would continue as a vehicle to help steer the Feasibility Study.

FOUR CASH FLOWS Anyone opening up a worst practice evaluation model is faced with a myriad of parameters and a convoluted array of computations that are understood only by the evaluation specialist. Key outputs such as NPV are buried amongst the worksheets. Auditing would be tedious and take days.

Best Practice has simplified valuations into four streams of cash flow:

1. Revenue Cash stream (production, stocks, sales, prices, debtors)

2. Capital Cash stream (capex, creditors, tax deductions)

3. Operating Cost Cash stream (opex, creditors)

4. Taxes Cash stream (Royalties and Income Tax)

Their sum represents the net cash flow each year and this can be simply discounted to give NPV, or used to compute IRR.

Second and third level computations such as working stocks, debtors, creditors and tax deductions for capex (‘tax depreciation’) are computed as high-level calculations within these four cash streams. Their impact on project managers’ and plant operators’ decision-making will be minimal so should be reduced to a few simplified rows.

THE TWO BOOKENDS OF ECONOMIC VALUE In most mining industry businesses there are two ‘bookends’ which dominate the economics: the ore resources in the ground and the market. One end determines how big and good the business can be and the other end determines how profitable it will be. In between are all the very important and exciting projects and operations to make it happen and make it improve. It has been joked that a whole team of engineers and metallurgists slave for ages on production, capex and opex while somewhere in a back office a few people generate price forecasts that swamp the valuation. None-the-less the production, capex and opex are critical and must be forecast with appropriate quality.

DO BIGGER PROJECTS NEED BIGGER ECONOMIC MODELS?

One of the world’s greatest iron ore mines was acquired in the 1970’s with the economic evaluation model provided to the company’s Board being just one page of very easy-to-follow, manual computations. Contrast it with the volumes of modelling and synthesis that would be required today. In a strange way that one page from the 1970’s was as potent as all the


evaluation study work that we generate today. There in amazing simplicity for the Board members was the heart of the acquisition decision, namely the risks in the forecasts of price, mineral resource, production and costs over the years.

In 2011 senior executives in the world’s biggest mining companies have made it clear they would love simple one page models of major investments, major acquisitions and life-of-mine plans.

This is not to suggest that all Evaluation models should be one page. Quite the reverse: a detailed evaluation model should be a centrepiece of a major Project or MetPlant. It should draw together all the component parts as a business so the Team understands the relative importance of each part, where to focus and how to optimise the overall design. It should be a tool used every day by members of the Project Team/MetPlant to challenge and test their ideas. Detailed evaluation models are needed to steer the Project toward the best configuration and to assess the risks.

Ironically, it would seem that Management/Board needs a simple, easy-to-understand evaluation “one page” model to help it understand the big decision of whether to invest in the Project whereas the Project Team might need a big, detailed, working model to optimise the Project’s configuration and the design of its component parts.

HALL MARKS OF BEST PRACTICE ECONOMIC EVALUATION From Day 1, Project Managers and MetPlant Operators must set their expectations and demand that economic evaluation is: -

Easy to follow – it may not be simple but anyone in the team should be able to follow it like a school text book

Fit for purpose in that detail matches importance – begin as simply as possible and add complexity only when warranted.

Rigorous, transparent and intuitive – trust what you see

Fully documented – readily see where every piece of data was sourced

Graphs – to find errors and to give quick visual understanding

Audited – both the mechanical computations and the results.


APPENDIX A – ILLUSTRATING BEST PRACTICE IN SPREADSHEET MODELLING

An extract from the worked example in AusIMM Guidelines .


Metallurgical Plant Design and Operating Strategies (MetPlant 2011) 12

8 - 9 August 2011 Perth, WA

Keeping Projects on the Rails

J Canterford1

1. FAusIMM(CP), Process Technologies Australia Pty Ltd, PO Box 210, Deloraine Tas 7304. Email: [email protected]

ABSTRACT

Although we have access to a wide range of sophisticated process design and improvement tools in combination with a significant range of case studies, the mineral resource industry still manages to be plagued by the negative publicity surrounding operating plant failures because they were neither smart nor safe, together with proposed projects that fail at the first or second hurdle. The general public and the finance industry are generally totally uniformed about projects that ramp-up to name plate capacity on time and on budget and even more so when ahead of schedule and under budget. As the saying goes, good news does not cut the mustard. The following is a summary of the author’s sometimes jaundiced view about why metallurgical plants do not live up to their technical and commercial expectations. In no way should the summary be construed as a negative reflection on the skills of the metallurgical profession. Rather, it is to be taken as a positive acknowledgement of those skills and is intended to highlight some of the major considerations that need to be assessed as a potential project transforms into a sustainable technical and commercial reality. While a considerable portion of the summary is directed at developing hydrometallurgical flowsheets, the general principles are equally important to all aspects of extractive metallurgy. For convenience, the summary covers the following five basic discussion points – project, people, process, patents and politicians.

INTRODUCTION

On any given day, every successful metallurgical plant can be characterised as being a positive compromise between a series of potentially conflicting criteria, the more significant including:

The mineralogical complexity of the feedstock, Market requirements, Owner/developer expectations, Technical complexity, On-going operating cost scenarios, and Sustainable environmental footprint.

Although we can talk about generic flowsheets, no two flowsheets are identical when dissected in detail. While we certainly can and should learn from past and current experience, what we have to accept is that the balance between the conflicting criteria will almost certainly alter during the life of any individual project. In turn, this means that the



“best” flowsheet should have a reasonable level of flexibility to cater for a realistic range of inputs and outputs, both metallurgical/environmental as well as corporate. Some of the anticipated variations may be self-induced while others will be a result of external forces over which the project owner/developer has no control. If the conflicting criteria cannot be balanced and that balance maintained then the outcome will almost certainly be closure or at least the requirement for major modifications to plant layout, operating procedures, changes to volume and specifications of outputs, etc. The capital and operating cost implications of such modifications can be mind blowing. Without going into any detail, experiences at Ravensthorpe and Goro are stark reminders of what can go wrong. As a person with no practical metallurgical plant design experience but is sometimes asked to comment on what, how and why a flowsheet appears to have failed or at least throw enough hand grenades at the evaluation stage of flowsheet development, the simplistic answer provided is itself multi faceted covering:

The chemical/mineralogical integrity of all testwork, Realistic technical risk analysis, Treatment of process and financial models with a fairly high degree of scepticism,

and Understand what the market will buy and why and how these requirements can be

met.

It will come as no surprise that I am very strongly of the view that hand grenade throwing and proper evaluation of the metallurgical plant design should be initiated at the scoping stage and should be a continuing exercise well into the operating phase at, or preferably above, name plate capacity.

SUCCESSFUL PLANT DESIGN DRIVERS

First and foremost, it is essential to match output volumes and qualities with a realistic assessment of market requirements. As noted below, some 15 years ago magnesium metal was considered to be the metal of the future. Numerous proposals were being bandied about. In Australia alone, the combined capacities of the proposed magnesium metal projects in Tasmania, Queensland, Victoria, South Australia and Western Australia between 1995 and 2005 was in excess of the 300% of then total world consumption and more than 200% of the projected demand in 2015. Individually each of the proponents considered their proposed projects to be superior to all others and that only they would attain commercial status. Of course reality was quite different and not one of those proposed Australian projects has seen the light of day. As I participated in some of these proposed projects I guess I have to confess to being guilty by association. Consistent with my novice status as a metallurgical plant “expert”, I would like to put forward what I have termed the “Five P” set of criteria that should be seen as one of several sets of criteria that need to be considered when developing and maintaining a state-of-the-art metallurgical plant. This set of criteria is made up of PROJECT, PEOPLE, PROCESS, PATENTS and POLITICIANS. Clearly there is a strong interplay between each of these components and each may individually or in combination lead to initiation of an outstanding technical and commercial success but conversely to an absolute failure.



From my perspective it is important to expect the unexpected and avoid any major deviation from the KISS principle. The title of this presentation indicates acquisition of information by osmosis and as such the following comments cover some of the more relevant observations made over the past 40 years. Project Geometallurgy is now properly recognised as a critical evaluation step in establishing a technically and commercially viable process flowsheet. Unless a detailed evaluation of the chemical and mineralogical complexity of the resource under consideration is executed then failure in some form is highly likely to follow. One of the real challenges is to ensure that testwork samples should be properly characterised and that the selected samples are representative of the expected resource over all aspects of the entire project life. This means that there has to be close cooperation and coordination between the mine planners (including the geologists), the client’s project team and the process flowsheet developers. While the distribution of the marketable metal(s)/product(s) has a major influence on mining and processing options, in reality it is how, where and in what form the non-value components are present and how to reject them efficiently that will be the key to success or failure. Geometallurgy is also a major factor that assists in the determination of realistic product outputs. Projects that can be described as conventional in that they yield standard products such as gold bullion via heap leaching of an oxide gold ore or cathode copper via smelting/electrorefining of high grade chalcopyrite concentrates can be relatively straightforward in that product specifications are well known and there is an active, open market. The situation with projects where each end-user sets their own specific product specifications is quite different. For example, virtually every end-user of dead burned magnesia will have quite specific requirements for crystallite size and shape in addition to other chemical and physical properties. There are no universal product specifications. While magnesite calcination may seem to be a relatively straightforward metallurgical process, it is the physical structure of the magnesium carbonate raw material as well as the calcination conditions that determine the properties of the magnesium oxide product. It follows that the owner/developer of a magnesia production facility must be fully aware of detailed market requirements and understand that it may not be practical to service a broad spectrum of customers. Thus there has to be an on-going assessment of potential markets as it is not a good option to tie production to a very limited number of customers. Practical metallurgical projects are generally based upon consumables that are widely available although some, such as sodium cyanide, will require a high level of process control and mitigation strategies to be put in place to overcome any plant failures however caused. People As practicing metallurgists, it is important to maintain a high level of realistic technical input into senior management considerations and, where appropriate, curb the misplaced



conviction by the technically illiterate that reliance on simplistic in-ground evaluations and avoidance of locked cycle continuous testwork at the pilot/demonstration level are just two of several ways of achieving a project that is not technically and commercially sustainable. It is not unusual for initial capital and operating costs to be considerably higher than originally anticipated so that cost-cutting procedures are instigated, particularly by the corporate finance management team. This may involve, for example, reduction in pipe and cable runs, reducing overhead access, reduction of automation and monitoring instrumentation, etc. While there may be a saving on structural steel, concrete, process control facilities etc., the end result will ultimately lead to increased maintenance costs, more complex occupational health and safety issues, less reliable adherence to operation within the designated band width and potentially extended shut-down periods. These technical issues must be properly flagged. The project’s technical champion has a very significant role to play. Bean counters need to be controlled so as to avoid equipment selection based on simple cost and availability terms, rather than on a genuine “fit for purpose” basis. There is no point in gold plating the processing plant, but it should not be deficient in practical operational and safety terms. Similarly there is no value in over-promoting the economic benefits that might accrue. Process It is possible to develop a process flowsheet that is chemically sustainable, at least at the theoretical and initial pilot/demonstration scale, but which never reaches commercial status because it is “impossible” to engineer and operate it and/or economic reality sets in. There are quite a few proposed processing flowsheets that are best described as the “tail wagging the dog”. This is particularly the case where the flowsheet incorporates recovery of every possible product in the purest possible form. For example, many of the proposed nickel laterite flowsheets incorporated recovery of metallic magnesium, given that the magnesium content of the pregnant leach liquor is many times greater than that of nickel. Only a little bit of evaluation from the other side of the fence clearly indicates that this approach is not sustainable. If the project does not make commercial sense based on the nickel cash flow with a nominal cobalt credit then the presumed financial benefits from additional by-products will just be an illusion. As a general but not totally universal comment, the “best” process flowsheets are developed and commercialised by project owners/developers. This is particularly the case with mineralogically complex resources since exploitation of the resource will not proceed without the successful development of the required flowsheet. This comment should not be taken as an unqualified criticism of R and D undertaken by universities, research organisations and independent companies. They certainly generate some great concepts. However, such groups rarely have the financial and technical expertise and abilities to undertake and achieve commercial status for their concepts and are reliant on selling their know-how. One must always question their ability to provide the necessary support when their know-how is implemented by others. In other words, the end user should always very seriously question the real value (technical and commercial) of the licence to use purchased technology.



As noted above, magnesium metal has been vigorously promoted as the metal of the future. Australia was at the forefront of developing what was being claimed as novel state-of-the-art processing regimes. Three major areas of sustained R and D centred on (a) purification and dehydration of concentrated magnesium chloride liquors, (b) development of alternative electrowinning electrolytes, and (c) improved cover gas technology. Despite spending in excess of $50M on their demonstration facility, amongst other things the AMC purification and dehydration technology failed to deliver the anticipated outcomes. The chemical/thermal decomposition of the organic dehydration agents was soon found to be quite deleterious. Electrowinning magnesium from fused salt electrolytes is highly energy intensive as well as challenging in engineering and operational terms. Molten magnesium metal is not the simplest material to handle. Energy consumption during electrowinning can be reduced by modification to the composition of the fused salt electrolyte. One such option is to add neodymium chloride. Unfortunately one side effect is that rather than pure magnesium being produced at the cathode, the end product is a magnesium-neodymium alloy. Such alloys have a number of useful physical properties, but it was soon worked out that the “loss” of neodymium to the alloy product was such that the volume of make-up neodymium chloride would soon exceed that currently available. In other words, the proposed magnesium metal production facility would be dependent upon the establishment of an on-going rare earth production facility with a significant neodymium output. The technical and commercial constraints so imposed basically canned the concept of the use of neodymium chloride additives to the magnesium electrowinning cell house. The cover gas technology developed by AMC and CSIRO was and remains technically astute – It is unfortunate that it has not been commercialised due to the total collapse of the AMC project. As noted previously, it is important to expect the unexpected, especially when considering some of the more complex hydrometallurgical flowsheets. For example, during carbon dioxide leaching of a number of caustic calcined magnesia feedstocks it was found that the resultant liquors had quite a high soluble ferric concentration (several g/L) even through the bulk pH was in excess of 9.5. Under the operating conditions it is possible to form soluble iron(III) carbonato complexes even though all the available thermodynamic data suggested that this was “impossible”. Development of a suitable technique for selective iron removal from the pregnant liquor proved to be quite challenging. During the 1970’s and 1980’s the vast majority of hydromet flowsheet development centred on the initial leach step, with less emphasis on the downstream purification and recovery steps. For example, chloride hydrometallurgy of base metal sulphide ores was seen as the panacea of all environmental ills associated with sulphur dioxide abatement with conventional pyrometallurgical operations, although it is appropriate to note that commercial reality is yet to be achieved. While detailed knowledge of the leach step was generated, many of the challenges with chloride hydrometallurgy relating to separation and recovery of elemental sulphur, recovery of precious metals from leach residues, deportment of nasties such as arsenic, regeneration and recycle of the leachant, sulphate control, etc., were subjected to far less rigorous evaluation. Fortunately this situation of unbalanced unit step development for most hydrometallurgical flowsheets is now being corrected.



Two of the driving forces for a balanced assessment of all hydromet unit steps are the need for and influence of a range of internal recycle stages and the maintenance of a workable process water balance. In some location and for some flowsheets there may be an excess of process water that needs to be discharged into the local environment in an acceptable, benign manner. In other situations there will be a potential deficiency of fresh suitable process water at reasonable cost, so recovery and recycle of process water is a requirement that adds to both capital and operating costs. Desalination plants have their own problems including power supply, plant duplication for continuation of supply and waste disposal. For hydrometallurgical flowsheets in particular, all definitive testwork from feed preparation, which may involve grinding and flotation, right through to final product recovery, must be carried out using process water that will be available at site. The physical and chemical properties of the process water will have a significant effect on parameters such density, viscosity, redox potential, oxygen solubility, ionic strength, leach and precipitation kinetics, etc, let alone materials of construction considerations and actual equipment design. It also follows that care must be taken in preparing process models since many of the thermodynamic inputs for most models are clearly deficient. Infrastructure requirements and power generation (particularly if “peak” oil status is fact not fiction) will impact on all remote sites as a cost burden as governments continue to bail out of these areas. Patents While it is not unreasonable to protect intellectual property by means of filing and executing patent applications, from a practical point-of-view the situation is tending to get out-of-hand. For example, since 2000 more than 90 patent applications on sulphuric acid leaching of nickel laterites have been filed via the World Intellectual Property Organisation (WIPO). The claimed flowsheets cover heap, vat, atmospheric and high pressure leach circuits, sometimes in combination with downstream unit steps such as iron precipitation, intermediate product recovery, solvent extraction and electrowinning. Sometimes it is quite difficult to discern any realistic technical differences between two or more claimed flowsheets. Part of the problem relates to the definition of patent laws and concepts such as novelty and inventiveness. It must also be remembered that a patent is basically an idea – it is not necessary to prove chemically and/or physically that the concept actually works. In fact in some instances a little technical nous indicates that the concept cannot work. Another feature of patent law that is sometimes difficult to fathom in technical terms is that known processes and concepts can be incorporated into primary and subordinate claims. For example, many of the nickel laterite sulphuric acid leach applications incorporate current commercial practice such as nickel/cobalt separation by solvent extraction using Cyanex 272. From a practical point-of-view, it seems difficult to understand how any claims relating to the use of this reagent in recent applications are sustainable. One apparent reason for the proliferation of patent applications relates to the fact that patent challenges/litigation is extremely costly in terms of time and dollars and user licence fees and conditions are often unacceptably onerous. Thus “new” flowsheets are devised to overcome claims of existing novelty/inventiveness even though such flowsheets may not be technically and/or commercially optimum.



It is now common practice to initiate a “freedom to operate” review of potential patent infringement as part of the overall risk analysis procedures during completion of feasibility studies. There are numerous examples of what are best described as “nonsense” patents. Probably the most technically “challenging” – and that description is certainly offered with a high level of derision – relates to the “neutralisation of an acidic stream using sized limestone”. It was originally allowed by the USPO on the basis of the word sized. Fortunately industry as a whole finally got the USPO to accept a dose of technical reality and the application was forced to lapse. Politicians The mining and metallurgical industries face a number of challenges as they are forced to cope with ever increasing volumes of more and more diverse bureaucracy. It is unfortunate that our politicians and their advisers as a whole are technically illiterate and are easily persuaded by pseudo scientific commentary. This can and does lead to constrained operating parameters because of the imposition of illogical rules and regulations as a means of placating vocal opponents of our industries. Perhaps the most bizarre was the push to ban the use of stainless steel because nickel has some rather ill-defined carcinogenic properties. Our industries must and do accept the imposition of logical environmental constraints but it is a pity that the overall record of compliance is not recognised by our politicians and the general public. Accidents do happen, but I suspect our record is somewhat more positive compared with those of many other primary and secondary industries. The challenge for us is to educate our politicians.

CONCLUSIONS To continue to learn by osmosis it is constructive to periodically assess why and how “wayward” metallurgical plants come into being. This presentation indicates that this occurs because of scientific-engineering incompetence, project owner/developer avarice, insufficient attention to future market/product requirements, imposition of impossible operating constraints, right through to obvious bad luck, but it is more likely to be a combination of all such factors.



Predicting Variations in Mill Feed

P McCarthy1 1. FAusIMM(CP), Chairman, AMC Consultants, 19/114 William Street, Melbourne Vic 3000.

Email: [email protected]

ABSTRACT The relatively new field of geo-metallurgy promises to improve the predictability of mill feed quality and enhance processing outcomes. However, it is subject to the same limitations as the prediction of grade using geostatistics and has additional limitations due to the spatial distribution of the variations being measured and the sample sizes required. Schedules based on geo-metallurgical models will, in turn, be subject to variations arising from geological control and practical mining constraints.

This paper discusses the key drivers of variations in mill feed quality and quantity from a mining perspective. It concludes that geo-metallurgical modelling addresses only one aspect of the problem. A case study is provided to illustrate poor predictability of metallurgical performance even with very detailed geological data.

INTRODUCTION The quality and quantity of ore delivered from the mine to the plant stockpile will vary at all time scales from hourly to monthly and beyond. The scale of these variations must be predicted early in the feasibility study so that an appropriate stockpile management strategy can be developed and so that metallurgical sampling and pilot plant testing can be arranged to cover the full range of likely conditions.

The ability to predict variations in the quality and quantity of ore delivered from the mine depends on good geo-metallurgical data collection and modelling, and is subject to the predictability of the mining plan. The reliability of the geo-metallurgical model will depend on the inherent variability of mineralisation, on the density of sampling and on the range of data collected. The predictability of mining outcomes depends on the mining method and its flexibility, the rate of mining relative to the scale of the orebody, the thickness and physical variability of the mineralisation, the effectiveness of grade control, the reliability of equipment and the capabilities of mine management. All of these factors must be understood, to some measure, when designing the processing plant to cope with short-term variations and longer term changes in the material coming from the mine.

GEO-METALLURGICAL DATA COLLECTION AND MODELLING – A MINER’S PERSPECTIVE The process should begin with a good three-dimensional geological model developed by an experienced geologist familiar with the deposit type. Geological domains must be identified, such that a common set of rules can be applied to determine local variations in metallurgical responses within each domain. The domain boundaries may be structural, mineralogical, alteration or lithological. Poor geological modelling and domaining are the leading causes of failure in geostatistical modelling for grade estimation and will be the same for modelling metallurgical parameters.

Domains should be defined beyond the “orebody” to include all material that could find its way into the ore stream. Metallurgically, adjacent domains may have little or nothing in common. For example, the waste rock adjacent to the orebody across a sharp contact may be much harder and more abrasive than the ore, and if the mining method will cause 25% mining dilution of ore with waste rock then the crushing, grinding and wet plant performance of the waste rock must be thoroughly understood. Similarly nearby carbonaceous shales (such as a hanging-wall zone) may be preg-robbing and a talc-rich fault zone may impact on filtration, although neither is considered part of the orebody proper.



Once the domains (both within and near the orebody) have been described (as “wire frames” or solid objects in a three dimensional computer model), representative samples from each domain can be subjected to laboratory-scale test work to determine the rock’s response to each mineral processing operation. Conventional geostatistics can then be used to model the distribution of metallurgical responses throughout each geological domain. This results in a model comprising a large number of blocks in three–dimensional space, each block being assigned all the geological, geotechnical, geo-metallurgical and other characteristics needed for mine planning and scheduling. The mine scheduling process can then produce from the geological block model not only a schedule of tonnes and grade but also a schedule of metallurgical performance and other characteristics such as ground support requirements or water inflows to the mine.

The geostatistical approach used to model metallurgical performance need not be complex. Even the simplest approaches using the “rule of nearest points” (which says each block in the model is likely to perform in a similar way to the nearest sample) or the “rule of gradual change” (which calculates a distance-weighted average of characteristics based on nearby samples) is likely to provide a significant improvement in predictability of plant performance when compared with having no geo-metallurgical model. However, the more advanced geostatistical methods are not difficult to apply and will further refine the result. Selection of the best techniques is the subject of ongoing research.

Samples for metallurgical testing are usually composited from diamond drill core. Hardness testing (for crushing and grinding) typically requires 10-20kg of sample, with some tests requiring 100kg, while flotation testing usually requires at least several kg of sample (Barrat and Doll, 2008). Large metallurgical samples excavated from near the surface of a deposit are unlikely to be representative of the orebody at depth. Shafts sunk for the purpose of obtaining large metallurgical samples may also yield unrepresentative samples, or samples that represent performance in only one geological domain. Such exercises may be compared to searching for a lost wallet under a streetlight, because it is too dark to search elsewhere.

In order to be useful in developing a geo-metallurgical model, test results must satisfy the following (ibid):

Results must reflect the properties of a “small”, identifiable interval of drill core. The location of the interval must be identifiable in three-dimensional space (to connect it to

the block model). The values being distributed through the orebody must be reasonably additive, allowing

unknown blocks in the model to be estimated by interpolating two or more known samples.

Sufficient sample material to achieve these aims may be available from drill core for a large porphyry copper open pit which has a large Selective Mining Unit (SMU or minimum mining block size). For smaller, more complex deposits and many underground mines where assays are obtained for each 1m sample interval, the production of composite samples of sufficient size for metallurgical testing may defeat these aims and blur the modelling results. For example Barratt and Doll propose sample intervals of 15m of HQ (63mm) core for a JK SMC (drop weight) test and 45m to 105m of HQ core for a Bond test. It may be necessary to develop local correlations between the large scale tests and other properties such as point load strength, RQD, fracture frequency and mineralogy in order to obtain sufficient data to create a meaningful three-dimensional model.

PREDICTABILITY OF THE MINING PLAN The specifications for the processing plant should reflect real hourly or daily mining outcomes, not a smoothed and idealised schedule. Table 1 summarises the unpublished results of benchmark studies of 44 underground mines and 21 open pit mines conducted by AMC Consultants. The monthly variability shown is the average absolute difference (as % of budget) between the mine budget and production over a 12 month period. High variability indicates that a mine is not operating as intended, and that daily variability may be much greater than the steady hourly throughput for which the plant was designed. It can be concluded that:



Surface mines suffer significantly higher variability in ore production from budget than underground mines. This is surprising, and may be in part because it is easier to designate each truck load as ore or waste, causing hourly or daily variability. Underground mines just throw unforseen waste into the ore stream.

Despite the mining variability, processing rates in surface mines run closer to budget, perhaps due to having larger stockpiles.

Variability in mined and processed grades is similar, suggesting that little use is made of stockpiles for blending and that grade control outcomes are similar in the surface and underground mines studied.

Table 1 Monthly Variability from Budget

Underground Mines

Surface Mines

Min Max Ave Min Max Ave

Ore Mined (t) 5% 43% 14% 7% 68% 29%

Head Grade 7% 33% 13% 3% 38% 11%

Ore Processed (t)

5% 54% 12% 4% 39% 9%

Processed Grade 3% 33% 13% 2% 31% 10%

Another useful measure is volatility, the average % change in a measure from one time period to the next. The level of planned and actual volatility drives the stock requirements and levels. When volatility is low, mining and processing are efficient and capacity is being used effectively with costs minimised.

Table 2 shows the month to month volatility from the benchmark study mentioned above. It can be concluded that:

Ore tonnage mined is nearly twice as volatile in surface mines compared with underground mines

On other measures, underground mines are more volatile than surface mines

Table 2 Month to Month Volatility

Underground Mines

Surface Mines

Min Max Ave Min Max Ave

Ore Mined (t) 6% 34% 14% 11% 62% 25%

Head Grade 5% 28% 11% 3% 22% 9%

Ore Processed (t)

3% 32% 13% 4% 22% 10%

Processed Grade 5% 23% 12% 3% 16% 8%



Carter (2010) observes “Volatility and variation in any business process creates uncertainty, whether in determining mining volumes and plant feed or in optimising maintenance schedules and supply-chain management. The same holds for the management of working relationships. It is this variation and volatility which skews outcomes, and which can be reduced by implementing rigorous planning, scheduling, resourcing and execution processes, and most importantly, by clarifying roles and accountabilities at each level.”

When volatility and variability are assessed on an hourly or daily basis they are seen to be much greater than the monthly measures in Tables 1 and 2. For example, Figure 1 shows daily plant tonnes at Anglogold Ashanti’s Mponeng mine. The figure also shows the results of a business improvement initiative which contributed to a 15% increase in throughput over the historical average. This is an especially significant result at the Mponeng plant, long regarded as the flagship operation within the group. Before the initiative, ore from the Mponeng mine would regularly be trucked to neighbouring plants for processing, as the mill struggled to cope. Now the plant has improved productivity to the point that it now has spare capacity. Emphasis on stabilised processes has also resulted in a 20% reduction in sodium cyanide consumption.

Figure 1 Mponeng Mine Daily Plant Tonnes (over 12 Months) (after Carter, 2010.)

MINING RATE RELATIVE TO THE SCALE OF THE OREBODY The optimum plant capacity for a new mine is usually based on empirical studies or “rules of thumb”, subject to confirmation by detailed scheduling of the proposed mining operation. The assumption that “economies of scale” will result from increasing throughput rates needs to be balanced by an awareness of the adverse effects on grade of increasing the rate beyond a level that is supportable by the resource. This effect is known to people at operations but is not generally recognised in current ore reserve estimation methodology. Most economic studies vary the throughput without adjusting the grade.

This effect was the subject of a study by the author (McCarthy, 2010). If historical production data is analysed whereby grade is expressed as a percentage of the weighted mean historical grade, and annual production is expressed as a percentage of the average annual historical production, then a dimensionless constant k can be defined such that

k = - ΔG(%) / Δt(%)……….(1)

This is simply a statement of the rate at which grade declines as tonnage throughput increases.

Historical values of k for Australian underground and open pit mines cluster around a mode of 0.3 and 70% of them lie between 0.2 and 0.5. These k values should be used with caution because they incorporate the influences of changes in cutoff grade, planned dilution, unplanned dilution and in some cases mining technology, as well as the effects of changes in orebody characteristics over



time. Nevertheless, they reflect valid choices of operating points for these orebodies and it is better to use these values than to assume that the mining rate has no effect on head grade.

Example

A mine is operating at a rate of 1.0 Mtpa at a head grade of 2% Cu. An expansion to 1.5 Mtpa is proposed and a value of k = 0.3 is inferred from similar deposits. The new head grade can be calculated from

G2 = G1 (1 – k (t2-t1)/t1)………………(2)

G2 = 2 x (1 - 0.3 x (1.5-1.0)/1)

G = 1.7% Cu.

The mining rate also affects production volatility, which can increase to the point of causing mine failure in extreme cases. In general, equivalent vertical advance rates greater than 60 metres per year are likely to be unsustainable (ibid).

PRACTICAL MINE OPERATION The ability of the mining team to deliver predictable quantity and quality will depend on the thickness and physical variability of the mineralisation. The more challenging the mining situation, the greater the stock levels need to be including developed (exposed) stocks, drilled stocks, broken stocks, and ROM pad stocks. If these stock levels are adequate then variability and volatility can be reduced to a minimum. The mine should be designed so that all stockpiles, including orepasses in an underground mine, have adequate capacity to smooth the short-term surges to a level acceptable for the system as a whole, including ore processing. This is a commonly overlooked requirement.

Some orebodies are amenable to visual grade control, others require assays on a short turnaround to allow ore selection decisions to be made. The latter can suffer from dilution and high grade volatility if the grade control programme is not well designed and given priority at the laboratory.

The reliability of mining equipment has an effect on ore quality. Delays in mine development (accessing ore in an underground mine or pre-stripping in a pit) can lead to periods when low-grade or high-impurity ore is all that is available. Breakdowns in ore-production equipment can lead to increased dilution because it is human nature to be less concerned about dilution when there is insufficient ore available to feed the mill. Hence the old saying “waste plus ore equals more ore!”

Proper mine design, planning, scheduling and maintenance require good management. Ultimately, the capabilities of the mine management team will determine the quality and regularity of mill feed.

CASE STUDY – PREDICTING THE UNPREDICTABLE IN GILL REEF The Bendigo orebodies have been shown to suffer from an extreme nugget effect, making prediction of grades from drill data difficult. For this reason Gill reef was mined, processed and reconciled in small batches and provides a more detailed picture of ore variability and plant performance than is usually available. The author examined the data to see whether metallurgical recovery could have been predicted from the predicted head grade of each batch.

Gill reef is a distinct quartz reef at the Kangaroo Flat mine that was mined between 2009 and 2011. A total of 55 discrete ore blocks totalling 243,497 tonnes from Gill Reef were mined, stockpiled and processed separately. The pre-mining block grades were estimated from diamond drill assays by geostatistical methods using a 150 g/t Au top cut. Overall, Gill reef was estimated from drilling to have a grade of 7.5 g/t Au which reconciled (for comparable blocks) with a mine head grade of 7.6 g/t Au. Individual block reconciliations, with a linear best-fit line, are shown in Figure 2



Figure 2 Predicted and actual block grades, Gill reef.

It was possible to reconcile mill recovered grade against mine head grade for 45 blocks ranging from 2064 t to 9799 t, with a mean size of 4426 t. Head grades ranged from 2.2 g/t to 12.6 g/t with a weighted mean of 7.6 g/t. Reconciled metallurgical recoveries ranged from 71.5% to 96.7% with a weighted mean of 89.1%. Individual block reconciliations, with a logarithmic best-fit line, are shown in Figure 3. A constant tail grade model was not a good fit to the data.

Figure 3 Head grade and metallurgical recovery, Gill reef

With sufficient experience in similar orebodies it should have been possible to predict from the drill results that mining about 240,000 tonnes of ore would yield a head grade of about 7.5 g/t Au and a metallurgical recovery of about 90%. However, these long-term averages would be of little use in predicting or optimising process plant performance on a daily or weekly basis at the actual mining rate from Gill reef of 2,000 to 3,000 tonnes per week. Figure 4 shows the predicted metallurgical recovery for each mining block (based on the predicted head grade and the grade-recovery relationship) and the actual metallurgical recovery, with a weak linear best-fit relationship. It is clear that metallurgical performance was not predictable at a useful level.



Figure 4 Predicted and actual metallurgical recovery

While attempts could be made to improve the metallurgical model by modelling other geological features such as carbonaceous shales or associated sulphides, the innate geological variability makes success unlikely at the actual scale and rate of mining. At much higher mining rates, or when considered over longer periods, the volume-variance relation would allow greatly improved predictability.

CONCLUSION The new techniques of geo-metallurgical modelling are useful in improving process plant performance and predictability in large orebodies such as porphyries, mined with big equipment. In these situations zonations may allow prediction of changes over a period of years. However, many medium to small-scale mines may not be amenable to the techniques being developed. The problem is that variability of the factors affecting metallurgical performance occurs at a scale smaller than can be sampled for metallurgical testing, and at that scale the measured properties are not well correlated with performance. The same problem exists for geotechnical modelling, where very limited success has been achieved in predicting ground conditions and stability using geostatistical methods.

ACKNOWLEDGEMENTS The author acknowledges AMC Consultants Pty Ltd for permission to publish benchmarking results and Unity Mining Limited for permission to publish the Gill Reef data.

REFERENCES Barratt, D.J., and Doll, A.G., Testwork Programs that Deliver Multiple Data Sets of Comminution Parameters for Use in Mine Planning and Project Engineering, Procemin 2008, Santiago, Chile, 2008

Carter C.: Project One Holistic Transformation Plan to Achieve Strategic Goals, Anglogold Ashanti 2010. http://www.anglogoldashanti.com/subwebs/InformationForInvestors/Reports10/financials/project-one.htm

McCarthy, 2010. Setting Plant Capacity. Mineral Processing and Extractive Metallurgy (Trans. Inst. Min. Metall. C) Vol 19 No 4. pp184-190



Setting Processing Plant Projects up for Success in Australia

R Young1 and A Hollonds2

1. Regional Director, Asia Pacific, Independent Project Analysis, Level 1, 56 Burgundy Street, Heidelberg Vic 3084. Email: [email protected]

2. MAusIMM, Senior Project Analyst, Independent Project Analysis, Level 1, 56 Burgundy Street, Heidelberg Vic 3084. Email: [email protected]

ABSTRACT

Independent Project Analysis (IPA) is in the business of benchmarking the performance of capital projects on behalf of owner companies in the processing and extractive industries. IPA assesses between 500 and 800 projects each year. The projects are conducted globally and range in size from about $1 million to over $30 billion. Using the data, we are able to benchmark the performance of Australian projects against the rest of the world. This presentation focuses on the performance of large (greater than $100 million) projects that involve chemical or thermal processing.

We regularly hear calls for more processing to be conducted in Australia but on the other hand industry leaders are concerned about the poor success rate of solids processing projects in Australia. Data gathered by IPA indicates that there is good reason for the concern. Out of a sample of 31 large processing projects conducted in Australia, 23, or 74 percent, were classed as failures. This performance is significantly worse than for similar projects in other parts of the world. The presentation analyses why the performance is worse and demonstrates that many of the reasons behind the failures were manageable by the project team. Furthermore, members of the project teams on at least half the failed projects knew that the projects were on a path to failure before the project was authorised but could not be heard.

The presentation discusses how large projects that involve process design and, in many cases, new technology, need to be set up for success.



Reducing Mining and Mineral Processing Plant Fatality Rates

N Cann1, S Casey2, R Mills3 and J Ross4

1. Principal Risk Consultant, GHD, Melbourne Vic 3000. Email: [email protected]



4. MAusIMM, Principal Risk Consultant, GHD, Melbourne Vic 3000. Email: [email protected]

ABSTRACT

In the 10 years from 1998 to 2008 the annual number of fatalities in the Australian mining industry has plateaued (MCA, 2007-08 and abc.net.au website). Despite significant efforts to eliminate fatalities, this is not a comfortable statistic in this very public performance indicator.

GHD has been working with the mining and minerals processing industry to identify the key risks that lead to fatalities utilising a semi quantitative risk assessment technique (SQRA®), initially developed to assist Major Hazards Facilities prepare Safety Cases (OH&S, 2007-08) in response to the Longford Incident in 1998 (Dawson and Brooks, 1999).

Workshop cross functional teams are assembled from the sites and led by the GHD facilitators to identify the incident types that can have fatal consequences, the causes of those incidents and the controls that are utilised (or could be utilised) to reduce the risk.

The risks for each incident type are quantified to produce a risk profile for the facility and the critical controls – the controls that have the greatest impact at reducing the risk of fatality – are selected. These controls are assessed for their adequacy and improvement actions are identified. Prioritisation of the actions is achieved by re-estimating the incident risk profiles.

A review has been undertaken using a small sample from the more than 100 individual studies that GHD has completed at mining and minerals processing sites. The five most significant risks reported by sites are:

1. Fall from Heights,

2. Vehicle/ Vehicle and Vehicle/Pedestrian interactions,

3. Entanglement,

4. Molten Metals and Materials, and

5. Electricity.

The most significant reductions in risk identified were found to be by implementing engineering controls. Examples include:

1. Chutes to contain splashes and spills of molten materials,

2. Redesigning instrument and equipment layouts to take personnel away form high risk areas,

3. Designing remotely operated equipment to remove personnel away from high risk areas,

4. Proper fire fighting equipment provisions such as deluges around flammable materials, and

5. Adequate risk assessment to ensure equipment is designed to handle the real situations and not a design case to meet an arbitrary specification.



All of these cases and many of the actions identified in the studies could have been handled through better design and construction early in the life of a project rather than retrospectively fitted after operations have begun.

INTRODUCTION

In the 10 years from 1998 to 2008 the annual number of fatalities in the Australian mining industry has plateaued as illustrated in Fig. 1 (MCA, 2007-08). Despite significant efforts to eliminate fatalities there has not been a reduction in this very public performance indicator over that 10 year period. However, there has been significant growth in the mining industry over this period so in relative terms the fatality rates will have declined.

Fig. 1 Fatalities 1994-95 to 30 June 2008 (MCA, 2007-08)

Over the last six years GHD has been working with the mining and minerals processing industry to identify the key risks that lead to fatalities. A semi quantitative risk assessment technique (SQRA®), initially developed to assist Major Hazards Facilities preparing Safety Cases (OH&S, 2007) in response to the Longford Incident in 1998 (Dawson and Brooks, 1999), has been utilised.

The SQRA process is a phased risk assessment method that utilises a workshop approach with participants drawn from a wide range of engineering and operation functions within the site being studied and led by experienced GHD risk facilitators. This ensures the process benefits from site experience, site knowledge, operating history and can be completed in a timely manner.

The SQRA process consists of seven steps, summarised in Fig. 2 below:



Fig. 2 SQRA process

The SQRA process has been developed to allow sites to understand their most significant fatality risks, often referred to as major accident hazards (MAHs). It identifies the MAHs, the pathways that lead to those accidents and the controls that are in place (or could be introduced – called potential controls) to prevent or reduce the likelihood of the event happening or to mitigate the consequences.

To put a line in the sand, the risk of the MAH leading to a fatality is semi-quantified by working through scenarios from the initiating event to the fatality, with the workshop team assigning probabilities along an event pathway. This produces a numerical estimate of the risk namely the current Potential Loss of Life (PLL) for the MAH.

As not all controls are equal in their ability to reduce risk, the most significant ones are chosen by the workshop team and these are deemed to be Critical Controls. These controls are examined in further detail to establish their adequacy by using parameters such as dependability, practicality, survivability, workplace involvement, performance monitoring and training. This analysis usually identifies a number of actions for improvements that can be undertaken to increase the adequacy of the critical controls.

Through completing the actions and introducing the potential controls identified earlier, it is expected that at some future point the MAH risk will be reduced. This predicted risk reduction is calculated by revisiting the PLL figures for the scenarios examined in the risk assessment step.

This paper looks at the collective results from a small sampling (9 to date) of that work to get an understanding of the common MAHs, the range of PLL figures for each, the key controls utilised by sites for the higher risks and the common actions for improvement being identified to reduce the risk of fatalities on sites.

1. Hazard Identification

QRA

Safety Management

System

2. Hazard Dynamics (Bowtie Diagrams)

3. Risk Assessment

4. Identification of Critical Controls

5. Critical Control Adequacy Assessment

6. Selection of Risk Reduction Actions & Predicted Risk

7. Reporting

Preparation, planning, facility description

ALARP Workshop



SCOPE OF THIS STUDY

The study is based on an initial assessment of databases to hand for the examining team and thus only represents a small proportion of the work that will be examined in the future.

The sites examined cover the following types of operations:

zinc mining and concentration, gold mining, bauxite mining, alumina refining, and aluminium smelting.

Three separate corporate entities operate these sites; although with changes in ownership occurring within the industry five distinct work place cultures are reflected. The data has been examined to identify the most significant and common MAH risks identified for the operations with the range of PLL figures and median PLL identified for the current risk. A similar process has been completed for the predicted risk profiles so that the most common improvements could be identified.

The common controls that are in place and utilised by the industry have been identified and the most common high priority risk reduction actions have been reviewed for the MAHs where at least a half an order of magnitude risk reduction in the average MAH PLL has been recorded.

CURRENT MEDIAN RISK

Across the nine operations included in the study 16 broadly grouped MAHs were encountered. These are shown in Fig. 3 where the range of PLL risks are indicated by the vertical lines and the median result in each MAH category is indicated by the diamond. Categories with a limited range across the sites are indicated by an asterisk.

With the exception of the hazards associated with molten materials, the top risks are all general risks associated with heavy industry and not the inherent hazards associated with mining and minerals processing. In existing operations these risks are the residual risks left by the designers and the construction crews. By understanding these outcomes in operating plant, designers can work towards eliminating risks that pose a threat to the operating crews that have to work with the results. Consideration needs to be given to how crews will operate the plants so that risks from falling from heights, entanglement in conveyors and rotating equipment and electricity are eliminated.



Current Mining and Minerals Processing Risks

1.E-10

1.E-09

1.E-08

1.E-07

1.E-06

1.E-05

1.E-04

1.E-03

1.E-02

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Min Risk Value:

Max Risk Value:

Median Current Risk Value:

Fig. 3 Current Mining and Minerals Processing Major Accident Hazard Risk Profile

CRITICAL CONTROLS

In order to manage these MAH risks, Operations have introduced a large number of controls. For the top eight risks (that have sufficient data to be relevant) the Critical Controls are listed in Table 1. Overall the dominant controls are procedural with specifically designed engineering solutions a close second. This is followed by competence and training and then inspection and testing.

In many well-known incidents such as Piper Alpha, Bhopal, Texas City and BP/Transocean Deepwater Horizon (Keltz, 1994, Keltz, 1993 Hopkins, 2000 and 2002 and www.cbs.gov), these have all had either compromised controls, critical procedures not followed and equipment known to be critical not maintained and in poor or non – operational condition. Therefore part of any SQRA process is to examine the Critical Controls i.e. the ones that offer significant risk reduction or mitigate the consequences of the MAH.

This adequacy assessment looks in detail in a number of broad areas to understand how a control is managed and maintained. The details examined include (but are not limited to) parameters such as: dependability, practicality, survivability, workplace involvement, performance monitoring and training. If the process is being undertaken for regulatory purposes then documentation, records and evidence will be sought to verify the assertions made in the workshops.

As a result of this analysis all operations have identified actions that need to be undertaken to lower the risk of MAHs. For more detail on the types of actions undertaken see Table 2 later in this report.



Table 1 Most Common Critical Controls for Major Accident Hazards

MAH Risk Most Common Critical

Control Second most common

critical control Third Fourth Fifth

Fall from Heights Procedural and rule based controls

Inspection and Testing Engineered solutions Competence (training and induction)

Access Control (Locks, barriers, signs)

Vehicle Interactions

Competence (training and induction)

Engineered solutions Access Control (Locks, barriers, signs)

Procedural and rule based controls

Preventative Maintenance

Entanglement Procedural and rule based controls

Engineered solutions Competence (training and induction)

Permit to Work Access Control (Locks, barriers, signs)

Molten Metals and Materials

Engineered solutions Access Control (Locks, barriers, signs)


Event Alarms Inspection and Testing

Electricity Engineered solutions Procedural and rule based controls

Industry Standards Competence (training and induction)

Condition Based Maintenance

Confined Spaces

Permit to Work Procedural and rule based controls

Management Plans Competence (training and induction)

Emergency Response Plans

Dropped & Falling Objects

Engineered solutions Inspection and Testing Procedural and rule based controls

Competence (training and induction)

Preventative Maintenance

Engulfment/Fall into Liquid


Engineered solutions Permit to Work Competence (training and induction)

Access Control (Locks, barriers, signs)



PREDICTED RISK REDUCTION

With the actions for improvement identified in the critical control adequacy assessment and the potential controls identified earlier in the hazard identification phase, the SQRA® process then revisits the risk calculations to show how implementing the improvement actions will lower the operations MAH risk.

These changes have been summarised in Fig. 4 and these results can be contrasted with the earlier PLL results presented in Fig. 3. Note that the order of the risks has not been changed.

The changes in the median PLL for each risk can also be seen in Fig. 5 so that a direct comparison of the improvement can be made.

Fig. 4 Predicted Mining and Minerals Processing Major Accident Hazard Risk Profile

Predicted Mining and Minerals Processing Risks

1.E-10

1.E-09

1.E-08

1.E-07

1.E-06

1.E-05

1.E-04

1.E-03

1.E-02

1.E-01

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Min Risk Value:

Max Risk Value:

Median Predicted Risk Value:



Fig. 5 Median Risk Reduction Opportunity Identified

Median Risk Reduction at Mining and Minerals Processing Plants

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Median Current Risk Value:

Median Predicted Risk Value:

COMMON ACTIONS THAT LEAD TO SIGNIFICANT RISK REDUCTION

In order to better understand the opportunities for designers and constructors to minimise risk in the final operation, a review of the improvements that operations have identified that provide the greatest risk reduction (as illustrated in Fig. 5) has been undertaken. The results are presented in order of greatest improvement in the median PLL figures and appear in Table 2.

Table 2 Significant Risk Reduction Actions

MAH Risk Typical Improvement Actions

Fall from Heights

Improvement actions include:

1. Remove the need to work at height, provision of access/permanent access platforms or relocation of equipment currently outside of handrails.

2. Inspection of walkways, handrails, ladder ways, pit covers, floor gratings to ensure integrity is maintained.

3. Training or refresher training for the use of fall restraint/arrest equipment, training in permit to work system.

4. Barricading open holes, installing fall containment to hatches with potential for access and ensuring required equipment is available for maintenance of scaffolding.




Molten Metals and Materials

Actions focussed on removing and limiting personnel access to the area and improved equipment design.

Equipment design changes include:

1. Chutes designed to prevent splashing of molten metal.

2. Crucible covers to prevent splashes and spills.

3. Bail arms designed to prevent point loading.

4. Cameras to aid observation and avoid overfilling of crucibles.

Removal of personnel from area achieved by defining exclusion zones, segregating personnel and providing designated pedestrian access paths.

Entanglement

The most common action across all sites was to conduct a guarding audit or in the case where it had already been undertaken, implement the recommendations.

Engineering controls to include:

1. Anti-rollback conveyors.

2. Installing dead man switches.

3. Interlocks on gates.

4. Developing methods for tracking conveyors without removing guards.

Signage indicating pinch points and fencing around conveyors to prevent unauthorised access were also listed.

Vehicle / Pedestrian Interactions

The key action for this risk across the different operations was the segregation of vehicles (light vehicles, forklifts etc) and pedestrians. This was achieved by restricting vehicle access to certain areas, providing dedicated pedestrian access routes and walkways, and the enforcement of an exclusion zone around heavy vehicles.

Collision avoidance systems (including reversing/blind spot cameras) and proximity alarms were also listed, as well as speed limiters on vehicles and speed limit reductions. Improved equipment maintenance programs, ensuring procurement of correct equipment and replacing older vehicles with new vehicles (which have improved safety features).

Improvements to radio communications and revision of HV and LV tickets, training and assessment packages and ensuring currency of these is maintained.

Traffic and road management plans and standards were established to have a responsible person assigned to ensure better road conditions are maintained as well as the re-designing of problem intersections, entrances and pedestrian crossings.




Confined Spaces

Actions focussed on developing or updating a confined space register to include hazards/controls/rescue plans specific to each confined space.

Updating the confined space entry permit to include the need for site specific risk assessment/controls/rescue plans and training workforce in hazards associated with each confined space.

Recruitment and retainment of an Emergency Response Team (ERT) and implementing a system to prevent confined space entry being scheduled while ERT is occupied.

Natural Forces

Risk reducing actions include the predetermined evacuation trigger levels particularly in relation to tsunamis and exclusion zones established around unsecured structures during cyclones.

Ultimate risk reduction is obtained by replacing the unsecured structures.

Hazardous Chemicals

A range of actions specific to all chemicals were identified including:

1. Removal of personnel from area permanently (by redesigning equipment to move location of pressure indicators) or during certain tasks (commissioning/re-pressurisation of equipment)

2. Introducing and managing alarms and trips to enable evacuation from the area

3. Isolating/guarding of equipment (e.g. equipment covers, enclosed cages)

4. Improvements to PM programs to reduce the rate of equipment failure

5. Redesign of relief system to cater for volume of liquid released.




Fire and Explosion

Actions focussed on engineering controls to prevent fire/explosion and mitigation controls to reduce severity of event.

Engineering controls include; quick fill connection system; electronic detonators; tracking roller frames to reduce friction; alternative method of breaking oversized rock to reduce the need to use explosives, hardware control to prevent uncontrolled raising of anodes.

Mitigation controls include; deluge system on conveyors and review of PM frequency, installation of fire detection system/alarms and designated evacuation route, blast/fire walls and an audit of fire protection system to ensure water is readily available.

Removal of personnel from area using auto casting and relocation of steam lines.

Procedural and associated training material changes to refuelling and scrap management.

In looking at the range of actions to be undertaken, it can be concluded that the largest risk reductions are obtained by instituting engineering controls to move people away from the hazards. Designers should pay attention where people have to interact with hazards and separate them. Remedies identified early in the design phase will be less costly (in cash and fatalities) than those implemented once a site becomes fully operational.

CONCLUSIONS

By reviewing the MAH risks as recorded from nine mining and minerals processing SQRA studies the top five risks are:

1. Fall from Heights;

2. Vehicle/ Vehicle and Vehicle/Pedestrian interactions;

3. Entanglement;

4. Molten Metals and Materials; and

5. Electricity.

Significant reductions in risk were identified by implementing engineering controls. Examples include:

1. Chutes to contain splashes and spills of molten materials;

2. Redesigning instrument and equipment layouts to take personnel away form high risk areas;

3. Designing remotely operated equipment to remove personnel away from high risk areas;



4. Proper fire fighting equipment provisions such as deluges around flammable materials; and

5. Adequate risk assessment to ensure equipment is designed to handle the real situations and not a design case to meet an arbitrary specification.

All of these cases and many of the actions identified in the studies and presented in this paper could have been handled through better design and construction early in the life of a project rather than retrospectively fitting after operations have begun.

ACKNOWLEDGEMENT

The authors would like to thank Gleb Shinkarsky and Karen Barns for the detailed data collection and processing necessary for this report and to Mike Erskine and Eric Duprey for their review.

References

1. Minerals Council of Australia - QUARTER 4 • 2007-08 SAFETY SURVEY REPORT

2. http://www.abc.net.au/news/stories/2007/08/24/2013822.htm

3. Part 5.2 – Major Hazard Facilities, Occupational Health and Safety Regulations

2007, S.R. No. 54/2007, Victoria

4. Dawson, D and Brooks, BJ – Report of the Longford Royal Commission, June 1999

5. Kletz, T – Lessons from Accidents, 1994, 2nd edition, Oxford: Butterworth-Heinemann Ltd.

6. Kletz, T – Lessons from Disaster, IChemE,

7. Hopkins, A – Lessons from Longford, CCH, 2000

8. Hopkins, A – Lessons from Longford: the trial, CCH, 2002

9. http://www.csb.gov/investigations/detail.aspx?SID=96


Professional Development for Metallurgists – Improving Technical Skills

D Drinkwater1, D Bradshaw2, P Tilyard3 and P Munro4

1. MAusIMM, Technical Program Manager, SMI Knowledge Transfer, JKTech Pty Ltd, 40 Isles Road, Indooroopilly Qld 4068. Email: [email protected]

2. Professor, Julius Kruttschnitt Mineral Research Centre, University of Queensland. 3. FAusIMM(CP), Group Metallurgist, Minerals and Metals Group (MMG). 4. MAusIMM, Senior Principal Consulting Engineer, Mineralurgy Pty Ltd.

ABSTRACT

In 2009, Peter Munro and Peter Tilyard examined the “human capital” in the mineral processing

sector in a paper entitled Back to the Future (Munro and Tilyard, 2009). One of their major

concerns was a general decline in technical skills and process knowledge, exacerbated by lack of

appreciation of the impact this can have on process plant performance and, ultimately, profitability

or sustainability of the industry.

There are many reasons for this decline, not least of which are the work schedules of today’s

metallurgists which have a very different balance between technical and non‐technical responsibility

than those in the past. However, a shift towards a more technical focus has potential to deliver a

range of benefits for industry at large and for the individuals in it

This paper discusses some options for development of core metallurgical and process engineering

skills that can be used by graduates, employers and training providers.

Key skills are identified and mapped against available opportunities for development, such as

existing graduate programs, on‐the job training and external courses. The authors also review a

number of past and present high‐quality graduate development programs, including the graduate

development program in place at Mount Isa Mines in the 1980s and the Anglo Platinum Graduate

Development Program (AGDP) currently operating in South Africa.

INTRODUCTION

In 2009, Peter Munro and Peter Tilyard examined the “human capital” in the mineral processing

sector in a paper entitled Back to the Future (Munro and Tilyard, 2009). One of their major concerns

was a general decline in technical skills and process knowledge, exacerbated by lack of appreciation

of the impact this can have on process plant performance and, ultimately, profitability or

sustainability of the industry.

As a consequence mining companies, focussed as they are on day‐to‐day profitability, tend to

neglect development of high calibre technical skills at the expense of non‐technical training in

financial management, safety and personnel management, all of which can be clearly demonstrated

to have an impact on the so‐called bottom line.


Their conclusions were that a return to a more technical focus will deliver financial benefits to

industry at large and job satisfaction to the individuals within it.

This paper presents some guidelines as a follow‐up, for development of core metallurgical and

process engineering skills that can be used by graduates, employers and training providers, and

some suggestions about how to deliver them in the metallurgical work environment of 2011. Key

skills are identified and mapped against available development opportunities, such as existing

graduate programs, on‐the‐job training and external courses. The authors also reviewed a number

of high‐quality past and present graduate development programs, including the graduate

development program in place at Mount Isa Mines in the 1980s and the Anglo Platinum Graduate

Development Program (AGDP) currently operating in South Africa.

Ultimately, the responsibility for technical competency must be taken by individuals, but the

industry as a whole can contribute to the process by articulating the value and providing the

framework for training and career development.

STATEMENT OF THE PROBLEM

Nobody disputes the value of technical expertise. Time and time again, market research studies

come back saying that industry professionals want training, and they see the value of training, and

they would like courses available in a whole range of topics which they are happy to specify in long

wish‐lists. And yet training providers frequently have to cancel or postpone courses because of a

shortfall in registrations. Clearly there are some obstacles to take‐up of training, either in the

appropriateness of what is offered or in the appetite of the client.

Head office support is essential for any training program and technical training is no exception.

Most metallurgists understand the practical importance of technical knowledge about gold

mineralogy, flotation chemistry or energy efficiency in grinding and how this can contribute to

improved plant performance. It cannot be assumed, however, that this understanding is shared by

professionals in human resources or finance departments, many of whom might think processing

plants are just like tomato sauce manufacturing plants, with simple and homogeneous inputs and

outputs. The metallurgists are trained engineers, so surely they know how to run their plants!

Major decisions about graduate development and training are often made by people in these non‐

technical areas, so it is important to engage them on this topic rather than waiting for them to

somehow stumble onto it. Obstacle number one is therefore a failure to engage support from non‐

technical professionals.

Even with support from head office, however, there are still some obstacles to effective training for

early career professionals. The detail is covered elsewhere (Munro and Tilyard 2009, AusIMM et al

2001, Department of Industry Science and Resources 2002, Chamber of Minerals and Energy of WA

2008), but in summary we can say that the obstacles include smaller and busier workforces, a

changing balance between technical and non‐technical responsibilities, the impact of FIFO on the

way professionals organise their leisure time, a focus on narrow, short‐term performance targets

and many others.

The Julius Kruttschnitt Mineral Research Centre conducted a market research study in 2001 to

identify the major obstacles to take‐up of formal technical training (Drinkwater, 2001). Key findings


then remain as issues today. One issue was time ‐ specifically the difficulty in finding time for

production staff to leave site. Even on‐line courses were considered to be difficult to manage, as

they still require substantial time away from busy schedules.

Another issue was the need for technical training to compete with (often compulsory) courses in IR,

personnel management, OHS and other non‐technical areas. As these are often company‐

sponsored, they are given precedence over any other training. The HR professionals running

graduate development are well aware of the benefits of these types of training to a graduate’s on‐

the‐job performance. They are often less aware of the benefits of technical training and

development, believing as many do that new graduates coming into industry already have all the

“engineering” knowledge they need but in reality “operational” knowledge is almost non‐existent

and must be learned by other means.

Informal learning and training (mentoring), which was such a valuable part of many metallurgists’

development 30 years ago is much less significant now; as there are fewer potential mentors in the

modern process plant environment and those who are have less time to spend on research or

development projects that might provide opportunity for knowledge transfer.

This is not to say that young metallurgists are not learning about the technical aspects of their jobs

and making a contribution to the expansion of knowledge in flotation, comminution,

hydrometallurgy, mineralogy and more. Indeed they are, and our audience surely includes many

fine and knowledgeable early career metallurgists.

DEFINING APPROPRIATE TECHNICAL SKILLS

In the past it has been assumed that metallurgists need to be able to conduct accurate surveys of

their plants and use this information for optimisation studies, or run special laboratory or pilot

investigations to enhance various aspects of process performance. The first issue needing to be

addressed is whether the ability to do this kind of work is still relevant. If not, then the discussion

should be changed to considering not how to deliver training, but who should be taking

responsibility for technical development work.

Assuming there is still the desire to have this work done in plants by the operations personnel, the

next need is to identify the requisite skills. This has been done by Munro and Tilyard (2009), who

provided a very detailed list of specific skills under the heading “Mineral Processing Basics”. They

can be further categorised into the following groups:

Knowledge of specific processes (flotation, grinding),

Understanding of basic mineralogy,

Skills in collecting metallurgical data (surveys, equipment specs, ore characterisation),

Skills in handling numerical data,

Understanding of experimental method, precision, accuracy and error,

Process modelling skills,

Financial modelling skills,

Skills in problem solving, and

Writing clear and concise technical reports on which operational and management decisions

will be made.


This list can be further divided into two groups ‐ specific knowledge pertaining to mineral processing

and metallurgy (the first three) and more general engineering competencies (the remaining six). The

latter group should already be well developed in a general sense when graduates leave university as

these are doing an excellent job of developing well rounded engineers. However, they are not

necessarily applied in the metallurgical context, and some direction is often needed to demonstrate

metallurgical conventions, tools and applications.

The first three items on the list above relate specifically to the metallurgical field of study. Many

young graduates employed as metallurgists start work with limited or no specific technical training in

those first three areas. Many qualified as chemical engineers. According to figures from the

Minerals Tertiary Education Council (MTEC) in 2003, there were only 40 ‐ 50 graduates per year in

Australia from designated Minerals Processing and Metallurgy programs (Tuckwell and Way, 2003),

and numbers have remained relatively stable in the years since.

Chemical engineers may never have been exposed to analysis of non‐homogeneous systems, or

dealt with particulate material in a quantitative way. They often know nothing about the science of

rock breakage, and although they understand fundamentals of surface chemistry, know very little

about the practical aspects of mineral flotation.

On the upside, Chemical engineers often have very rigorous training in core engineering topics like

thermodynamics and reaction engineering.

Even within the Metallurgy programs, course content has changed, with a greater focus on

development of broad engineering skills and less on specialty skills development. Courses with

fewer than 30 students are discouraged by University administrators on a cost basis so many

specialised courses in comminution, flotation, practical hydrometallurgy, process mineralogy and

process modelling have become unviable.

The question has been asked in previous publications (Minerals Council of Australia, 2008) whether

minerals industry skills should be delivered as part of undergraduate training or after graduates

come into industry. The authors suggest that it makes more sense for the skills to be delivered after

a graduate begins work. Learning outcomes from a course in comminution modelling and

simulation, for example, are always better when the participants already have some knowledge of

the systems they are modelling and all their complexities.

In the past, individual employers have taken the lead in delivering this kind of skills development.

One example is the graduate training program run by Mount Isa Mines in the 1980s. A list of core

technical capabilities required by graduates includes (Munro 1990):

Size By Size Mineral Particle Behaviour,

Mass Balances,

Chemistry of the Process,

Laboratory and Pilot Plant Techniques,

Unit Operations and Processes,

Process Dynamics and Control, and

Experimental Technique.


A similar list of core competencies was developed by the AusIMM in the 1990s.

Structured programs ensured that graduates were given opportunities to develop their skills in all

these areas over their first two or three years of employment, and promotion to senior positions

would not be contemplated until graduates had completed this program. Cases where a graduate

was considered sufficiently “developed” to bypass this system were the exception, not the rule.

A list of technical reading was also recommended as follows:

“Engineering & Mining Journal” – monthly,

“Mining Engineering” – monthly,

“Mining Magazine” – monthly,

Proceedings of Australasian Institute of Mining & Metallurgy – quarterly,

“Canadian Mining & Metallurgical Bulletin” – monthly,

Proceedings of Institution of Mining & Metallurgy (London): Section C “Mineral Processing

and Extractive Metallurgy” ‐ ~ quarterly,

“Journal of the South African Institute of Mining & Metallurgy”,

“International Journal of Minerals Processing” – quarterly,

“Minerals And Metallurgical Processing” – quarterly, and

“Minerals Engineering” – quarterly.

Attendance at key technical conferences was encouraged, and conference proceedings from forums

such as the Canadian Mill Operators Conference were routinely circulated to technical staff.

In the leaner workforce environment of 2011, this model is no longer feasible as the resources are

simply not available. In this paper, some alternative models for delivering the same sort of skills

development are examined.

A TRAINING MODEL FOR THE GRADUATE METALLURGIST IN 2011

Today’s young metallurgists and process engineers are busier, carry more responsibility, and have

more autonomy than the previous generation and often are employed on a FIFO or 12 hour shift

basis. Typically they would come directly from university into a production‐based role, where they

take direct responsibility for performance of a section of process.

Mostly they are four year trained engineers or applied scientists with honours, though significant

numbers do have three year degrees. As mentioned, university courses have changed over the last

30 years. They would have been in larger classes than the previous generation, encountered smaller

staff/student ratios, and had considerably fewer contact hours while engaged in study. A really

significant difference is that they will have had much less laboratory time. Metallurgists studying 30

years ago would have spent up to 10 hours a week in laboratory and practical sessions with a high

staff to student ratio, while many current programs have fewer than 10 hours in a semester in

laboratories. Laboratory work, when it is done, is increasingly in the form of “demonstrations” to

deal with the large student group numbers, so there is even less exposure to hands‐on activities.

The trade‐off for this reduction in practical skills is supposedly a more integrated approach to

developing graduates as complete engineers and life‐long learners (Duderstadt, 2005). Universities


have a stated commitment to producing graduates with a specified set of attributes like the set

below (from The University of Queensland School of Engineering):

In‐depth knowledge of the field of study,

Effective communication,

Independence and creativity,

Critical judgement, and

Ethical and social understanding.

The aim is to produce engineers with the skills required to continue developing and learning

throughout their lifetimes, rather than fully‐formed engineers who know everything they need.

A training program needs to provide specific metallurgy content for chemical engineers (and

metallurgists who weren’t paying attention in their undergraduate years!) and provide opportunities

for practical application of the knowledge delivered in the program as well as opportunities to

expand and develop knowledge and skills already in place. Content is important (for example the

explanation of relationship between energy input and size reduction in different kinds of

comminution machines) as is application of this knowledge (laboratory tests and scale‐up for

different comminution machines). The best learning outcomes are achieved when the learner is

given an opportunity to assimilate the new knowledge with pre‐existing knowledge and create their

own unique solutions to a problem (how does what I just learned help me operate my grinding

circuit more productively?).

WHO IS RESPONSIBLE FOR GRADUATE DEVELOPMENT?

This paper makes no attempt to give a definitive answer to the question above. Companies and

organisations address graduate development in many different ways. Often responsibility is shared

between graduates, their immediate supervisors and human resources professionals at head office.

Implementing a comprehensive, structured program for a cohort of graduates will be easier in

companies where head office commands a significant portion of the training budget. On the other

hand, a graduate with specific training requirements is more likely to get support when his or her

supervisor has control of the training cheque‐book. Immediate supervisors are also more likely to

recognise an individual area where training can have an impact, for example process control.

In any case, however, it is essential that HR professionals in an organisation are aware of the

potential impact that good technical understanding of process characteristics can have on overall

productivity and profitability of an operation.

FEATURES OF GOOD DEVELOPMENT PROGRAMS

The requirements for a good quality graduate development program should be no different from

requirements for any other good quality educational experience. The course content needs to be

current, relevant and clear, and the presenters need to know their topic well. Good quality program

delivery, however, is the key to really good learning outcomes.


For best results, a program should support learning with practical activities that are relevant to the

learner. The catch‐phrase “what the student does is actually more important in determining what is

learned than what the teacher does" (Shuell 1986) is more relevant now than ever before.

If the learning activities in a graduate development program are directly relevant to that person’s

day‐to‐day responsibilities, there is even more likelihood that the learning will stick, and there is an

added benefit that the activity is likely to be supported by the graduate’s peers and supervisors.

Good learning outcomes are also associated with a problem‐based approach. Here the learner is

presented with a problem and given access to resources, without being told exactly how to develop

the solution. This is an ideal scenario for workplace‐based training, as most sites would have little

difficulty identifying an area that could be used as a focus for such a learning activity. A graduate

learning about flotation chemistry could be set a task based on optimising cleaner performance or

increasing selectivity in a scavenger bank, and learning would be enhanced by better knowledge

retention, deeper understanding of applications and more opportunities for creativity and

innovation (Biggs 1999).

Another feature that can add value to the learning experience is group work. Putting learners into

groups adds a social element to the problem solving process, meaning that individuals have to

negotiate the scope and then the solution, leading to better engagement with the problem and

more sophisticated outcomes (Johnson and Johnson 1999). The flotation cleaner optimisation task

would now involve detailed discussions about strategy, prioritising options and progress reviews. In

the longer term, the collaborative experience should lead to development of a community of

learners who become a support network for each individual in any future activities.

The learning process still needs facilitation, and learners require prompt and effective feedback from

subject experts as they work through any problem‐based task. Resources also need to be easy to

find and properly explained, so there is still scope for some classroom‐type activities to familiarise

the participants with appropriate tools and techniques.

In summary, a good graduate development program can be delivered to a group, be activity‐based

and be focussed on solving a workplace related problem. Some classroom time can be devoted to

teaching fundamental skills or techniques, but the bulk of the learning will actually take place when

participants put their learning into context.

An example of a scheme that is built on these principles is discussed in the next section. The ideas

also form the basis of the MetSkill model that JKTech is developing, and discussing with a number of

corporate clients. For JKTech there is an added benefit of this model; development of relationships

between researchers and practitioners that can increase awareness of new developments and

innovations in the field.

OUTSTANDING EXAMPLES Some outstanding metallurgy development examples are given here, including a long term

integrated workplace training program and others that expose students to high level plant

optimisation project work. In all cases the students or graduates are placed into a plant

environment and actively engaged in investigative work.


The Anglo Platinum Graduate Development Program (AGDP) in South Africa is now in its 8th year.

The program was designed to deliver formal, advanced education for technical graduates. The

company entered into an agreement with the University of Cape Town who have been managing

and delivering the program (Sweet et al., 2006).

The program runs over two years, with approximately 20 new graduate metallurgists in each intake.

Every year the company selects a plant to be the focus for a whole group activity – a major survey of

comminution and flotation sections. Over 11 weeks of each year is spent in classes, seminars and

workshops learning about experimental method, comminution, flotation, mineralogy, mass

balancing and numerical modelling. A rigid formal assessment program is included, and each

module contains exams and assignments. Performance in the program is linked to promotional

opportunities and graduates are not engaged in major production roles until they finish.

The company formerly known as MIM or Mount Isa Mines Limited has also run some very good

graduate development programs, including a program starting in the 1960’s which employed

vacation student groups under the leadership of a relevant university staff member to examine a

section of operating plant in detail with the intention of then using the data to improve plant

operations. Apart from OH and S and training in plant operating rules, the groups operated

independently of plant personnel with the exception of a liaison metallurgist. Areas studied in this

way included the zinc flotation circuit, the copper refinery tank‐house and the lead smelter. Several

students involved in these activities later joined MIM as graduate metallurgists, with obvious

advantages to the company.

THE METSKILL MODEL The MetSkill program run by JKTech is an attempt to create a similar program in the Australian

context. This model was discussed by Viscarra in a paper at the AusIMM New Leaders conference in

2011 (Viscarra and Drinkwater, 2011). Key elements of the program are:

Graduates work in groups,

A plant is selected for a whole group activity,

The training program is integrated with a plant survey, and graduates submit a plant

optimisation report on completion,

Long timeframe, say one or two years,

Incorporate short workshops in comminution, flotation, mineralogy, statistics, mass

balancing, and

Provide competency based assessment. No exams; pass or fail based on plant

optimisation report.

CONCLUDING REMARKS The focus of this paper is more about raising questions than answering them. After a decade of

substantial transformation in the workspace, and even more substantial transformation in our

tertiary education institutions, the authors believe that the time is ripe for Mineral Processing

practitioners to have a good hard look at their discipline.


Is the general level of knowledge in our processing plants about the latest ideas, equipment and

processes all that it is supposed to be? Can today’s metallurgists sense a good idea or are they

afraid to be the first “off the rank” when something new comes along? More importantly, are they

sufficiently aware of leading practice to ensure that opportunities are not missed for major plant

improvements? If the answer to these questions is “yes”, then nothing needs to be done. If not,

then action is needed now to ensure that operations are achieving maximum return to the bottom

line.

On‐site training will always occur but this is not the same as expert mentoring. Access to potential

mentors and facilitators with the knowledge needed is a critical need to achieve this, but they are a

dying (or at the very least, retiring) breed. If industry is serious about capitalising on the current

commodity price boom there should be a move to capture these potential mentors and identify how

best to involve them with incoming metallurgists.

REFERENCES 1. Australian Institute of Mining and Metallurgy and the Department of Education, Training

and Youth Affairs, 2001, Rising to the Challenge: Building Professional Staff Capability in the Australian Minerals Industry for the New Century

2. Biggs J, 1999 Teaching for quality learning at university, Buckingham, SRHE and Open

University Press.

3. Brown J S, Duguid P 2002, The Social Life of Information Cambridge, MA: Harvard

Business School Press.

4. Chamber of Minerals and Energy of WA, 2008, Submission to the Review of Australian

Higher Education, Discussion Paper

5. Department of Industry, Science and Resources, 2002, Mining Technology Services Action Agenda ‐ Background paper on issues affecting the sector

6. Drinkwater D, 2001, Report on visit to N‐W Queensland Mine Sites, Internal Report,

JKMRC

7. Duderstadt JJ, 2005, Engineering Research and America’s Future: Meeting the

Challenges of a Global Economy, University of Michigan Millennium Project; 2008

8. Johnson D; Johnson R, 1999, Making Cooperative Learning Work, Theory Into Practice,

Spring 99, Vol. 38 Issue 2

9. Minerals Council of Australia, 2008, Higher Education Review Submission, Discussion

Paper

10. Munro PD, Tilyard PA, 2009, Back to the future – why change doesn’t necessarily mean progress, Tenth Mill Operators’ Conference, Adelaide, South Australia.

11. Munro P, 1990, Personal communication

12. Shuell T J, 1986, Cognitive conceptions of learning, Review of Educational Research, 56,

411‐36.


13. Sweet CG, Sweet JA, Harris MC, Powell MS, Lambert AS & Knopjes LM, 2006, Industry

taking the initiative in developing high calibre technical staff, Proceedings of XXIII

International Mineral Processing Congress, Istanbul, Turkey, 3‐8 September.

14. Tuckwell K, Way A, 2003, Where Do Graduates Go? in The AusIMM Bulletin No. 5,

September/October

15. Viscarra T, Drinkwater D, 2011, Challenges and Opportunities in the Development of

Technical Expertise in the Minerals Industry, Paper presented at The AusIMM New

Leaders’ Conference, Newcastle, New South Wales.



Differences between the Engineering Cultures of Australia and Brazil

C Fountain1, P Libânio2 and G Lane3

1. MAusIMM, Senior Technical Consultant/Senior Study Manager, Ausenco Services Pty Ltd, Level 1, The Podium, 44 St Georges Terrace, Perth WA 6000. Email: [email protected]

2. Regional Director, Ausenco do Brasil Engenharia Ltda, Rua Pernambuco 1077, 6° andar, Funcionários, Belo Horizonte CEP 30 130-151, Minas Gerais, Brazil. Email: [email protected]

3. MAusIMM, General Manager – Technical Solutions, Ausenco Services Pty Ltd, 144 Montague Road, South Brisbane Qld 4101. Email: [email protected]

ABSTRACT

Australian mining and engineering companies have increasingly been investing in the mining

industry in Brazil. It is a country of enormous mineral and agricultural wealth, with a well-

established mining sector. Some Brazilian mining companies and personnel have been

journeying in the opposite direction. There are significant differences in the approaches to

engineering between the Australian and Brazilian industries that can cause confusion for

companies operating in the other country. These range from differing expectations during the

various stages of feasibility studies to style of plant and contracting strategies during the

construction phase of projects. Some of the differences are driven by tax regimes, labour

costs, or different settlement patterns. Others are the result of different histories and pressures

on the industries in the two nations.

This paper describes some of the differences in approach found in the two countries and some

of the reasons for them.

INTRODUCTION

Brazil is a rapidly growing nation of 194 million people (World Bank, 2011). It is ranked

fifth in the world by area (being about 11% larger than Australia), had a gross national

income per head of US$8040 in 2009 and, like Australia, features the Southern Cross on its

flag.

Brazil has a much longer history than Australia, being initially colonised by the Portuguese in

1500. Mining has been an important part of the Brazilian economy throughout much of that



history. Like Australia, gold was an important driver of early settlement, with a gold rush

triggered by the discovery of gold in what is now the state of Minas Gerais (‘General Mines’)

in 1695. Gold and later diamonds, extracted from places with names like Ouro Preto (‘Black

Gold’) and Diamantina, were brought down the Estrada Real (‘Royal Road’) through Minas

Gerais to Rio de Janeiro.

The mineral wealth of Brazil supported the Portuguese crown and led to the transfer of the

colonial capital from Salvador to Rio de Janeiro. The capital eventually shifted from Rio de

Janeiro to the present-day capital of Brasília, which, like Canberra, is a city created from

nowhere to be a national capital.

The similarities with Australia do not end there, nor with the eucalypt trees lining Brazilian

and Australian roads. As in Australia, the importance of gold mining has declined, but mining

iron ore has become an important part of the economy and companies in both countries find

their ability to export iron ore constrained by lack of port and transport infrastructure. Both

countries play host to major global mining companies, with Brazil’s Vale competing with

Australia’s BHP Billiton and Rio Tinto as one of the world’s three largest diversified mining

companies, all three operating in many countries. The currencies of both countries have been

buoyed against the US dollar by the high prices of their mineral exports, which helped them

avoid most of affects of the recent ‘global financial crisis’, and both have benefited

significantly from the rise of the Chinese economy.

There has been movement of mining companies in both directions across the Pacific, with

Australian companies, large and small, attempting to develop mines in Brazil and with Vale,

in particular, moving into Australia. Global engineering companies, including Australian

players, have also seen opportunities in Brazil and either purchased local engineering firms or

set themselves up to grow organically. Mining and engineering professionals have also been

crossing the Pacific to work in each others’ homelands.

However, despite the similarities between the industries of the two nations, there are some

significant differences in the engineering cultures of the two countries that extend beyond

speaking different languages. These differences can lead to misunderstandings and

frustrations between people and organisations that have made the journey from one continent

to the other.



This paper describes some of these differences from the authors’ perspective. It is hoped that

this is a helpful guide to industry professionals and organisations looking to work in the other

country. Please note that this is not intended to present the approach in one country as being

better than that of the other as they are simply different for a variety of reasons that include

history, industry structure, taxation and labour costs. However, changes in the economic and

social environments of both countries might mean that each could learn from the approaches

of the other.

The key differences described include:

the approach to feasibility studies,

the type of engineering design, and

the project implementation phase.

THE APPROACH TO FEASIBILITY STUDIES

Differences in the approach to feasibility studies start with the tendering process and extend

to the content and management of feasibility studies.

The tendering process

Tendering for feasibility studies in both nations commonly involves sending a request for

proposal to a group of engineering companies experienced in the delivery of feasibility

studies. However, there is a significant difference in the level of information required by most

Brazilian mining companies and in the style of the proposal documents.

Deliverables

The typical practice in the mining sector in Brazil is to pay by deliverables rather than hours,

and a deliverable is measured as the equivalent of an A1 drawing. This means that all

documents have to be converted to an A1 equivalent (8 A4 pages = 1 A1 drawing, 4 A3

pages = 1 A1, etc). The cost of the proposal is calculated as an all-up A1 equivalent cost.

Table 1 shows a list of activities extracted from a Brazilian feasibility study proposal.



Table 1 – A list of activities from a typical Brazilian feasibility study proposal

Activity A1 A4 Total equivalent A1

Coordination and planning of multidisciplinary activities 480 60.00

Process and systems 20 700 107.50

Piping 5 60 12.50

Mechanical 35 95 46.88

Structural and civil 2 255 33.88

Architectural 2 35 6.38

Infrastructure 2 40 7.00

Electrical 2 100 14.50

Instruments and control 1 100 13.50

Cost engineering (including procurement support) 370 46.25

Total 69 2235 348.38

The concept that a 16 page options study is equivalent to two A1 drawings seems strange to

people raised in the Australian engineering environment. The amount of time, thought and

effort required could be significantly greater than that required to produce a couple of plant

general arrangement drawings. Also, the approach of paying for the number of A1 drawings

produced does seem to provide an incentive to overproduce drawings to maximise income.

However, in the context of the role of engineering companies work in the Brazilian industry,

it does make some sense, as will be discussed later.

Some mining companies go as far in the requests for proposal as prescribing the number of

deliverables and the number of hours to be spent on each.

Lump sum studies

The next key difference in the tendering process is that Brazilian companies almost always

require lump-sum (‘preço global’) contracts for studies, while reimbursable contracts are

more common in Australia.

The preference for lump-sum contracts in Brazil is largely driven by the complex taxation

system, where taxes are payable monthly to various levels of government and where

forecasting of the next month’s income is very important. A Brazilian engineering company

cannot invoice a client for an amount exceeding the initial contract without an official

variation document because of the tax implications. In discussions with Brazilian clients, the

comment has been received that a reimbursable approach is okay, as long as it does not

exceed a certain fixed number.



Each approach has advantages and disadvantages. The key advantage of lump sum contracts

to the client is that it provides a greater degree of certainty over the cost of the study.

However, the disadvantage is that it reduces the flexibility required of studies, particularly in

the early stages of the feasibility process where an examination of options is recommended.

The number of tenderers

Another difference in tendering approaches can be the number of engineering companies

asked to tender for a study. In Australia, mining companies will often approach an

engineering company with whom it has worked satisfactorily previously and negotiate a price

for a new study or even an EPCM (engineering, procurement, construction and management)

project, sometimes through an open-book approach. This is unusual in Brazil, where tenders

are normally called for each stage of a project, from concept study through to final

construction.

Going to a full tender for each stage of a project has several disadvantages, including the

delays inherent in issuing requests for proposals, answering questions from potential

tenderers, assessing each tender, seeking clarification, making an award and agreeing the

wording of the final contract. In our experience, it typically adds two to three months before

work can start on a study. In addition, potential contractors understand that there are likely to

be multiple tenderers and minimise their costs by producing boiler-plate proposals that might

not address the client’s needs, and pad them with contingency for the lump sums bids

required.

Another major disadvantage is the lack of continuity in the engineer from one phase to the

next. This means additional time lost as the new team gets up to speed, and some rework due

to the inevitable differences in design philosophies between various engineering companies.

The tendency to seek multiple tenderers in Brazil is, in part, driven by its historical culture of

corruption and kick-backs. There is a strong desire to minimise the opportunity for this to

occur. Consequently, there is a large focus on ‘proper process’ when assessing tenders and

awarding contracts.

Assessing the proposals

The focus on proper process in Brazil has resulted in differences in the style of proposals

between the two countries.



In Brazil, two separate proposals are normally produced—a technical proposal and a

commercial proposal. This frequently extends to quotations from equipment suppliers, for the

same reasons. In Australia, there is commonly only one document, particularly for studies.

The client submits each of the proposals to different teams. Normally, the technical group

examines the technical proposals and makes a recommendation, and then the commercial

group looks at the cost.

Technical proposals usually describe the scope of services, the engineer’s team and the team

members’ qualifications, the schedule and the deliverables. The commercial proposal

includes the cost of the work, the hourly rates, the taxes, expenses and the terms and

conditions.

The specification that the deliverables be reported as an A1 equivalent allows the technical

team to make a rapid comparison between the various proposals.

Often, the cheapest tenderer wins the work, regardless of prior history with the client.

However, considerable weight is put on the average hourly rates and the cost per A1 drawing

in case there are variations. Thus, the company that produces the cheapest overall tender for

the prescribed scope of work might miss out because the average hourly rate is greater. This

can have the effect of penalising engineering companies with greater hourly productivity (and

lower lump sum cost) but higher average rates.

Again, the Australian approach can be different. Australian engineering companies will often

put their most experienced (and therefore more expensive) people on a study team to increase

productivity and the quality of the study, particularly options studies, as this is where the

chance to add value to a project is the greatest. In Brazil, this approach can disadvantage a

tenderer, in part because the expectations of the role of the engineering company are

different.

Conducting studies

The approach to studies

There are often differences in the approaches to studies in the two countries.

To some extent, the approach to engineering and feasibility studies in Brazil is similar to that

which prevailed in Australia two or three decades ago. The Australian mining and



engineering industry was subjected to pressures that are only now beginning to be felt in

Brazil.

When the authors began their careers in the mining industry, it was common for mining

companies, particularly the larger ones, to develop and manage projects themselves.

However, during the lean times of the 1990s, Australian companies began to outsource much

of this engineering work as they struggled to survive and contracting became the norm for

many service providers in Australia. EPCM contracts became much more common, with the

engineering companies expanding into procurement, construction and project management.

These days, it is common in Australia for engineering companies to provide full project

services, including developing and managing test work programs, selecting flow sheets,

purchasing equipment, and managing construction contractors. The process design and

engineering capabilities of Australian minerals industry engineering companies is often well

developed.

In contrast, Brazilian engineering companies employ fewer process engineers. The clients are

normally responsible for the metallurgical test work and much of the process selection, either

through their own staff or through consultants. In feasibility studies, the engineering company

is typically expected to focus on sizing the equipment for the flow sheet that has been

specified by the clients, producing the associated engineering drawings, and producing

capital and operating cost estimates. Combined with the emphasis on measurement by

drawings, this means that an engineering company in Brazil typically makes a smaller

contribution to the process optimisation and engineering than in Australia.

Some Australian clients are also responsible for metallurgical test work (and/or use

consultants for this purpose) and specification of the flow sheet. However, they often expect

some form of warranty from the engineering company. The merits of such warranties are

debatable and have been discussed elsewhere (Lane et al., 2007), but this expectation

increases the level of involvement of the engineering company in the review of test work

outcomes. Australian clients often want the engineering company to contribute to, and sign-

off on, the metallurgical test work as part of a study.

It is uncommon for Brazilian engineering companies to provide such warranties. One client

even expressed disbelief at the very concept that an engineering company would do such a



thing. Consequently, Australian engineering companies working in Brazil experience greater

resistance from clients to the offer of input into process design.

The level of detail in studies

The various levels of studies—concept study, prefeasibility studies and feasibility studies—

typically have a greater level of engineering in Brazil than in Australia. A tendency to

investigate the engineering of options to greater detail in Brazil than in Australia before they

are discarded has been observed.

This is partly due to the lower cost of engineering labour in Brazil. The shortage of

engineering and drafting labour (and the consequent high cost) in Australia has placed the

emphasis on doing the optimum amount of work necessary to achieve the level of accuracy of

the estimate.

One example of the difference is the requirement of most Brazilian clients that the

engineering company produce piping and instrument diagrams (‘P&IDs’). While many

Australian engineering companies would produce P&IDs for a feasibility studies of complex

plants in which they have little experience, P&IDs are considered a waste of time and money

for an ordinary gold or base metals plants. The level of engineering for a feasibility study is

insufficient to produce the P&IDs that might be required for construction purposes, so they

would have to be redone during the detailed engineering phase of the project, and experience

has shown that it is often more accurate to use the factors for piping and instruments obtained

from benchmark projects that have already been constructed.

The cost of P&IDs is not trivial. If a plant design has 20 process flow diagrams (PFDs), there

would be a minimum of approximately 40 P&IDs. At 50 hours each to draw, this totals about

1000 hours, or about A$150 000 for a set of drawings that are not used for anything. The cost

is higher when process and instrument engineering input is considered. Hence, while the

precision of engineering output is increased, accuracy of subsequent cost estimates may not

been improved.

Another area of difference is the inclusion of the mass balance on the PFDs. While this is a

convenient way of conveying information, it can add unnecessarily to the cost of the study.

Every time there is a change to the mass balance as additional information becomes available,



the drawing has to be updated as well, and there is the potential for transcription errors and

inconsistencies between various documents.

Australian engineering companies try to minimise the number of drawings, while the

payment by drawing emphasis in Brazil provides and incentive to maximise the number.

Estimating

Many Australian companies work with earthworks, civil, construction, and mechanical and

electrical installation contractors to estimate the cost of building plants. This enables a gauge

of the productivity of construction workers in various parts of the world.

In Brazil, estimators typically determine the costs by building them from first principles,

including known labour rates (which are set by union agreements) and allowing for taxes,

expenses and the contractors’ margins.

Trying to obtain similar information to that which can readily be obtained in Australia and

despite the fact that Brazilian contractors approached were asked not to provide quotations

for the job, this only resulted in lump sum quotations with insufficient detail to determine the

rates. Incidentally, the quotations varied enormously, from R$80 million to R$210 million for

the same scope of work.

A second difference in the estimating process is coping with the Brazilian system of taxes.

The Brazilian tax system is diabolical in comparison to the Australian tax system, consisting

of an alphabet soup of federal, state and local government taxes.

Table 2 lists the taxes that are directly relevant for estimating purposes.

Note that many of the taxes in Table 2 are variable. For example, the federal government is

responsible for the IPI and II taxes. These vary depending on the item. To determine the tax

rate, it is necessary to know the NCM code (Nomenclatura Comum do Mercosul – a standard

code used in the Mercosul trading bloc of South America) for the item of equipment. This

code is based on the internationally-accepted Harmonised Commodity Description and

Coding System, with two extra digits added. The NCM code is available on Brazilian

government internet sites.



Table 2 – Brazilian taxes

Tax Full name Jurisdiction Rate (%)

PIS Programa de integração social – Social integration program Federal 1.65

COFINS Contribuição para o financiamento de seguridade social – Social security contribution

Federal 7.60

IPI Imposto sobre produtos industrializados – Tax on industrial products

Federal Variable, mostly 0% to 20%,

depending on item

II Imposto de importação – Import tax Federal Variable, depending on

item

ICMS Imposto sobre circulação de mercadorias e prestação de serviços – Goods and services tax

State Variable, depending on

state

ISS Imposto sobre serviços – Municipal tax on services Municipal Variable, depending on municipality

The ICMS tax varies between states. If an item is purchased from a state with a lower rate of

ICMS and shipped to a state with a higher rate, the differential between the two rates has to

be paid to the government of the recipient state.

The ISS rate varies from municipality to municipality, and it is levied by the municipality

within which the service is undertaken, not the municipality within which the person

performing the service is normally based. Thus, if a person living and working in Belo

Horizonte in Minas Gerais travels to Cuiabá in Mato Grosso to perform a service, the ISS for

that service is paid to the Cuiabá city council.

To further complicate matters, there are various exemptions, rebates and concessions that

might be applicable, depending on the region in which the mine is to be developed, and

whether the mineral products are being exported or consumed locally.

There is also a difference between the way the taxes are calculated in Brazil and Australia. In

Australia, the Goods and Services Tax is 10% of the untaxed price. In Brazil, most of the

taxes are expressed as percentages of the final selling price, and the absolute value of one tax,

say ICMS, might depend on whether there is a concession on other taxes, say PIS and

COFINS. The effect of IPI, if it is payable, on the ICMS, PIS and COFINS depends on

whether the item being taxed is considered to be consumed in the production process, as there

are two different methods of calculating the IPI’s contribution to the taxable total.



The ICMS levied by the government of the state of origin is levied as a percentage of the

selling price in that state, but the ICMS differential levied by the government of the recipient

state is also a percentage of the original selling price.

The rules for calculating PIS and COFINS were changed some years ago. However, some

companies still pay these taxes at the old rates of 0.65% and 3% that formerly applied when

the tax was cumulative rather than the current non-cumulative nature.

Quotations from equipment suppliers usually state the tax assumptions made in calculating

the selling price of the equipment. However, the assumptions vary from supplier to supplier.

Some suppliers do not include taxes, some include all the taxes, and some assume that the

client will be eligible for the concessions that are available. The quoted prices cannot,

therefore, be compared directly. They first have to be normalised for the applicable taxes.

Determining the applicable taxes can take two to three weeks, and should involve discussions

with the client to determine the applicable concessions for which it is eligible. These

concessions are granted by the Receita Federal do Brasil on application by the client, and

their existence must be noted on purchase orders to receive them.

Implementation

EPCM projects

The practice of EPCM contracts is common in Australia, but rare in Brazil.

This is partly due to the different industry structure and history and partly for tax reasons.

Brazil’s mining industry was, until recently, largely state controlled. Even now, the Brazilian

government remains the major shareholder in Vale, which it used to run as Companhia Vale

do Rio Doce (‘CVRD’). As such, Vale runs its own projects, much as CRA and BHP used to

in Australia. It has a strong procurement arm and project management team.

Since much of the work was, and still is, with Vale, engineering companies in Brazil had

little incentive to develop their own procurement and project management expertise.

In addition to this history, the way the Brazilian tax system operates discourages the EPCM

approach. Every time an invoice is issued, taxes are payable. If a contractor subcontracts part

of the job and is invoiced by the subcontractor, taxes are payable. Then when the contractor



invoices the client, taxes are again payable. There is a strong incentive to minimise the levels

of invoicing during a project.

The EPCM approach is used in Australia, particularly by small and mid-tier mining

companies, to avoid the problems of directly hiring project staff when there might be a long

drought between jobs.

The typical approach in Brazil is to use an engineering company to provide the engineering

design and drawings. The client then engages a construction company. Earthworks and civil

engineering is often done by one construction company and mechanical and electrical

construction by another. The client normally pays for fabrication of equipment, structural

steel and other materials directly, to minimise the number of invoices involved and to take

advantage of any tax concessions that might apply.

Standards of construction

In Brazil, mine infrastructure is ‘built to last’. Even short-life mines tend to have a high

standard of building, with large workshops, solid office buildings and restaurants, and

buildings to house the processing equipment. This is reminiscent of the glory days of Mount

Isa or Broken Hill, with enclosed halls for the grinding mills and flotation cells.

A cost-effective approach has become more common in Australian engineering. In part, this

arose during the 1980s as small, short-life gold mines were being developed (Close, 2002),

and it was realised that the standards that applied to long-life ore bodies, such as those of

Mount Isa and Broken Hill, were not affordable.

In Australia, it is now common to construct mines with modular transportable offices that can

be moved to the mine site as flat-packs or on the back of trucks. The same applies to control

rooms, and MCC (motor control centre) rooms can be assembled in a major city and

transported to the mine site for connection to the electrical systems. Mills, pumps and other

equipment are often installed with minimal covering. In one plant recently constructed by

Ausenco, the vehicle workshop was effectively stacked sea containers and canvas (see

Fig. 1).



Fig. 1 – Low-cost workshop design

The housing of equipment in buildings means that it is difficult to use mobile cranes for

maintenance. Consequently, gantry cranes and smaller hoists are more common in Brazilian

plant designs than Australian ones. These can add to the cost of the plant.

It is also more difficult in Brazil to hire mobile equipment. Rental costs can be very

expensive, with the payback time for buying equipment often being less than a year.

Prefabrication is less common in Brazil. For example, the ‘donga’ that is a mainstay of

Australian mines has not been seen on any Brazilian mine sites. An attempt to obtain

prefabricated MCC rooms resulted in a double-walled steel arrangement that was more

expensive than the common Brazilian concrete block construction.

One of the key reasons for the lack of prefabrication in Brazil is the relatively low cost of

construction labour. In Australia, the emphasis is on doing as much offsite work as possible

to minimise the number of people required on site. This is less of an issue in Brazil. In

addition, the condition of the roads in many parts of the country can make it difficult to bring

prefabricated structures to the site. It can be very difficult to move such items as excavators

through the narrow streets of some villages.



Another difference observed is the construction of ROM pads. In Australia, concrete is very

expensive and efforts are made to optimise its use. However, this is not evident in Brazilian

constructions, where large concrete retaining walls are commonly used on ROM pads. An

approach increasingly favoured by engineering companies in Australia is to use concrete

vaults with reinforced or unreinforced earth walls, again to minimise the use of concrete.

Fig. 2 is an example of a crusher and ROM pad wall recently constructed by Ausenco.

Fig. 2 – Recently-constructed crusher and ROM pad

CONCLUSIONS

There are some significant differences between the approaches to projects in the Brazilian

and Australian mining industries. These have arisen for a variety of reasons, including

history, labour conditions, and taxation regimes.

In Brazil during the study phase, the owners are typically responsible for:

sample selection and metallurgical test work,

flow sheet development,



conceptual design,

developing the project design basis,

developing the contracting and implementation plan, and

study management, including supervision and management of the subconsultants.

Brazilian engineering companies are typically responsible for:

engineering deliverables, including equipment lists, data sheets and layouts, and

capital cost estimates, including material take-offs, rates and equipment pricing.

During the study phase, Australian engineering companies are typically responsible for:

sample selection and metallurgical test work,

flow sheet development,

conceptual design,


engineering deliverables,

developing the contracting and implementation plan,

capital and operating cost estimates, and

study management.

In Australia, the owners are typically responsible for the overall study management and some

supervision and management of various subconsultants (e.g. geology, mining and

environmental work).

During the engineering phase in Brazil, the owners are typically responsible for:


procurement,

contracting strategy,

construction management, and

commissioning.



During this phase, Brazilian engineering companies’ services are typically limited to:

front-end engineering design and detailed design, and

assistance with the development of procurement packages.

Most Australian engineering companies can provide full EPCM services, including:


front-end engineering design and detailed design,

full procurement services,

contracting strategy,

project management,

construction management, and

commissioning services.

Owners are typically responsible for:

permitting and legal,

review and approval of designs and drawings,

sign-off on purchase orders for equipment and contractors, and

operational readiness.

The Brazilian mining industry is experiencing some of the pressures that have driven the

Australian industry toward the leaner approach to feasibility studies, engineering and

construction. As Brazil’s economy grows, there are growing shortages of skilled engineering

and drafting personnel. Salaries are rising as the standard of living increases.

The authors expect that, over time, there will be some convergence in the engineering

approaches of the industries in the two nations. This is already evident in some of the smaller

gold projects, where Brazilian and international companies operating in Brazil, are looking

for new approaches to reduce capital and operating costs, and to increase profitability. The

inertia in the Brazilian system due to government regulation may counteract some of the

potential for change.



REFERENCES

Close, S.E., 2002, The Great Gold Renaissance, The Untold Story of the Modern Australian Gold Boom 1982–2002

(Surbiton and Associates, Melbourne).

Lane, G, Davis, M, McLean, E and Fleay, J, 2007. Performance testing—when, what and how? in Proceedings Project

Evaluation Conference 2007, pp 197–202 (The Australasian Institute of Mining and Metallurgy: Melbourne).

World Bank, 2011. Brazil. Available from: <http://data.worldbank.org/country/brazil> [Accessed 26 March 2011].



Crushers – An Essential Part of Energy Efficient Comminution Circuits

R Bearman1, S Munro2 and C M Evertsson3

1. MAusIMM, Director, Bear Rock Solutions Pty Ltd, PO Box 150, Melville WA 6956. Email: [email protected]

2. MAusIMM, Director, Met Dynamics Pty Ltd, PO Box 204, Woodvale WA 6026. Email: [email protected]

3. Associate Professor, Chalmers Rock Processing Research, Machine Elements, Product and Production Development, Chalmers University of Technology, Göteborg SE-41296, Sweden. Email: [email protected]

ABSTRACT The importance of crushers in the mining industry declined with the introduction and resultant

dominance of AG and SAG based circuits, but the recent move to consider more energy efficient

circuits has caused the industry to re-consider the role of crushers.

Different types of comminution equipment have different efficiency in the way the energy is applied in

the size reduction process. Because of the general trend towards more sustainable production

processes and emission of carbon dioxide, energy use has become more important. To fully

understand how crushers can contribute to optimising the energy application in comminution circuits,

detailed models of process performance are required. Such models need to take into account the

mechanical design factors and how the crusher performance changes in response to feed conditions,

the control strategy utilised and the wear of the crusher liners.

It is also critical to analyse the total circuit, not just the individual crushers. To fully understand the

performance of the entire circuit, dynamic simulation should be applied. Without the holistic

understanding of the total circuit and the incorporation of transient effects, the full circuit performance

cannot be adequately determined and the true energy picture will not be quantified.

Based on detailed mechanistic modelling of crushers by the authors and the application of dynamic

simulation, the opportunities to optimise the use of crushers for energy efficiency are examined.

INTRODUCTION Crushers, particularly cone type machines, were the mainstay of the hard rock concentrator

comminution plants from the development of the Symons crusher through to the 1960’s. It was

common to see entire coarse comminution plants consisting of secondary, tertiary and quaternary

crushers feeding ball mill circuits. The advent of SAG mills greatly reduced the deployment of

crushing plants as part of base metal concentrators. The main reasons were reliability and

robustness, simplicity of the process layout, investment and operating costs. Over the years the

importance of overall energy utilisation and energy costs have increased.



Crusher application became restricted to primary and secondary crushing and in some instances,

pebble crushing associated with SAG circuits. As a result of this more limited deployment, the

interest in cone crushers for mining became restricted to more specific applications such as

lump:fines iron ore operations, where a coarse product size was required.

Manufacturers continued to develop new and improved crushers, but the focus on their performance

and how they could be optimised in their own right, or as part of the overall system, declined.

It has been known for a considerable time that the breakage of rocks in compression based cone

crushers is, relatively speaking, energy efficient. Typically, the breakage efficiency of cone crushers

is 60-80% compared to ideal single particle breakage. By comparison, the figure for AG or SAG mills

is typically 20-40%.

This efficiency is greatest for larger particles, i.e. secondary crushing, but declines through

subsequent stages of crushing. By the time material reaches the quaternary stage the efficiency has

dropped significantly. The most efficient breakage mode in compressive crushing is single particle

breakage which occurs when a rock particle is stressed to fracture between two crushing plates i.e.

mantle and concave.

The increased focus on energy efficiency has led many people to re-examine the role of crushers in

comminution circuits and some of the factors that are critical to optimising the efficiency and how the

circuits are impacted are examined in the following sections.

OPTIMISING THE EFFICIENCY OF CRUSHING The cone crusher relies on delivering crushing energy into the material contained within the crushing

chamber, or cavity. Many factors influence the effectiveness of the crushing in the chamber and the

amount of size reduction achieved. The factors can be split into:

Mechanical Design Variables (MDV),

Feed Material Variables (FMV),

Machine Operating variables (MOV),

Machine limits,

o Volumetric (maximum feed size, capacity),

o Power,

o Force, and

Flowsheet design and interaction with other equipment.

Mechanical Design Variables Mechanical Design Variables (MDV) are those variables that are specified during the design of cone

crushers. The main variables are listed below:



Cone head pivot point,

Cone head angle,

Eccentric throw,

Eccentric speed , and

Chamber and liner design.

In cone crushing the MDV significantly affects energy application, crushing efficiency, reduction ratio,

product size/shape and liner wear life. The combinations making up the MDV are usually fixed for a

particular model of cone crusher (with the exception of the chamber/liner design), although some

manufacturers do offer limited variability for a set base design.

The most significant effect of varying combinations of MDV is the variation in crushing force achieved.

In terms of the MDV parameters listed above, they are somewhat machine specific. The base design

of cone crusher still falls into the two broad categories, namely:

Top supported – consisting of a spider arrangement holding the top of the mainshaft. These

tend to have a steeper crushing chamber, a smaller eccentric throw and run at slower

eccentric speeds.

Base supported – crushing head sits on an eccentric assembly that rotates around a

stationary mainshaft. These tend to have a flatter chamber, a larger eccentric throw and a

higher eccentric speed.

As a result of the market being split between two competing designs, the advantages and

disadvantages of the two approaches and the variables used are the subject of great discussion. Due

to this sensitivity it is difficult to be definitive, but general points can be made as follows.

It is generally agreed that the liners on a steep cone head will have more uniform wear profiles than

those on shallow cone heads. Very shallow cone head angles, as found on some fine crushing

cones, are prone to ‘dishing’ of the manganese liners. Dishing is generally accepted as being a

region of accelerated wear caused by numerous factors, the most common of which is localised

crushing due to a predominance of certain particles sizes in the feed. Another cause can also be

“wash-out” due to the flow of material through the chamber.

The position of the pivot point in association with the eccentric angle and the eccentric throw has a

significant impact on the crushing force and power. In several of the newer cone designs, the pivot

point has been raised to improve process performance. The reason is to increase the relative throw in

the upper part of the crushing chamber.



The eccentric throw is, in process terms, the difference between the open side setting (OSS) and

closed side setting (CSS). In mechanical terms it is the eccentricity of the bush. As the eccentric bush

usually sits quite close to the discharge point of the crusher, the approximation that the throw is OSS

– CSS is mostly adequate for equipment descriptions, but for detailed analysis the eccentricity at the

bush is the most important. As throw controls the difference between open and closed points in the

chamber, it therefore controls the reduction ratio of the chamber. The throw also affects the dynamics

of the interaction between the crusher and the rock material. The general comment regarding

eccentric throw is that it has a directly proportional impact on throughput and increased throw –

matched to changes in eccentric speed – will provide a finer product. Although in most cases a

change to eccentric requires a major maintenance shutdown to replace the eccentric bush, in some

machines it is possible to make small adjustments to throw with minimal intervention, i.e. Metso

GP300-S.

It has been found that the effects of eccentric speed on performance can vary depending on the style

of crusher and the duty. Increased speed will provide more hits and thus more opportunities for size

reduction. The speed will also affect the so called effective throw which in turn is used for

compression and thus the breakage. For modern crushers increased speed decreases the effective

throw and thus also the throughput (capacity). Fig. 1 shows the variation in product size distribution

for a Hydrocone style crusher in a fine crushing application for changing eccentric speeds.

Fig. 1: Change in product size distribution with eccentric speed for a Hydrocone crusher (Hulthén and

Evertsson, 2008)

From most texts it appears there is a throughput increase with speed up to a limiting value, above

which point throughput starts to drop. Hulthén, (2010), has extensively examined the use of eccentric

speed as a joint control parameter alongside the closed side setting. Part of the reasoning behind the

use of speed control is to try and compensate for changes in crusher size reduction with liner wear.

Fig. 2 illustrates the type of performance decay that is normally seen, without any advanced control

able to compensate for the wear of the manganese liners.



Fig. 2: Loss of crusher performance with liner wear and CSS drift.

Hulthén’s (2010) work shows that there is value in such speed control. Fig. 3 shows the “eYe”

algorithm where the speed control acts to maintain a required throughput. The algorithm is based on

true measured performance of an HP-style crusher. The use of such control is claimed to offer

extended performance life for liners and the ability to maintain higher performance levels compared to

manual operation. The appearance of the performance function clearly shows that a fixed speed is

sub-optimal with respect to liner life.

Fig. 3: “eYe” speed control algorithm for cone crushers.

The benefits of such control can be seen in Fig. 4, where the traditional decline in crusher

performance with liner wear is alleviated by the use of the speed and CSS control. Quantitatively, the

impact of the control exhibited in Fig. 4, where the liner life is increased by 27% whilst maintaining

crusher performance (Hulthén, 2010). The increased liner life may be explained by a more uniform

distribution of the work hardening of the manganese surfaces of the crushing chamber.



Fig. 4: Comparison of throughput over mantle liner life using speed and CSS control compared to using

a set eccentric speed.

The design of crushing chambers is absolutely critical to cone crusher performance as it defines:

Dynamics governing the flow of the rock material,

Energy distribution within the chamber and therefore the process performance and the

efficiency of the energy input into the particles in the chamber,

Minimum plan cross sectional area, i.e. the control point for throughput, and

Breakage modes.

In the design of the cavity, the distribution of force which produces the crushing also leads to the

development of liner wear. It is therefore important to have a cavity that not only performs when the

liners are new, but maintains its performance over the liner life. The development of poor liner

profiles is not only considered one of the main causes of degraded process performance, but it is also

a major factor in mechanical failures of crushers. A typical scenario is poor wear profiles generating

high crushing forces, which in turn cause ring-bounce ore packing. If ring-bounce is not addressed, it

leads to damage and wear of the bowl seats, which is usually a pre-cursor to major damage to the

mainframe of the machine. A range of other possible premature fatigue failures may also occur.

In terms of process performance, the liner profile must induce a good flow of material via a suitable

nip angle between the concave and mantle liners. The correct design of nip angle is essential.

Generally nip angles of 19° or less tend to induce high crushing forces and pressures. With nip

angles above 25° (depending on the rock to manganese coefficient of friction) the tendency is to eject

material rather than capture/fracture.



Machine Operating Variables The two main Machine Operating Variables (MOV) are closed side setting and feed presentation.

The closed side setting of a cone crusher is the most well-known adjustable mechanical parameter.

The CSS is located at the discharge point of the crushing cavity and it is the minimum distance

between the mantle and concave liners. The CSS can be adjusted either hydraulically, or

mechanically (screw thread), dependent on the style of crusher. The means of adjustment has no

material process impact, except in terms of the frequency of the adjustment procedure, which in

instances of high wear applications can be significant. Loss in production due to the requirement to

stop the crusher and adjust is illustrated in Fig. 2.

Closed side setting can be adjusted to control product size, energy consumption and throughput.

Bearman et al., (1991) produced relationships between the CSS and the fracture toughness of the

feed and both product size and energy consumption, for a common feed size (Fig. 5).

(a) (b)

Fig. 5: Relation between closed side setting, fracture toughness (Kcb) and (a) energy consumption and (b)

eighty percent passing size of product (P80).

Feed presentation is often neglected, with the only factor generally discussed being the choke feed

condition. Choke feeding, whereby a head of material is maintained in the chamber, is a desirable

feature as it aids the stability of operation, but there are many other factors encapsulated in feed

presentation that are equally important. Some of the lack of focus on feed presentation stems from

the older attitude to crushers, whereby they were regarded as brute force devices that would take

practically anything that could be thrown at them and still produce a reasonable product. Such an

attitude is manifested in situations such as:

Crushers fed directly from screens,

No feed rate control,

No surge capacity ahead of crushers,



Segregation of particle size in the feed, and

Poor distribution of material around the crushing chamber.

Crushers from the mid-20th century managed to cope with some sub-optimal conditions due to their

heavy mechanical construction and the lack of expectation around their size reduction performance.

In installation manuals from the 1950’s and 1960’s it is clearly stated and emphasised that the

presentation of the feed should be very carefully addressed.

Developments in crushing technology have led to a new generation of cone crushers which are rated

to much higher throughputs and capable of delivering significantly higher power and hence finer

product. The downside to the newer cone crushers is that they are less tolerant of poor feed

conditions and such sub-optimal conditions will not just lead to reduced process performance, but

mechanical integrity can also be compromised. Over the years the tonnages in process plants and

concentrators have increased dramatically, which in turn have resulted in increased dimension of the

equipment whilst the particle sizes have remained roughly the same. For example conveyor belts are

wider today compared to the past. This leads to increased problems with segregation and proper

handling of bulk materials.

To achieve optimal performance, feed rate must be controlled, ideally via bins and feeders that deliver

a consistent and controllable feed rate. Such arrangements ensure that crushers are not run empty

and the mantles do not free-spin. Crushers are designed to have a consistent force acting through a

certain position in the machine to keep defined loads on bushes. Poor feed control does not allow

such conditions to exist and the subsequent reduction in mechanical availability can be highly

detrimental.

Feed distribution is again known to impact crusher performance. The main problems relate to the

quantity of particles reporting to specific areas of the chamber, or alternatively segregation in feed

size that sees finer material in one area of the crushing chamber compared to other areas. Mostly

such issues have been regarded as a part of normal operation and, although not ideal, not hugely

detrimental to crusher performance.

Evertsson (2010) has shown that segregation and poor distribution are not simply a minor problem,

but are a cause of high crushing forces and inefficiency in crushing.

Fig. 6 shows a photograph of feed to a cone crusher. On initial inspection the feed appears to be well

centred and distributed. It appears to be fully choke fed and as such it is expected that this crusher

will perform well.

Traditionally data sampling of crusher instrumentation is not undertaken at high rates. In this instance

a high sampling rate was used to investigate the performance of the machine and this identified an

issue that was not obvious from visual inspection.



Fig. 7 shows high rate sampling of the crusher running at the conditions shown in Fig. 6 and a similar

pressure plot for the same crusher once segregation was eradicated. It can be seen the pressure

signal after the segregation had been removed shows a much less erratic variability and the overall

magnitude of the pressure is much lower. The effect on the crusher is that there will be fewer

hydraulic overloads and the CSS will be maintained over much longer periods of time. From a

process view this means that the product size distribution is much more consistent and the amount of

recirculating load is reduced. The crusher can thus operate at a smaller CSS which results in a finer

product. From a mechanical standpoint the incidence of mechanical failure will be reduced due to the

lower pressure being experienced by the crusher.

Fig. 6: Photograph of crusher feed hopper, with size segregation shown by the smaller particles

migrating to the back of the hopper (Evertsson pers. comm., 2011)

Fig. 7: Pressure-time plots for crusher with and without, particle size segregation (Evertsson pers.

comm., 2011)



Feed Material Variables All rock materials, ores and minerals are associated with natural variations in properties. The Feed

Material Variables (FMV) of importance are:

Feed material strength,

Bulk density,

Feed size distribution and maximum feed size, and

Feed moisture content.

The measure of rock strength most applicable to crushing has been the subject of considerable work

over the years. Many measures are used with varying degrees of success. The main measures used

include:

Bond Crushing Work Index,

Fracture Toughness,

Uniaxial Compressive Strength,

Tensile Strength, and

Drop Weight Test.

Regardless of the test used from the above list, each provide estimates of the strength and therefore,

using appropriate relationships, can be used to predict power consumption and product size

distributions. Depending of the level of information retrieved from a test procedure, different levels of

accuracy in the predictions can be achieved.

Cone crushers are capable of accepting a variety of feed sizes depending upon the combination of

mantle and concave. Most manufacturers supply a range of mantles and concaves from extra-fine to

extra-coarse for any given diameter of cone head. Generally it is accepted that the feed size should

not exceed 80% of the maximum open side feed opening. The definition of feed size is however open

to interpretation. Cone crushers will accept slabby or elongated feed, although these reduce capacity.

A mismatch of feed size and crusher manganese configuration is also potentially detrimental to

crusher performance. For example, partial bridging in the feed material just above the inlet to the

crushing camber may occur which reduces capacity, but this cannot be determined by visual

examination.

Manufacturers of cone crushers have recommended that material equivalent to or smaller than the

closed side setting is removed from the feed. Over the years this suggestion has been modified,

particularly where product shape is important, i.e. in the aggregates industry. The fines increase the

bulk density of the feed and thus more interparticle breakage and higher compaction pressures are



achieved. Fines in the feed still have a tendency to cause the crusher to draw high power and in the

tertiary and quaternary stages of crushing the fines will often lead to packing and “ring-bounce”. Once

“ring-bounce” occurs the crusher has hit its force limit, process performance deteriorates and the

incidence of mechanical damage increases.

Manufacturers often impose moisture content limits of 4%. As a general rule the impact of the

moisture content is dependent on the size distribution, fines content and mineralogical composition of

the feed. In many instances crusher performance will deteriorate with moisture content, but at a

certain point free moisture can aid the crushing process. Other problems associated with moisture

content can be accelerated liner wear and high crushing forces, where the moisture enhances

packing.

Machine limits In terms of the mechanical and geometric limits for cone crushers, they can be divided into:

Volumetric Limit,

Purely the maximum rate at which material can flow through the crusher,

Maximum feed size able to enter the inlet of the crushing chamber,

Power Limit,

Point at which the crusher draws the maximum motor power, and

Force Limit,

Point at which the force limit is exceeded and the upperframe bounces, or the

hydraulic system relieves.

Of these limits the most obvious and easy to understand is the volumetric limit. Any crusher has a

maximum chamber, or cavity volume, that will pass feed through the crusher at a rate determined by

a raft of factors, but essentially controlled by the minimum cross sectional area. Once this volumetric

rate is exceeded material will overflow the crusher. The limit is not normally seen in practice, as

monitoring and instrumentation can detect such build-up and reduce the feed rate. It should be noted

that the wear of the mantle and concave liners can have a major impact on the volumetric limit and

hence, particularly with shallow head angle machines, there can be quite dramatic changes in the

volumetric limit of a crusher over the total wear life of the liner set, see Fig. 4

The power limit refers to the point at which the power required for crushing the material in the

chamber reaches and exceeds the installed motor power. Such a situation is mainly seen in primary

and secondary crushers, where the reduction ratio is high. Power limits can be encountered

prematurely due to inappropriate crushing chamber selection or due to the development of poor wear

profiles that cause the crushing to concentrate feed into a certain section. The other main cause of



high power is the crushing of extremely hard feed material. It is known that power draw increases

significantly with increasing rock strength, see Fig. 5.

The least understood limit for crushers is the crushing force limit. This is often confused with the

power limit, but they are clearly different and manifest in different ways. The force limit is reached

when the vertical component of the force generated in the chamber exceeds the force holding the

mantle and concaves in the crushing position. In bowl adjusted crushers there is a clamping force

exerted by either hydraulics or springs, whilst in head adjusted machines a hydraulic system holds the

head in the crushing position. All “clamping” devices, whether hydraulic or mechanical, have a set

point or limit. Once exceeded, the crusher will relieve. In the case of hydraulic machines, such relief

will be in the form of the hydraulics engaging to separate the bowl and mantle. With a spring based

machine the springs will simply relieve enough to reduce the crushing force, but if the condition

continues it becomes known as ring-bounce. Ring-bounce is where the machine is constantly at, or

above, its crushing force limit and the spring or hydraulic relief mechanism is continually operating.

Ring-bounce is a problem as the process performance is degraded, but it can also cause extensive

mechanical damage.

A mathematical framework coupling all of the above mechanical and feed parameters with the

traditional metallurgical measures of throughput and product size offer a powerful method for

analysing and optimising crusher performance. Such approaches are examined in the following

section.

CRUSHING CIRCUIT ANALYSIS The current state of public domain crusher modelling is typified by the Whiten population balance

approach (Whiten, 1974) that is found in many commercially available software platforms. The model

applies a set of selection and breakage functions to a feed size distribution, predicting the complete

product size range and power consumed by the crushing action. The Whiten model’s chief limitation is

the use of empirical parameters and functions to define the crusher machine’s performance (e.g. K1,

K2, K3 and t10). There is a subsequent reliance on extensive sample data to characterise a given

crusher.

Advances in crusher modelling have been ongoing for several decades (Bearman, (1991), Briggs,

(1997), Evertsson, (2000)). These examples are typified by a particle dynamics approach, where the

motion of particles freefalling or sliding down surfaces in the crusher chamber are explicitly calculated.

Particle nipping is detected based on crusher liner geometry and particle sizes, and material-specific

breakage functions are applied. Liner wear rates are approximated from the internally developed

crushing forces (Lindqvist, 2003).

This approach requires the complete specification of mantle and bowl liner profiles, machine

eccentricity, head rotation speed and feed material properties. This eliminates the empiricity inherent

in the Whiten model.



The advantages of machine and material specification include accurate prediction of product size,

power draw, throughput limits and wear rates for any chosen liner set, machine operating parameters

or combination of feed materials. The test data required for the model relates to the feed material

strength properties, as defined by the Drop Weight Test, and the feed particle size distribution to the

crusher. The framework is applicable to both gyratory and cone crusher machine configurations.

Recent efforts have focused on expanding this advanced crusher modelling approach from an

individual unit operation to the flowsheet context. The authors have furthermore introduced the time

domain to the flowsheet approach, with the advanced crusher model being successfully integrated

into the SysCAD dynamic simulation platform.

The SysCAD dynamic simulation platform offers the ideal method to analyse the energy and

processing requirements of a crushing circuit, particularly related to changing feed conditions, liner

wear, process control and operating philosophies.

Several case studies exploring the use of the dynamic simulation are presented below. The same four

stage crushing circuit feeding a milling operation was analysed in each.

Changing feed conditions The properties of the Run-Of-Mine (ROM) feed to a comminution circuit are ever-changing, often

exhibiting both continuous and discrete step changes as mining encounters ore/waste boundaries,

geological discontinuities and operationally induced variation, i.e. blast performance.

These changes have an immediate effect on a processing plant’s throughput and product quality

performance. Knowledge of how changing feed may affect a comminution circuit is vital to developing

a strategy for mitigating the risks to production and optimising energy input.

In Fig. 8 (a), the forecast composition and fragmentation of ROM material from an open pit mine

varies over a two year period. The ROM material is fed to a crushing plant prior to milling. Many of the

plant’s operating parameters are fixed, having been identified over time as the ‘best’ way to run.



(a)

(b)

Fig. 8: (a) Trend of hard/soft fraction and ROM P80; (b) Circuit energy consumption arising from alternative liner sets.

The advanced crusher model, applied in the dynamic time domain, offers a useful method to

benchmark and visualise plant performance under the changing feed conditions.

Fig. 8 (b) forecasts how the crushing plant product size and energy consumption will vary with the

changing ROM when the existing Coarse liner set is maintained.

Consideration of the crushing mechanisms in action during the finer feed periods does, however, offer

the opportunity of improving plant performance beyond expectations. The main changes that can be

introduced with fine feed include the use of finer crushing chambers and the optimisation of CSS

across the circuit. Finer feed tends to allow more stable operation of the crushers and more effective

crushing. The move away from coarse feed also changes the type of crushing limit from a power limit

to a force limit.

Simulating such mechanical changes to the crushers, in the same dynamic flowsheet context as

above, suggests the potential improvement summarised in Fig. 8 (b). The fine liner sets and CSS

modifications are able to achieve identical size reduction for greatly reduced energy consumption, but

only in the instances where the motor power and mechanical force limits of the machines can be

satisfied. The period between May 2012 and June 2013 offers a clear opportunity to reduce energy

use and ultimately cost, via a simple, proactive change in operating philosophy.



Whilst such opportunities are often obvious to plant metallurgists, the framework presented provides

an analytical tool for quantifying the benefits of changes to crusher operating parameters, in a ‘virtual’,

risk-free environment.

Liner wear The advanced crusher modelling approach described previously includes a liner wear module. The

module uses the crushing forces developed upon the mantle and bowl during a single head revolution

to estimate the rate of mass removal from each surface (Lindqvist, 2003). The liner geometry is

altered along the length of each surface depending on the forces applied at each point. The amount of

crushing performed on feed of any given particle size distribution and material hardness, by any given

liner geometry, is used to explicitly determine wear for the specific duty and application.

Furthermore, the altered liner geometry resulting from a stage of wear will affect how crushing forces

are subsequently developed and distributed. Many small, discrete stages of wear are therefore used

to estimate the continuous wear experienced by real liners.

The wear-induced re-distribution of crushing forces also shifts the distribution of power consumption

from the crushing actions. Fig. 9 demonstrates how the redistribution of power from wear can result in

a coarser product size for similar total energy consumption. This loss of reduction ratio must clearly

be compensated for further downstream, perhaps by subsequent crushing stages or other

comminution operations. Total circuit energy consumption will increase because of the reduced

efficiency of the worn liners.

Fig.9: Simulated mantle and bowl liner profiles, chamber power distribution and product sizes for new

and worn liner sets. Total power is similar in each case (257 kW initial and 242 kW worn).



To illustrate the circuit-wide consequences of liner wear, the four-stage crushing plant was

dynamically simulated using the advanced crusher modelling framework. The simulated plant feed

properties were held constant to isolate wear effects from transient feed effects. Fig. 10 demonstrates

how the burden of crushing is shifted towards the tertiary crushers as the secondary crusher liners

wear. The recirculating load feeding the quaternary crushers also increases, non-linearly, as the

tertiary crushers wear from the increased feed rate. The crushing circuit product P80 increases in

parallel.

Fig.10: (a) Change in feed rate to tertiary and quaternary crushing stages due to liner wear; (b) Crushing

circuit product P80 resulting from liner wear.

Having isolated and observed these effects, it is possible to devise a strategy for improving energy

consumption. Fig. 11 shows the change in overall circuit energy consumption where the crusher liner

sets are replaced at more frequent intervals identified from the simulation. The improved crushing

energy utilisation means a finer average product is transferred to downstream milling, resulting in

lower mill energy consumption, i.e. a cascading benefit.



Fig. 11: Decreased energy consumption arising from increased liner replacement frequency.

Product size and near-size Every crushing circuit will have a required product size, which may be a final size as in most

aggregates operations, or alternatively a transfer size to some subsequent comminution process.

The delivery of a final product size relies on the interaction of crushers and a classification device,

mainly vibrating screens. It is critical to examine the likely performance of the crushers in the

flowsheet compared to the desired final separation size. The main reason for this analysis is to

ensure that there is not a bottleneck caused by the size reduction capabilities of the crushers and the

separation performance of the screens. Such a bottleneck is often referred to as a “near-size”

condition.

The near-size condition exists where the size distribution generated from the crushers produces a

great deal of material close to the separation size of the screens.

Fig. 12 (a) uniquely highlights how near-size particles are contributing to a crushing circuit’s

recirculating load issues, manifested as a ‘spike’ between the crusher CSS and the screen aperture

size.



Fig. 12 (a): Recirculating load versus size composition, current state, showing the ‘Targeted’ CSS of

8mm, the ‘Effective CSS’ of 16mm and associated screen aperture of 17mm.

Fig. 12 (b): Recirculating load versus size composition, potential state, showing a ‘Sustainable’ CSS at

12mm and associated screen aperture of 17mm.



Fig. 12 (c): The decrease in circuit energy consumption delivered by the conditions behind Fig. 12(b)

compared to Fig. 12(a).

The crushing circuit in question is characterised by both difficulty and poor discipline in measuring and

achieving closed side settings, feed presentation issues manifesting as ring-bounce, and wear profiles

continually shifting the perceived, or ‘effective’, CSS upwards. A total recirculating load in excess of

200% is regularly experienced, frequently resulting in plant feed reductions as screens and bins

struggle to manage the load.

The advanced crusher modelling framework, applied in a flowsheet context, becomes a powerful tool

for exploring options to mitigate the near-size phenomenon, stabilising throughout and improving

energy consumption.

Fig.12 (b) shows a more optimal situation, where feed presentation has been improved, and a more

appropriate liner set and sustainable CSS have been selected to remove the spike and decrease the

overall recirculating load. Fig. 12 (b) and (c) demonstrate that whilst the recirculating load has

decreased significantly, crushing energy consumption has increased. A finer product is transferred to

milling where the energy required to maintain feed to subsequent processes has decreased

dramatically. Crushing is, in this case, playing a key role in energy efficient comminution.

CONCLUSIONS Understanding the operation and design features of crushers can significantly improve the energy

efficiency of both the individual machines, but more importantly the overall circuit. Many of the

improvements that can be made to the operation of crushers are straightforward and involve minor

costs in comparison to the overall benefit. As has been illustrated, changes involving the following

can deliver major improvements in energy efficiency:

Feed presentation,

Matching of feed size distribution to crusher cavity design,



Optimised liner change-out philosophy, and

Reduction of “near-size” material in closed circuits.

In combination with appropriate control of crushing and screening operations, an overall operational

philosophy can be developed that can be trialled using dynamic simulation techniques. Such an

approach will ensure that the full energy benefit of crushers is delivered and that their use in

association with other energy efficient processes will provide the best overall outcome.

REFERENCES Bearman R A, 1991. The application of rock mechanics parameters to the prediction of crusher

performance, PhD thesis (unpublished), Camborne School of Mines, UK.

Bearman R A, Barley R W and Hitchcock A, 1991. Prediction of power consumption and product size

in cone crushing, Minerals Engineering, 4(12):1243-1256.

Briggs C A, 1997. Fundamental model of a cone crusher, PhD thesis (unpublished), University of

Queensland, Australia.

Evertsson C M, 2000. Cone crusher performance, PhD thesis (unpublished), Chalmers University of

Technology, Gothenburg, Sweden.

Evertsson C M, 2011. Personal Communication.

Hulthén E and Evertsson C M, 2008. On-line optimization of crushing stage using speed regulation

on cone crushers, Proc. of XXIV International Mineral Processing Congress, 2 pp.2396-2402, Beijing,

China.

Hulthén E 2010. Real-time optimization of cone crushers, PhD. thesis (unpublished), Dept. of Product

and Production Development, Chalmers University of Technology, Gothenburg, Sweden.

Lindqvist M, 2003. Prediction of worn geometry in cone crushers, Minerals Engineering, 16(12):1355-

1361

Whiten W J, 1974. A matrix theory of comminution machines, Chemical Engineering Science, 29:31-

39.



Reducing the Energy Required in Grinding Clinker to Cement – Some Case Studies

H Benzer1, N Aydogan2, H Dundar3 and A J Lynch4

1. Professor, Hacettepe University, Department of Mining Engineering, Beytepe, Ankara 06800, Turkey. Email: [email protected]

2. Lecturer, Hacettepe University, Department of Mining Engineering, Beytepe, Ankara 06800, Turkey. Email: [email protected]

3. PhD Student, Hacettepe University, Department of Mining Engineering, Beytepe, Ankara 06800, Turkey. Email: [email protected]

4. HonFAusIMM, Mineral Engineer, 9 Adelaide Street, West End SA 4101. Email: [email protected]

ABSTRACT Grinding clinker to finished cement consumes more energy than any other comminution process.

World production of cement in 2009 was 2840 million tonnes and the energy required for

comminution was 40-60 kWh/tonne depending on the clinker characteristics and the fineness of the

finished cement. Machines used include single and multi compartment ball mills, HPGRs, vertical

roller mills, centrifugal air separators, V separators, screens and crushers. The circuit most commonly

used consists of a closed circuit HPGR followed by a single compartment ball mill and an air

separator, and in operation the variables are adjusted to obtain maximum throughput with no loss of

quality of the separator fines.

A common objective in cement plants is to reduce energy consumption without affecting the particle

size in the finished cement and simulation is used as one procedure to achieve this. Models are now

available for all the machines commonly used in clinker circuits and circuit optimisation studies using

simulation have been carried out in many plants leading to decrease in energy consumption per tonne

of cement. In this paper the models used will be discussed briefly and several case studies will be

reviewed which will highlight the effect of change in ball size in a two compartment mill, the role of

the HPGR in an HPGR – ball mill – air separator circuit, and the use of a high intensity vibrating mill

for producing finished cement with high additive ratio.

INTRODUCTION Cement production is an energy intensive process. It consumes 2% of the global primary energy and

5% of the total global industrial energy (World Energy Counc., 1995). Grinding is a high-cost

operation consuming approximately 60% of the total electrical energy expenditure in a typical cement

plant. The electrical energy consumed in the conventional cement manufacturing process is in the

order of 110 kWh/tonne, of which approximately 30% is used for raw materials preparation and 40%



for the finish milling (cement clinker grinding). This is the largest single consumption point of electric

power in the process of converting raw materials to finished cement (Bhatty et al., 2004).

The exact cost of energy is significant as a part of the total production costs in the cement industry

and better energy efficiency will reduce costs. Substantial potential for energy efficient improvements

exist in the cement industry and in individual plants. The cement industry is targeting technologies

focusing on increasing the mill throughput; energy savings; and minimizing the production costs

without negatively affecting product yield or quality.

The increasing demand for ‘‘finer cement” products, and the need for reduction in energy

consumption and green house gas emissions, necessitate the optimization of grinding circuits.

Opportunities exist at cement plants to improve energy efficiency while maintaining or enhancing

productivity both in raw and finish milling. This paper is concerned with using modelling and

optimisation to improve energy efficiency, and application of high pressure grinding rolls and

vibrating ball mills for cost effective operations.

CEMENT GRINDING MODELLING

Ball mill and separator models Conventional cement grinding ball mills, operated in closed circuit with air separators, have usually

two grinding chambers, which are separated from each other by a slotted diaphragm through which

the particles finer than the size of the slots pass to the second chamber for further size reduction.

Cement grinding mills are modelled by using the modelling approach described for closed cement

grinding circuits (Lynch et al., 2000; Benzer et al., 2001, 2003). According to this approach, each

grinding compartment is modelled by the perfect mixing model (Whiten, 1974) which considers a ball

mill or a section of it as a perfectly stirred tank. Then the process can be described in terms of

transport through the mill and breakage within the mill. Because the mill or section of it is perfectly

mixed a discharge rate, di, for each size fraction is an important variable in defining the product. The

parameter si indicates the mill content.

iii sdp (1)

The model for steady state operations includes two sets of model parameters, i.e. the breakage

function aij and a combined breakage/discharge rate ri/di function.

i

ji

i

ii

jjiji p

d

rp

dj

rpaf

1

0 (2)

The breakage function is a material characteristic and defines the size distribution of the product

formed after the breakage of the parent size fraction. It is determined by laboratory tests using twin

pendulum (Narayanan, 1985), ultra-fast load cell (Hoffler and Herbst, 1990) or drop weight apparatus



which has been practiced by several researchers for various material characterization purposes (Gross,

1938; Piret, 1953; Fairs, 1954; Schonert, 1972; Rumpf, 1973; Pauw and Mare, 1988). In this study the

breakage distributions were determined by drop weight tests. The combined breakage rate/discharge

rate function defines the machine characteristics and can be calculated when feed and product size

distribution are known and breakage function is available. For most ball mills r/d on particle size is a

smooth curve which can be fitted to a spline function with four knots. The ball mill model is

calibrated by determining the r/d values using the feed and product size distributions obtained under

particular operating conditions. Where the size distribution of the mill contents is available, breakage

rates and discharge rates can be calculated separately (Napier-Munn et al., 1996). The important

variable for each mill is r which is the tonnes broken per hour/tonne in the mill. This can be calculated

from d which is the tonnes discharged per hour/tonne in the mill and r/d which is calculated from

plant data. Schematical representation of the modelling approach (Benzer, 2000) used for two-

compartment ball mills is given in Fig. 1. This modelling approach requires the size distribution data

obtained from inside the mill.

Fig. 1 - Ball mill model structure (Benzer, 2000)

In the above model structure, Stream-1 represents the mill feed calculated by mass balancing around

the circuit. Stream-2 size distribution depends on the size reduction achieved in the mill usually

presented by the size distribution of the sample collected at the last 1 m of the first chamber length.

Mill powder rejected from the diaphragm after screening is represented by Stream-5 which is

predicted by simulating the diaphragm with its efficiency curve model parameters defined by Whiten

(1966). In the modelling of the screening effect of the diaphragm, efficiency curve parameters; C, by-

pass value, , sharpness of the screening, , fish hook effect, d50, cut-size are back-calculated. The

efficiency curve model is capable of defining the fish hook and is used for separators. The general

form of the equation is presented below:

E oa

2)exp()..exp(

)1))(exp(..1(*

*

X

XC (3)

where C indicates the by-pass of the separator, sharpness of the separation, fish hook behaviour of

the classification and X=di/d50c.



CASE STUDIES

Effect of change in ball size in finish grinding In this study, the capacity improvement and specific energy reduction have been achieved in the finish

milling stage of a cement plant in Izmir, Turkey by adjusting the ball size using modelling and

simulation tools. An extensive sampling campaign has been conducted around the circuit for

modelling and simulation purposes. The circuit consists of closed circuit ball milling with an air

classifier. Prior to the sampling campaign, steady state condition has been verified from the control

room data. After completing circuit sampling, the ball mill was “crash–stopped” for sampling inside

the mill. The simplified flowsheet of the circuit and sampling points are given in Fig. 2. During the

sampling campaign, four sets of data were collected for different types of cement relating to the

additive type and ratio. The details of the sampling are given in Table 1.

1- Clinker2- Pozzolana3- Limestone4- Gypsum5- Fly ash6- Mill overflow7- Filter return8- Mill discharge9- Separator feed10- Separator reject11- Product12- Mill inside samples

9

7

8

4

612

2 31

5

11

10

Two-chamber ball mill

Air classifier

Filter

Fig. 2 - Simplified flowsheet of the circuit and sampling points

Table 1 - Sampling conditions

Sampling #1 Sampling #2 Sampling #3 Sampling #4

Product Type A B C D Clinker (tph) 141.1 112.9 128.4 119.6

Pozzolana (tph) 0.0 39.6 10.5 0.0

Limestone (tph) 4.6 8.3 0.0 4.3

Gypsum (tph) 9.4 7.4 6.8 8.0

Fly ash (tph) 0.0 10.0 16.1 10.0

Total 155.1 178.2 161.8 141.9

Product fineness, passing% 32µm

83.8 83.7 88.4 90.8

The specific name of the products for sampling surveys is given in Table 2. Regarding the type of the

additive there are five main classes of cement denoted as CEM from I to V. (CEM I: Portland cement,

CEM II: Portland-composite cement, CEM III: blast furnace cement, CEM IV: pozzolanic cement,

CEM V: composite cement). The sub-classes are defined with the ratio of the additives and required



strength value (EN 197-1:2000). For instance, CEM II is the Portland-composite cement that can

comprise several additives. A and B defines the range of the clinker ratio i.e. A means higher clinker

ratio, B means lower. M indicates more than one additive is used and the symbols for the additives are

given in between parentheses; P for pozzolan, W for fly ash. The number at the end represents the 28-

days strength in MPa. N and R represent normal early strength and higher early strength respectively.

The cementitious properties of the additives are lower than that of cement clinker so introducing more

additive requires finer grinding to achieve the same quality cement in terms of the strength. As the

surface area of the product increases the strength increases.

Table 2 - Specific name of the products refer to EN 197-1:2000

Product Specific name of the product

A B C D

CEM I 42.5R CEM II/B-M (P-W) 32.5R CEM II/A-M (P-W) 42.5R CEM II A-(W) 52.5N

The samples collected from inside the mill and around the circuit including the fresh feed were

analysed by dry sieving connected with laser diffractometer sizing technique for the particle size

determination. The particle size distributions of the samples were determined starting from the top

size of the material down to 2 µm. By using the particle size distributions and the control room data,

comprehensive mass balance studies were carried out around the circuit. Mass balance calculations

were performed to determine adjusted size distribution values and tonnages required in the further

modelling studies. A summary of calculated data is given in Table 3.

Table 3 - Mass balances for four sampling surveys

tph Sampling pt. 1+2+3+4 5 6 7 8 9 10 11

Sampling #1 155.1 0.0 289.1 65.2 354.3 354.3 199.2 155.1

Sampling #2 168.2 10.0 251.3 81.0 332.3 342.3 164.1 178.2

Sampling #3 145.7 16.1 310.2 55.4 365.6 381.7 219.9 161.8

Sampling #4 131.9 10.0 293.5 51.3 344.8 354.8 212.9 141.9

The two-compartment ball mill modelling approach was used for the ball mill while the Whiten

efficiency curve model was used for air-classifier and filter at the model fitting stage. Modelling

studies were performed separately for each equipment in the circuit using JKSimMet software. Then,

the simulation studies were performed for the purpose of optimization, Fig. 3 gives the fitted size

distribution data with the mass balanced data for the key streams around the circuit. As can be seen

from these graphs of the data the modelling results fit the mass balanced data very well. This indicates



the capability of the model. The fitted tonnage values around the circuit are given in Fig. 4 along with

those from the mass balances.

0

10

20

30

40

50

60

70

80

90

100

0.001 0.01 0.1 1 10 100

% p

assi

ng

size (mm)

Sampling #1

mill feed

2.ch 0.m

sep. feed

product

f it

0

10

20

30

40

50

60

70

80

90

100

0.001 0.01 0.1 1 10 100

% p

assi

ng

size (mm)

Sampling #2

mill feed

2.ch 0.m

sep. feed

product

f it

0

10

20

30

40

50

60

70

80

90

100

0.001 0.01 0.1 1 10 100

% p

assi

ng

size (mm)

Sampling #3

mill feed

2.ch 0.m

sep. feed

product

f it

0

10

20

30

40

50

60

70

80

90

100

0.001 0.01 0.1 1 10 100

% p

assi

ng

size (mm)

Sampling #4

mill feed

2.ch 0.m

sep. feed

product

f it

Fig. 3 - Size distributions around the circuit after model fit

0

100

200

300

400

0 100 200 300 400

fitted TPH

mass balanced TPH

Sampling #1

Sampling #2

Sampling #3

Sampling #4

Fig. 4 - Fitted tonnage values vs. mass balanced tonnage values

During the simulation studies capacity improvement was the main concern for the same product

quality in terms of the fineness. For this purpose ball size optimization was considered. By detailed

analysis of the sizing data from inside the mill for both the first and second chambers, a new ball size

distribution was determined for each chamber. Table 4 gives the new ball charge with the existing

one. More than a 100 sets of industrial scale data coming from various cement plants reviewed to date

were compared and the result used to optimize the ball charge. The data base consists of survey data

acquired from ball mills which are different in diameter, length, ball charge composition and

processed cement type. Modelling those ball mills, using Benzer’s (2000) approach, results in



different breakage rates which are a function of various design and operational variables. A statistical

analysis of the data set led to identifying the relationships between the design or operational variables

and breakage rate. By using these relations new breakage rates were calculated and simulation studies

were performed. Simulation results showed that the capacity of the circuit can be improved up to 10-

17% for different product types with the new ball size distribution.

Table 4 - Ball size composition for the 1st and the 2nd chambers

weight %

Existing charge New charge

Ball size (mm) 1st chamber 2nd chamber 1st chamber 2nd chamber

90 19.2 - 13.6 -

80 34.6 - 28.8 -

70 21.1 - 23.7 -

60 25.1 - 22.0 -

50 - 7.9 11.9 -

40 - 10.0 - 8.3

30 - 20.0 - 12.5

25 - 24.2 - 16.7

20 - 22.9 - 28.1

17 - 15.0 - 34.4

Using the simulation results, the plant implemented the new ball charge and intermediate grate design

to the mill. After the implementation, sampling surveys were performed to see the change in circuit

performance. It was concluded that the capacity of the circuit increased up to 12.7-20.5% for different

cement types with the proposed optimization. Table 5 gives the capacity figures obtained after

optimization. For the product “B” operation, modelling studies were performed and the change in the

breakage rate after the optimization was assessed. Fig. 5 gives the change in breakage rate after

optimization for “B” type cement grinding. As can be seen from Fig. 5, the breakage rate of the finer

sizes was increased with the new ball charge; hence the fines generation was increased inside the mill.

Table 5 - Capacity of the circuit after optimization for different product types

Capacity, tph

Product type Before optimization After optimization

A 155.1 180.0

B 178.2 205.0

C 161.8 195.0

D 141.9 160.0



Fig. 5 - Change in breakage rate after optimization for “B” type production (a) 1st (b) 2nd chamber

Finish milling with HPGR and downstream ball milling After the initial applications of the HPGR in cement industry, more of the grinding required has been

transferred to the HPGR by introducing various circuit configurations such as pre-grinding mode,

hybrid-grinding mode, semi-finish grinding mode and finish grinding mode (Kellerwessel, 1993,

1996). In this case study, the effect of pre-grinding with HPGR in a finish mill circuit has been

assessed for open and closed circuit applications of the HPGR. In pre-grinding applications, the

HPGR reduces the work done by the downstream ball milling. In closed circuit HPGR applications,

this effect is more distinct compared to open circuit.

During blended cement production with open circuit pre-grinding with HPGR either the cement

clinker or all the raw materials can be processed together through the HPGR but in closed circuit pre-

grinding applications all the raw materials pass through the HPGR and the classify fines are sent to

ball milling.

In this study open and closed circuit HPGR applications in finish milling stage is compared with the

conventional finish milling circuit without HPGR in terms of energy efficiency. Aydogan et al. (2006)

discussed the benefit of the HPGR for different circuit configurations applied in the cement industry.

Two of the case studies published by Aydogan et al. (2006) are considered in this discussion. The

simplified flowsheets of the circuits are given in Fig. 6 and the details of the circuits are given in

Table 6. Case (a) is a conventional closed circuit ball mill operation. Case (b) recycles aportion of the

HPGR discharge material to its feed to improve the efficiency of the compressed bed breakage

between the rolls. In Case (c), the HPGR is running in closed circuit with the air separator with the

coarse product from the separator recycled to the HPGR for further breakage and the separator fines

processed through the ball mill circuit. Also since the HPGR circuit product is now fine enough, the

ball mill has been converted to single-compartment mill with a finer ball size composition (compared

to that in a conventional two-compartment mill) to maximize the benefit of this closed circuit HPGR

application.



Fig. 6 - Simplified flowsheets of the finish milling (a) without HPGR (b) with open circuit HPGR (c) with

closed circuit HPGR

Table 6 - Details of the circuits

Case study 1 conventional

Case study 2* open circuit HPGR

Case study 3* closed circuit HPGR

Raw materials fed Clinker (95%) Gypsum (5%)

Clinker (74%) Gypsum (5%) Trass (21%)

Clinker (94%) Gypsum (6%)

Capacity (tph) 55.3 71.7 132.0 F80 (mm) 14.5 16.0 18.5

Fresh feed WI (kWh/t) 11.5 11.84 11.07

HPGR – roll length (mm) – roll diameter (mm) – power draw (kW)

– – –

550 1200 640

1450 1000 1058

Ball mill – power draw (kW) 1811 1480 1800

*Aydogan et al. (2006)

During the comparison of the performance of these three circuits, the specific energy consumptions

were considered together with the reduction ratios. As can be seen from Table 6, the Bond work

indices of the fresh feeds are close to each other and therefore the effect of the material characteristics

on the energy consumption is eliminated for ball mills. Table 7 gives the specific energy

consumptions for the circuits and reduction ratios. Specific energy consumptions are given separately

for the HPGR circuit and overall HPGR-Ball mill circuit.

(a) (b)

(c)



Table 7 - Specific energy consumptions and reduction ratios

HPGR circuit sp.en.consumption (kWh/t)

Overall sp.en.consumption (kWh/t)

Overall reduction ratio (F80/P80)

Final product size (micron)

Case study 1 Case study 2* Case study 3*

– 8.93 8.02

32.74 29.57 21.65

483 235 660

30 68 28

*Aydogan et al. (2006)

From Table 7, it is obvious that the overall specific energy consumption of the finish milling circuit

can be significantly lowered with the closed circuit HPGR application. In case study 3 the overall

specific energy consumption is considerably lower and the final product size is also considerably

lower. The closed circuit HPGR is much more energy efficient. On the other hand the open circuit

application of the HPGR has no significant effect on the reduction of the specific energy

consumption. Even though the overall specific energy consumption in case study 2 is slightly lower

than in case study 1, the final product size in case study 1 is finer than in case study 2. So these two

conditions seem to be similar in terms of the energy consumption when considering the final product

sizes. Aydogan (2002) stated that the capacity of the finish milling circuits can be increased by open

circuit HPGR application with slightly reduced specific energy consumption. Results obtained in case

study 2 verifies this statement.

Further grinding of finish cement with more additives using vibratory ball mills for cost effective production For cost-effective operations and quality improvement in the cement industry, additives exhibiting

cementitious properties are blended in different ratios with reference to the cement type required (EN

197-1:2000). Various cement additives such as pozzolana, blast furnace slag, limestone, fly ash etc.

are fed to the circuit in different ratios along with the cement clinker and ground to different fineness

values, to meet the requirements for the specific cement type. Introducing more additives to the

system requires finer grinding to compensate the end-product quality in terms of the strength.

In this study, experimental work was conducted using vibratory ball mills. The finish cement from the

cement plant was further blended with additives and ground to finer sizes using 2-stage vibratory ball

milling. The schematic view of the experimental procedure is given in Fig. 7. It is known that the

efficiency of the ball milling decreases with grinding to finer sizes, therefore in this study, vibratory

ball mills which utilize more energy replace this ball mill for fine grinding. In the plant cement

clinker, gypsum and trass are interground to obtain a finish cement. In the context of the study the

finish cement was blended with fly ash for further grinding with vibrating mills at a capacity of 7 tph.

At the end of the study a similar quality cement is achieved efficiently with less clinker consumption.

As can be seen from Fig. 7, 24% clinker is added to the circuit. In the first stage of vibratory ball



milling relatively coarser media is used in this stage compared to that in the second stage. Table 8

gives the design parameters of the vibratory ball mills.

Fig. 7 - Simplified flowsheet of the cement grinding circuit followed by 2-stage VBM

Table 8 - Design parameters of the VBM

Vibration frequency (rpm) 1160

Motor power (kW) 75

Internal Diameter (m) 0.78

Internal Length (m) 1.2

Operational ball load % 90

Fig. 8 gives the size distributions of the raw materials fed to the circuit. In Fig. 8, finish cement

represents the air classifier fines. In the ball milling circuit the clinker, gypsum and trass are ground to

finish cement size. Then the finish cement is blended with fly ash. Samples were taken from the VBM

discharges. The product size distributions of the VBMs are given in Fig. 9 together with the VBM

circuit feed.

Fig. 8 - Particle size distributions of the raw materials



Fig. 9 - Size distributions around the VBM circuit

The efficiency of the circuit is assessed by considering the specific energy consumption and product

quality. For this purpose the circuit was considered in two parts. The first part is represented by the

closed ball mill in which the finish cement is obtained and the second part covers the whole circuit.

Table 9 gives the specific energy consumption of the circuit during finish cement production and for

the overall process. Table 9 also gives the 28 days strength values for the mortars tested in accordance

with EN 196-1:1994. Results showed that the product quality is satisfied by using only 24% cement

clinker. Applying the VBM after finish milling reduced the clinker usage from 60% to 24% with an

extra 7 kWh per tonne of energy used. When considering the energy required for 1 tonne of cement

clinker production which is 30-40 kWh, using energy efficient VBMs in fine grinding significantly

reduces the cost of the operation.

Table 9 - Specific energy consumptions and product qualities

sp.en.consumption (kWh/t)

28 days strength (MPa)

Finish cement End-product

28 35

37.6 38.1

CONCLUSIONS Optimizing the ball size distribution in finish milling of cement using modelling and simulation tools

resulted in higher production rate, reduced overall specific energy consumption and reduced

component wear cost.

The use of HPGRs in finish milling circuits as pregrinder increases the capacity of the existing

circuits by reducing the work done by the downstream ball milling. The closed circuit HPGR

application has significant effect on the reduction of the overall specific energy consumption of the

circuit.



Using vibratory ball mills after finish cement production to reshape the final product with increased

additive ratio results in a cost-effective process. However for high capacity finish milling circuits the

application of the VBMs is not practical due to their low capacity.

ACKNOWLEDGEMENTS Authors would like to thank to Hacettepe University Comminution Group members for their support

during the sampling campaigns.

REFERENCES Aydogan, N.A., 2002. Mathematical Modelling of High Pressure Roller Mills, Hacettepe University,

MSc thesis.

Aydogan, N.A., Ergun, L., Benzer, A.H., 2006. High pressure grinding rolls (HPGR) applications in

the cement industry, Miner. Eng. 19 (2), 130-139.

Benzer, A.H., 2000. Mathematical Modelling of Clinker Grinding Process, Ph.D. Thesis, Hacettepe

University, Ankara, Turkey, pp.138.

Benzer, A.H., Ergun, S.L., Oner, M., Lynch, A.J., 2001. Simulation of open circuit clinker grinding.

Miner. Eng. 14 (7), 701–710.

Benzer, A.H., Ergun, S.L., Oner, M., Lynch, A.J., 2003. Case studies of models of tube mill and air

separator grinding circuits. In: Lorenze, V., Bradshaw, D.J. (Eds.), Proceedings: XXII

International Mineral Processing Congress, pp. 1524–1533.

Bhatty, J.I., Miller F.M., Kosmotka, S.H., 2004. Innovations in Portland Cement Manufacturing.

Portland Cement Association, USA.

EN 196-1:1994, Methods of testing cement - Part 1: Determination of strength

EN 197-1:2000, Cement – Part 1: Composition, Specifications and Conformity Criteria for Common

Cements.

Fairs, G.L., 1954. A method of predicting the performance of commercial mills in the fine grinding of

brittle materials. Trans. Inst. Min. Metall. 63, 211–240.

Gross, J., 1938. Crushing and grinding. US Bureu of Min. Bull. 402, 1–148.

Hoffler, J.A., Herbst, J.A., 1990. Ball mill modelling through microscale fragmentation studies: fully

monitored particle bed comminution versus single particle impact tests. In: The 7th Euro

Symposium on Comminution, Ljubjana, Yugoslavia, pp. 381–397.

Lynch, A.J., Oner, M., Benzer, A.H., 2000. Simulation of a closed cement grinding circuit. ZKG, No.

10, pp. 560–568.



Kellerwessel, H, 1993. High pressure particle bed comminution of mineral raw materials,

Aufbereitungs-Technik, 34(5):243-249.

Kellerwessel, H, 1996. High pressure particle bed comminution – state of the art, application, recent

developments, Engineering and Mining Journal, pp 45-52.

Napier-Munn, T., Morrell, S., Morrison, R.D., Kojovic, T., 1996. Mineral Comminution Circuits:

Their Operation and Optimisation. In: Napier-Munn (Ed.), JKMRC, 342p.

Narayanan, S.S., 1985. Development of a Laboratory Single Particle Breakage Technique and its

Application to Ball Mill Modelling and Scale-up, Ph.D. Thesis, University of Queensland.

Pauw, O.G., Mare, M.S., 1988. The determination of optimum impact-breakage routes for an ore.

Powder Technol. 54, 3–13.

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Rumpf, H., 1973. Physical aspects of comminution and new formulation of a law of comminution.

Powder Technol. 7, 145–159.

Schonert, K., 1972. Role of fracture physics in understanding comminution phenomena. Trans. Soc.

Min. Eng. AIME 252, 21–26.

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University of Queensland.

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The Importance of Evaluating Grinding Performance

R L Koenig1 and K T Broekman2

1. FAusIMM(CP), Project Director Hawsons Iron Ore Project, Carpentaria Exploration Pty Ltd, 345 Ann Street, Brisbane Qld 4000. Email: [email protected]

2. MAusIMM, Project Development Manager, Carpentaria Exploration Pty Ltd, 345 Ann Street, Brisbane Qld 4000. Email: [email protected]

ABSTRACT

Previous papers at Metplant conferences have covered the options for crushing and grinding circuit selection. High Pressure Grinding Rolls (HPGR) may have considerable advantages over autogenous milling circuits but test work and evaluating different grinding circuits must be done to establish which circuit provides the best option.

The Hawsons low grade magnetite project provides a specific example in using thorough test work, careful simulations and detail in design to allow the evaluation of grinding circuit performance. Fortunately tools are available for this evaluation and the test work for the Hawsons iron ore is detailed in this paper. The results of the test work and simulation indicates that although the Hawsons iron ore is competent in terms of autogenous milling given the medium hardness of the ore in the Drop Tests, the HPGR testing indicates that the use of HPGR is likely to use half the energy requirement. This is a characteristic of the Hawsons iron ore and is likely to be valid for other iron ores in the Braemer Iron Formation which stretches from the Hawsons lease in New South Wales to Razor Back and South Dam in South Australia.

The Hawsons magnetite iron ore resource has been known from the early 1960s when CRA explored the area. However, the Pilbara region was discovered in this era and this led to development of high grade direct shipping grade hematite ores. Carpentaria Exploration (Capex) rediscovered the magnetite whilst undertaking multi-commodity exploration in the Broken Hill district during 2009.

The magnetite grade was found to be low at around 16% Davis Tube mass recovery (DTR) but despite this, the project provides a positive business case as a result of the following key factors:

The mineralisation is particularly soft for crushing and grinding with a Bond Work Index of 6.3.

The mineralised rock fractures easily preferentially along grain boundaries.

The location is close to mining infrastructure and a skilled workforce in Broken Hill, 60km to the North East.

It is close to road and rail infrastructure for access to the resource.

The only contaminant in concentrate is silica.

The resource contains many billions of tonnes of magnetite mineralisation.

The low Bond Work index is a key characteristic of the Hawsons iron ore and the use of autogenous milling and HPGR circuits is compared to other iron ores indicating why the plant energy requirement may be just as important as iron ore grades.


INTRODUCTION

Carpentaria Exploration is a Brisbane based exploration company with exploration leases mostly in the Broken Hill area in NSW. Carpentaria Exploration has outlined a large magnetite resource at their Hawsons Prospect near Broken Hill (Fig. 1). The resource is low grade at around 16% weight recovery (DTR) but has some advantages in terms of location and soft ore hardness.

Resource drilling for the project has been completed and this has provided the first JORC compliant 1.4 billion tonne of inferred resource for the project. Preliminary laboratory testing has indicated excellent concentrate grades at grind sizes less than 38 micron. While exploration drilling and testing have indicated a potential of ±11 billion tonnes of iron ore mineralisation, which indicates that the size of the project need not be limited.

Figure 1. Location map

Agreements have been signed for the funding of all exploration and feasibility work for a project to produce 20 Mtpa of magnetite concentrate per year. The resources are available for this scale of project.

The Hawsons Iron Ore resource is part of the Braemar Iron Formation, a geological structure that extends from the Carpentaria Exploration Lease area to South Dam about 250kms to the west parallel to the rail transport corridor. The successful development of the Hawsons Iron ore project could lead the way to development of a number of iron ore projects based on the Braemar Iron Formation.


Despite the low grade magnetite content, the project continues to provide a positive business case in feasibility studies indicating that perhaps grade is not everything for magnetite deposits. This paper compares the grades and power requirements with other operating and developing projects in magnetite to determine the critical indicators.

Mineralogy

A key to the softness of the ore and also the grind size requirements is indicated in the mineralogy. Fig. 2 is a reflected light magnification of the ore indicating the square shaped magnetite, the quartz and the carbonate. At this stage of the project, the carbonate appears to be the key to why the ore is so soft. Also note that the magnetite is free of silica inclusions.

Figure 2. Photograph of the mineralogy in the Hawsons magnetite iron ore.

The ore fractures easily and mostly around grain boundaries so that RC chip samples typically are 70% passing 25 microns and this material can be separated as high grade magnetite concentrate.

The mineralogy indicated is very consistent across the ore resource with an occasional enriched silica lens.

Ore Grades and Mining Costs

The magnetite grades at Hawsons as measured using Davis Tube testing, are low compared to other magnetite deposits. Hawsons grade averages 15.5% Davis Tube weight recovery but can vary from

Magnetite Muscovite,

+ biotite

Quartz

Carbonate


10% to 25% DTR at concentrate grades of >68% iron. This compares to Savage River Grades of 35 to 45%, Taconite grades of 30% to 50% and the Pilbara banded iron formations of around 30%.

Mining costs are influenced by both the magnetite weight recovery as well as the waste to ore ratio. Thus the shape of the orebody as well as the grade of the orebody can come into play. To illustrate the effect of each, Table 1 below indicates the amount of material required to be moved both with changes in grade as well as the waste to ore ratio.

To Produce 20Mtpa Concentrate Overburden at 30% grade

Grade ROM Waste : Ore All Material

% Million Tonnes

Ratio Mtpa

16.00% 125 0.25 83

30.00% 67 0.50 100

40.00% 50 1.00 133

50.00% 40 2.00 200

60.00% 33 4.00 333

Table 1. Tonnages Required to Produce 20 Mtpa of Magnetite concentrate at Different Grades

and Waste to Ore Ratios.

The Hawsons Iron Ore Project plans to produce 20 Mtpa of magnetite concentrate and has a grade of 16% but has an overall waste to ore ratio of 0.27. The advantageous waste to ore ratio is caused by the thick sedimentary deposit type as illustrated in the geological cross section in fig. 3. So despite the low grade, the ease of mining does reduce the overall mining costs.

Figure 3. Geological Cross Section


Comminution Energy Costs

Other papers have described the potential grinding process options (McNab et. al., 2009) and in testing the Hawsons magnetite ore, both autogenous and HPGR type circuit configurations were trialled. Test work and simulations were done on diamond drill core by JKTech Laboratory (2010), the results are provided in Table 2.

Ore Grade % DTR Weight Recovery 18%

Bond Ball mill work index kWh/t 6.3

JK Drop Weight Axb 51.1

Specific Gravity t/m3 3.00

HPGR Type A ESP (kWh/t) 2.22

Unconfined Compression Strength

MPa 50‐90

Table 2. Grinding Circuit Design Parameters.

These test results were used to simulate 3 different grinding circuits which included:

Hawsons AB Circuit consisting of an autogenous mill, rougher magnetic separation, ball mill and finisher magnetic separators.

Hawsons SS AG Circuit consisting of a single stage autogenous mill circuit similar to Hibbing Taconite.

Hawsons HPGR circuit with screening, rougher magnetic separation, ball milling and finisher magnetic separation.

The results of the simulations can be compared to the typical grinding options presented in the McNab et al. paper. Figure 4 shows the significantly lower power requirement for the Hawsons magnetite iron ore.

Figure 4. Circuit Energy Comparison of Different Hawsons Circuit Options and these Proposed McNab et. al. Options.


The energy requirements for the Hawsons magnetite ore are significantly below those of other magnetite ores and the HPGR options require less than half the power requirement of a typical autogenous mill circuit. These results have resulted in the selection of just the HPGR circuit for the feasibility studies and the fine grinding requirements have yet to be finalised.

To put this into a perspective which takes into account the grade of the ore, the grinding energy requirement has been recalculated in terms of kWh/t of concentrate. The energy has been calculated from available data on operations in Australia and North America as shown in Fig. 5. This fact has resulted in the energy requirement being considerably less at 39kWh/t of concentrate than existing and planned operations overseas despite the low grade at Hawsons.

Figure 5. Comparison of Grinding Circuit Energy Requirements per tonne of concentrate for Hawsons and other Magnetite Iron Ore operations.

The Hawsons magnetite grades are low by industry standards but this is more than compensated for with by the softness of the ore and studies to date have provided a positive business case for the project at long term iron ore prices.

Other resources along the Braemar Iron Formation in South Australia we believe have similar grades and ore characteristics and therefore the potential for additional iron ore projects along this formation similar to the Hawsons iron ore project is great.

CONCLUSIONS

The in situ Davis Tube Recovery grade and concentrate grade is not everything for a highly beneficiated (processed) ore like a magnetite resource. The power requirement per tonne of concentrate product is likely a better determinant of economic potential/performance. When assessing magnetite prospects it is essential to consider the likely Run of Mine Bond Work Index as this is potentially almost as important as in situ grade. This early stage engineering analysis could potentially be overlooked by exploration geologists who are typically tasked with making first pass assessments of undeveloped magnetite prospects. Magnetite is a special case where engineering parameters are just as important a consideration when evaluating the ore resource against other forms of iron ore resources.


The production of magnetite and pellets in North America in the 1950’s from Taconite ores was a difficult starting point with Taconite being one of the hardest ores. These recent Australian developments such as those for the Hawsons Project have certainly helped in developing world wide magnetite operations which can only benefit from assessments of the grinding performance of the ore in evaluating the magnetite resource.

ACKNOWLEDGMENTS

The authors acknowledge the permission of Carpentaria Exploration to publish this paper and the support of their Chief Geologist Doug Brewster.

REFERENCES

1. McNab, Jankovic, David, Payne 2009, Processing of Magnetite Iron Ores – Compared Grinding Options, Iron Ore Conference, Pert WA.

2. JKTech Laboratories - Confidential report JKTech Job No.10242 – September 2010



Increasing Capacity and Efficiency of Grinding Circuits with High Frequency Screens

J Wheeler1 and B Packer2

1. Product Application Engineer, Derrick Corporation, Buffalo, New York, USA.

Email: [email protected] 2. Director, Western Process Equipment, Perth WA 6106.


ABSTRACT

The objective of this paper is to examine the operation of high frequency screens in closed circuit grinding, where the outcomes achieved indicate favourable economic gains at the comminution stage and downstream processes, including flotation, dewatering, and filtration operations. The classification process in grinding circuits has been slowly evolving within the last century. The history of classification devices brings to mind rake classifiers and spiral classifiers, which although having higher efficiencies compared to hydrocyclones, were not able to keep pace with the evolution of concentrator equipment treating higher tonnages. This higher tonnage throughput has tended to drive the concept of utilising hydrocyclones as the classification stage in closed grinding circuits. The comminution process is commonly the most energy intensive stage of mineral processing. Researchers have determined that the classification stage in the comminution process has the greatest potential for improvement.

This paper is intended to inform mineral process engineers of the benefits that high frequency screens offer over hydrocyclones. This will be done by reviewing the purpose for the evolution of classification in the comminution process, documenting past and recent studies, and explaining the modern technology that is currently available.

HISTORY

Since the mid 1920’s when the early pioneering test work was conducted by E. W. Davis, (1925) to compare the performance of rake classifiers, spiral classifiers and vibrating screens in closed grinding circuits, some remarkable relatively recent developments have been seen. From this early work by Davis, it was determined that vibrating screens provided higher grinding mill capacities and superior control over the grind size. Unfortunately the apparent benefits could not be realized in practice at that time because of the limitations in screening technology. While screens were commonly used for mineral processing, finer separations were considered impractical due to low machine capacities and high maintenance requirements relating to the screen media which suffered from blinding and high wear rates. Another study by E. R. Albert (1945) determined that “lower grinding costs are possible with screen circuits”. The lower capital costs of rake and spiral classifiers led to the standardization of these devices for closing ball mill circuits. By the 1950’s advancements in high capacity pumps opened the door for hydrocyclones to become the preferred equipment for ball mill classification. Hydrocyclones allowed for higher tonnage with even less capital investment but did not improve classification efficiency. By the 1960’s researchers forecast a need for improved grinding efficiencies based on increasing energy costs and lowering ore grades. Two studies by Hukki and Eland (1965) and Hukki (1967) concluded that “the master key for great improvements in capacity and in energy consumption in closed grinding circuits is improved sharpness of separation”. These studies also concluded that to reduce over grinding and increase separating efficiency an accurate sizing device was required. Whilst the concept of classification by physical size was sound and clearly understood together with the understanding that the salient characteristic of the ideal classification



device is to immediately exit particles to the downstream process upon being ground to liberation size, at that time there were no cost effective technologies available to meet these objectives.

ADVANCES IN CLASSIFICATION TECHNOLOGY APPLIED TO GRINDING

CIRCUITS

As previously mentioned, the early classifiers used to close grinding circuits were mechanical rake and spiral classifiers however operational costs and capacity limitations were major impediments to the advancement of these technologies.

Hydrocyclones became the popular choice in the 1950’s to handle the increasing demand for higher production rates despite the fact that generally the separating efficiency still remained in the 45 % to 65 % range. The fundamental classifying principal of a hydrocyclone is a function of particle size and specific gravity and therefore does not provide a true size separation. This often results in heavy liberated mineral already at the target grind size being returned to the mill for overgrinding and generation of slimes instead of reporting to the downstream process. Although fines in cyclone overflow is typically more pronounced in processes where there is a significant difference in specific gravity between the target mineral and gangue, regardless of the ore, hydrocyclones have a tendency to bypass liquid with fines to the coarse underflow stream without classification.

Separation techniques based on true particle size by screening has always been understood and accepted as the recognized standard for coarse particle classification. However, for fine particle separations, screens were considered to be impractical and uneconomical due to factors such as low capacity, high screen panel consumption, and blinding.

Screen manufactures recognized that if an effective fine screen could be developed to close grinding circuits it would offer significant benefits to the industry. The identified requirements for a successful design were as follows:

Rugged construction and mechanical reliability, Large tonnage capacity, High separation efficiency, Resistance to blinding, Wear resistance surfaces, Energy efficient / Low power consumption, Ease of maintenance, Cost effectiveness, and Produce favorable metallurgical results.

FACTORS THAT AFFECT WET SCREENING

These factors (Valine and Wennen, 2002) include:

1. Feed Rate. The capacity of a screening machine is defined as the optimal feed rate to meet the desired product specifications. Feed rate, usually expressed as dry mass flow (t/h), is one of the more critical factors affecting screen performance. The capacity of the screen will determine the number of screening machines required. Exceeding capacity (or over feeding a screen) will result in the misreporting of undersize particles and fluid to the oversize stream as well as a reduction in screen surface life. Depending upon other factors, the optimal feed rate can be exceeded to some extent without a significant decrease in efficiency. The capacity of a wet fine screen is best determined by full-scale testing to optimize all factors affecting screen performance.



2. Feed Density. As understood, undersize particles are transported through the screen openings by the fluid and therefore, the volume fraction of fluid will affect screen efficiency. Screening efficiency will increase with decreasing feed pulp density. From a practical standpoint, a screen feed density of roughly 20 % solids by volume has been found to be a reasonable compromise, independent of dry solids specific gravity. To maximize undersize efficiency (the correct placement of undersize), the screen feed slurry could be even lower, perhaps as low as 10 to 15 % solids by volume. It has also been shown that it is usually more beneficial to add water to the screen feed slurry than to add the same amount of water directly to the screen surfaces with spray nozzles.

3. Feed Size Distribution. The size distribution of the material fed to a screen is one of the more important factors affecting both capacity and performance of a wet screening machine. The oversize particles must be conveyed off the screen and capacity usually decreases as the amount of oversize increases and creates congestion at the screen surface apertures. Another important factor is the amount of near-size material in the screen feed. Near-size material is defined as the particles that are 2 mesh-size equivalents larger and smaller than the screen opening. Near-size, oversize material inhibits the ability of the undersize material to get through the screen openings and, in some cases, can cause some plugging problems. Selection of screen media is quite important when dealing with significant amounts of near-size material.

4. Screen Opening and Open Area. The larger the opening, the greater the machine capacity. Conversely, as the desired separation size decreases, so does machine capacity. For example, say that full-scale tests determine that machine capacity is 100 t/h with a 2350 µm (60 mesh) screen opening. Machine capacity could drop to 20 to 40 % with 150 µm (100 mesh) openings. At a given size, the open area of a particular screen surface also affects capacity. To increase screen panel life, it may be desirable to use a more robust screen cloth with lower open area. However, doing so will result in a lower machine capacity.

5. Screen Width. Extensive research has demonstrated that screen width is the primary

parameter needed to maximize capacity and efficiency in fine wet screening applications

(Coleman, 1986). This information has led to the development of multiple feed point screens (Figs. 1 and 2) which achieve efficiencies above 90 % in numerous high tonnage applications.

Fig. 1: Multifeed Flow Distribution Fig. 2: Multifeed Screen

The unique Multifeed screen, having started to meet the required design objectives of finer wet screening applications at higher tonnages, has been used in a variety of mining operations. With the ever increasing demands of higher tonnage mineral dressing operations, the main problem relating to



such screens continued to be their ability to reach these new capacities at operating P80s demanded by modern plants. With a focus on this primary objective, an entirely new and innovative approach was required to produce efficient and low power consumption wet screening equipment to meet ever increasing operating tonnages (and probably lower grades) with reasonable Capex and Opex costs including acceptable screen life and low maintenance requirements. Now with increasing machine capacity as the main focus, the new machine to be developed had to be based on the understanding that the screen deck width is the main design criteria for separation performance.

DERRICK STACK SIZER TM

To meet these criteria, in 2001 the Derrick Corporation released a new high frequency screen called the Stack Sizer (Figs. 3 and 4). This is the highest capacity and most efficient fine sizing wet screen produced to date. The increased capacity widened the practicality of fine screen applications in mineral processing plants. The main difference of the Stack Sizer design, when compared to the early Multifeed screens, is the five short screens in parallel, stacked one above the other. This Stack Sizer has an effective total width of 5.1 metres with extremely low power input. Driven by 2 off 2.5 hp motors it has the lowest power input per metre width of screen available and has 2 to 5 times more capacity than the Multifeed screens.

Fig. 3: Stack Sizer Fig. 4: Stack Sizer Flow Distribution

HIGH OPEN AREA URETHANE SCREEN SURFACES

Perhaps even more revolutionary than the Stack Sizer, was the development of a tensionable high open area urethane screen surface. This provides long life, high capacity urethane screen surfaces with typical effective open area ranging from 35 to 45 %. The design of these urethane screen surfaces virtually eliminates blinding with screen panel life in excess of one year for several applications. The urethane screen surfaces are currently being made with apertures as fine as 75 µm.

CHARACTERISTICS OF CLOSED CIRCUIT GRINDING WITH SCREENS



According to Albert, (1945) and Hukki and Allenius, (1968), the most substantial changes that occur when substituting hydrocyclones with screens in closing circuit grinding are:

Increased productivity/throughput, Reduced circulating load, Lower power consumption per ton, Reduced overgrinding of valuable minerals/generation of slimes/tailings, Consistent product size distribution and density, Increased recovery of valuable minerals, Reduced reagent consumption in flotation, and Reduction in overall operating costs.

The grinding circuit media charges may need to be adjusted to balance for the decrease in circulating load and there is a very high potential for an increase in new feed tonnage. Areas that should be taken into consideration by the mineral process engineer are the downstream unit operations and because of the reduced overgrinding, beneficiation methods of mineral separation used in the plant. A careful characterization, by a device such as a QEMSCAN, of the classification circuit can give the optimum mineral liberation size prior to any changes. After substituting the hydrocyclones with screens the operator should keep in mind that the grinding circuit product will have a lower specific area. This can probably reduce the reagent consumption in the downstream flotation processes.

Another major benefit that screen size classification offers is control over the top size of the particles that leave the grinding circuit. For example, copper flotation circuits often have poor recovery of particles coarser than 70 mesh. A screen can be used to reduce the amount of particles coarser than 70 mesh that leave the grinding circuit. Plant operators should look at all unit operations to take full advantage of the improvements that screens offer in the grinding circuit. Other processes, such as filtration often benefit from a reduction of fines.

CASE STUDY – CIA MINERA CONDESTABLE S.A.

The following section refers to a case study by Delgado et al., (2007):

CIA Minera Condestable is an underground copper mine located south of Lima and has been in production since 1963. Condestable produces copper concentrates that contain Au with a process capacity of 4100 t/d. A recent expansion project applied Derrick Stack Sizer high frequency screens as the classifier to close the grinding circuit, completely replacing the hydrocylones. The results obtained are presented in Table 1.

Table 1: CIA Minera Condestable – Expansion Results with Stack Sizers

Actual

After Expansion

Gains/Losses

Daily Production TMSD 4100 5500 +34 %

Annual Production KT 1.497 2.008 +34 %

Feed Cu % Grade 1.329 1.245 -6 %

Cu Recovery 90 % 90 %

Concentrate Production TMS 68845 86516 +26 %

A 34 % increase in total production, to the amount of an additional 1400 t/d, was due to an increase in the grinding circuit classification efficiency. While the plant was operating with hydrocyclones it was processing 171 TMS/h. After changing to the high frequency screens it



increased production to 229 TMS/h. Prior to the installation of Derrick Stack Sizer screens at Condestable, full-scale screening tests were conducted at Derrick’s laboratory in Buffalo, New York, USA and the test results are presented below in Table 2.

Table 2. Screen Test Results with Derrick Stack Sizer Screen

Compañia Minera Condestable S.A. – Mill Discharge 1 and 2– 14 March 2006, Buffalo, NY, USA

Test

Feed 1 Oversize Undersize Efficiency 2 at 70 mesh

Slurry Dry Solids Solids Weight Solids Weight Solids Plus Minus Oversize Undersize Overall

No. (m3/h) (MTPH) (%) (%) (%) (%) (%) 70 Mesh 200 Mesh (%) (%) (%)

Tests with 0.23 mm (70 mesh) polyurethane panels1 187 126 46.8 52.3 81.7 47.7 31.8 9.8 52.2 91.2 88.3 89.8

2 204 137 46.8 52.7 82.1 47.3 31.6 8.2 55.1 92.7 89.4 91.1

15 170 130 51.2 53.4 84.7 46.6 0.0 10.3 56.5 90.6 88.4 89.6

16 187 143 51.2 54.5 84.3 45.5 0.0 8.5 59.5 92.5 87.5 90.1

Tests with 0.30 mm (50 mesh) polyurethane panels

3 204 137 46.8 50.5 82.5 49.5 32.4 11.3 54.2 98.7 84.8 90.5

4 227 152 46.8 50.3 81.8 49.7 32.6 11.5 53.7 98.5 85.1 90.6

13 170 130 51.2 55.6 83.9 44.4 0.0 10.5 56.2 95.3 81.9 88.3

14 193 148 51.2 50.5 83.5 49.5 0.0 13.1 52.8 92.6 86.0 89.1


5 227 149 46.0 46.3 81.9 53.7 33.4 17.1 50.6 97.5 85.1 89.6

6 250 163 46.0 45.0 81.5 55.0 33.9 18.8 47.7 96.9 85.2 89.4

11 182 141 51.6 51.2 83.4 48.8 0.0 13.0 55.0 96.9 83.6 89.1

12 199 154 51.6 52.7 82.4 47.3 0.0 13.1 55.1 96.8 81.4 87.7


7 256 167 46.0 39.4 82.5 60.6 35.7 25.5 45.3 96.5 89.8 91.9

8 284 186 46.0 39.3 82.2 60.7 35.8 25.3 45.8 96.3 91.1 92.8

9 216 167 51.6 40.0 84.8 60.0 40.9 25.6 44.6 96.2 89.2 91.5

10 250 193 51.6 43.4 83.5 56.6 39.9 23.5 47.0 97.5 86.4 90.0

The Derrick Stack Sizer screens continue to operate without any major problems. The operation parameters defined by the Derrick laboratory tests have been kept and are giving the expected results. The control of the percentage of solids in the screen feed has been determined as the most important parameter to produce the best screen performance. Table 3 shows the technical comparisons between the two classification systems, prior to and after the expansion.

Table 3: Technical Comparisons Hydrocyclone – Screen

Operational Parameters Circulating Load Classification Efficiency Tonnage Processed TMS/Hr Size Distribution Control

Hydrocyclone 204 %

62-64 % 59

Variable

Stack Sizer 96 % 85 %

68 Uniform

Other important results obtained at Condestable beneficiation plant was a 17 % increase in the amount of ore processed, better operational control, more uniform particle size distribution produced by the grinding circuit, and a reduction in power consumption per ton. The increased production rate was achieved without an increase in overall energy consumption. An economic evaluation of the investment showed the purchase and installation of one Derrick high frequency screen Stack Sizer was US$300,000.00. With the economic gain generated by the 17 % increase in production that resulted in US$2,630,000.00 per year, the return of investment was 1.4 months.



Several other mining companies in Peru have replaced hydrocyclones with Stack Sizers in their grinding circuits with similar results. As showing in Table 4, a total of 31 Stack Sizers are operating in Peruvian grinding circuits to date. This trend was started in 2004 by Sociedade Minera El Brocal. Brocal was considering the purchase of an additional ball mill to increase production rates. Following full-scale Stack Sizer screen tests at Derrick, the installation and operation of Stack Sizers at Brocal clearly demonstrated that production rates, metal recovery, and power consumption could be significantly improved with better classification (Aquino and Web, 2007).

Table 4: Application of Stack Sizers in Peru

PERU DERRICK STACK SIZER APPLICATIONS

YEAR COMPANY LOCATION MODEL QUANTPANEL

mm ORE TYPE Position

2004 SOC. MIN. EL BROCAL COLQUIJIRCA 2SG48-60W-5STK 3 0.5 Lead/Zinc Grinding replace cyclones

2005 CÍA. MIN. ARES EXPLORADOR 2SG48-60W-5STK 1 0.18 Gold/Silver Grinding replace cyclones

2005 CÍA. MIN. ARES ARES 2SG48-60W-4STK 1 0.18 Silver/Gold Grinding Mix with cyclones

2006 SOC. MIN. EL BROCAL COLQUIJIRCA 2SG48-60W-5STK 1 0.5 Copper Grinding replace cyclones

2006 CONDESTABLE CONDESTABLE 2SG48-60W-5STK 4 0.3 Copper Grinding replace cyclones

2006 COLQUISIRI COLQUISIRI 2SG48-60W-4STK 1 0.43 Lead/Zinc Grinding replace cyclones

2006 CÍA. MIN. ARES ARCATA 2SG48-60W-5STK 1 0.18 Silver Grinding Mix with cyclones

2006 CATALINA HUANCA C.HUANCA 2SG48-60W-5STK 1 0.3 Lead/zinc Grinding replace cyclones

2006 MILPO EL PORVENIR 2SG48-60W-5STK 1 0.3 Lead/zinc Grinding replace cyclones

2006 MINERA CORONA

2SG48-60W-2STK 1 0.23 Copper/Zinc Discharge Rod mill

2007 MILPO EL PORVENIR 2SG48-60W-5STK 1 0.3 Lead/Zinc Grinding replace cyclones

2007 SOC. MIN. EL BROCAL COLQUIJIRCA 2SG48-60W-5STK 1 0.7, 0.23 Lead/Zinc Grinding replace cyclones

2007 LOS QUENUALES ISCAYCRUZ 2SG48-60W-5STK 4 0.23, 0.3 Lead/Zinc Grinding replace cyclones

2007 MILPO CERRO LINDO 2SG48-60W-5STK 4 0.3, 0.35 Lead/Zinc Grinding replace cyclones

2008 LOS QUENUALES YAULIYACU 2SG48-60W-5STK 3 0.23, 0.3 Lead/Zinc Grinding replace cyclones

2008 MILPO EL PORVENIR 2SG48-60W-5STK 1 0.3 Lead/Zinc Grinding replace cyclones

2008 CÍA. MIN. ARES EXPLORADOR 2SG48-60W-5STK 2 0.18 Silver/Gold Grinding replace cyclones

CONCLUSIONS The history and advancement of grinding and classification technology during the last century has been a slow process compared to that in other mineral processing areas. The comminution process continues to be the most energy intensive operation in most mineral processing facilities. Studies conducted by researchers through the years have demonstrated that the largest potential improvement in the comminution process could be achieved by improved classification. Today, with the increasing cost of energy and the rising value of metals and minerals, the benefits of improved classification through the use of Derrick high frequency screens is now a practical reality. The commercial-scale replacement of hydrocyclones with Derrick high frequency screens in closed grinding circuits has demonstrated significant metallurgical and economic benefits, including increased production rates and lower power consumption per tonne treated. Downstream unit operations such as flotation and dewatering also benefit due to a reduction in overgrinding and specific surface area, resulting in improved product grade and recovery and reduced dewatering costs.



Whether it be a “Greenfield” project or a “Brownfield” expansion, high frequency wet screening in grinding circuits is here to stay.

REFERENCES

Davis, E W, 1925. Ball Mill Crushed in Closed Circuits with Screens. Bulletin University of Minnesota, No. 10.

Eared Albert, 1945. Characteristics of Screen-circuit Products. T.P. 1820, Min. Tech., May.

Hukki, R T and Eland, H, 1965. The Relationship Between Sharpness of Classification and Circulating Load in Closed Grinding Circuits. Society of Mining Engineers, Transactions September.

Hukki, R T, 1967. An Analysis of Mill and Classifier Performance in a Closed Grinding Circuit. Society of Mining Engineers, Transactions.

Hukki, R T and Allenius, H, 1968. A Quantitative Investigation of Closed Grinding Circuit. Society of Mining Engineers, AIME –Transactions Volt 241, December.

Colman K G, 1986. Selection Guidelines for Vibrating Screens in SME Mineral Processing Handbook (ed: Norman, L W).

Valine, S B and Wennen, J E, 2002. Mineral Processing Plant Design, Practice and Control (Society for Mining, Metallurgy and Exploration, Inc.: Ed: Molar, Hale, and Barrett).

Aquino, Vd. and Torres, Web, 2007. Reingeniería de Los Procesos Metalúrgicos en la Mina Colquirjica (XXVIII Convencion Minera Extemin - Perú).

Delgado, M, Diaz, G and Chambi, R, 2007. Expansión de Producción de Condestable com Innovaciones Tecnologicas de Clasificación de Molienda (XXVIII Convención Minera Extemin - Perú).


Recent Improvements to the Gravity Gold Circuit at

Marvel Loch

A Bird1 and M Briggs

2

1. GAusIMM, Plant Metallurgist, St Barbara Ltd, PMB 10, Southern Cross WA 6426. Email: [email protected]

2. MAusIMM, Metallurgical Superintendent, St Barbara Ltd, PMB 10, Southern Cross WA 6426. Email: [email protected]

ABSTRACT Gravity gold circuits have an important part in reducing operating costs and maximising recovery in modern

gold plants where coarse gold is present in the ore body. The design and selection of equipment has a vital

part in the overall effectiveness of the gravity gold circuit. There are a number of associated factors that need

to be taken into consideration when selecting the most suitable equipment for gravity gold circuits including

occupational health and safety, security and labour requirements.

This paper is a case study that outlines the validation behind the design, selection of equipment and

subsequent benefits of the recent upgrade of the gravity circuit at the Marvel Loch gold mine with the main

component to the upgrade being the installation of a Consep Acacia Reactor. Since the upgrade there has

been a significant increase in gravity recovery, which has reduced the amount of gold reporting to the

leaching circuit therefore reducing reagent consumption and overall operating costs. There has been an

increase in overall gold recovery and therefore increased gold production. The benefits to occupational

health and safety include the removal of toxic gases being released in the gold room during the previous

practice of gold concentrate calcination and the removal of the arsenic and nickel from the gold doré, which

poses a risk during refining. Manual handling and security has improved as there is no physical contact with

gold concentrates during the automated process using the Acacia reactor. The total time for gold room

personnel working in the gold room has also reduced from around six hours per day to less than one hour

per day.

With the continuing development of low grade gold deposits, there will be an increase in demand for low

capital, high recovery gravity gold circuits that can significantly increase overall production while reducing

operating costs.

INTRODUCTION St Barbara's Southern Cross Operations are centred at Marvel Loch, 30km south of the town of Southern

Cross and 360km east of Perth, Western Australia. The current operations based at the Marvel Loch

Underground mine as shown in Figure 1. Southern Cross Operations produced 121,870 ounces of gold for


the 2010 financial year. The immediate strategic focus of the operation is to reliably deliver to the plan

through to financial year 2013, while extending the mine life through expansion of reserves (St Barbara

Limited, 2008).

Figure 1: Location of Southern Cross Operations

Geology Marvel Loch is the largest of the many shear-hosted gold deposits found within the Southern Cross

greenstone belt. It lies in the neck of a mega-boudin formed by the mafic-ultramafic greenstone core of the

Polaris domain. The Marvel Loch gold deposit is hosted by a steep westerly dipping package of ultramafic,

mafic and sediments to the west, with gabbro, dolerite and sediments to the east.


Marvel Loch Underground Mine The Marvel Loch Underground (MLU) mine is the mainstay of St Barbara's Southern Cross Operations. Gold

mineralisation extends over a 1.3 km strike length and has been identified to depths of over 700 metres

below surface. The ore body comprises multiple lodes. Those currently being mined include Sherwood and

Undaunted at the North; Exhibition at the centre; and East and New at the South. These are shown in Figure

2. Mining methods include uphole benching and open stoping with rock fill where necessary. Ore production

from Marvel Loch Underground mine for the 2010 financial year was 969,519 tonnes at an average grade of

4.0 g/t.

Figure 2: Marvel Loch Underground Ore Lodes

Other Mines Open pit mining ceased at the Southern Cross Operations in July 2009. The most recent open pit mining

was from the Mercury pit from the Transvaal region 4km south of Southern Cross. The recoveries from the

Mercury ore ranged from 45-65% due to gold predominantly locked in Arsenopyrite resulting in lower overall

plant recoveries. Processing of the Mercury ore ceased in May 2010. The Nevoria ore body is currently

subject to a resource model and mining study. Elsewhere in the area, the company is seeking to leverage

off projects with a significant amount of drilling and a track record of past production. The Frasers, Corinthian

and Copperhead mines are yet to be fully evaluated following the St Barbara’s acquisition of the Southern

Cross land-holding. Extensive refractory mineralisation is present in the Transvaal region, which requires

further metallurgical test work to assess its viability (St Barbara Limited, 2008).


Marvel Loch Processing Plant

The current Marvel Loch processing plant was commissioned in 1987 by Mawson Pacific Ltd with an original

nameplate capacity of 1.2 Mtpa. The original plant consisted of single stage crushing, SAG milling, and

conventional carbon-in-leach (CIL) circuit. Since that time, the plant has undergone numerous expansions

and upgrades, including the addition of secondary and tertiary crushing, the replacement of a coarse ore

stockpile with fine ore bins, a third ball mill, extra leach tanks, the conversion of CIL to carbon-in-pulp (CIP)

and gravity circuit improvements.

The current processing plant arrangement consists of three stage crushing, fine ore bins, fine ore stockpile,

two stage grinding, gravity concentration, CIP circuit, split AARL elution, electrowinning, smelting and a

tailings storage facility.

The plant capacity is 2.5 Mtpa, however current annual mill throughput is 1.2 Mtpa with all ore mined from

the MLU mine. Since September 2009, the plant has operated on a 1 week on / 1 week off campaign milling

basis.

GRAVITY CIRCUIT PRIOR TO UPGRADE The gravity gold circuit at Marvel Loch prior to the upgrade in April 2010 consisted of screening and primary

and secondary gravity concentration of the cyclone underflow stream to recover coarse free gold.

Specifically, two bleed streams of the cyclone underflow were fed to a gravity feed box where the slurry

density was reduced prior to being screened in parallel over two 3.0 m x 1.5 m vibrating screens with an

aperture size of 3 mm. The coarse fraction (+3.0 mm) was returned to the secondary ball mills for further

grinding while the minus fraction (-3.0 mm) was collected into a hopper before being pumped to the magnetic

separator for tramp iron removal prior to primary gravity concentration using two 30 inch Knelson

concentrators (one CD model and one XD model). The gravity concentrate from the Knelson concentrators

was then transferred to a 1 m3 storage hopper in the gold room after regular intervals depending on Knelson

concentrator cycle times. On a daily basis, the gravity concentrate was batch fed, using an auger feeder, to a

40 inch Mineral Cone for secondary gravity concentration. The concentrate was upgraded by the rejection of

low specific gravity materials to produce a smaller volume of concentrate at a higher gold grade.

The final concentrate was filtered using a filter press to reduce moisture content prior to being calcined in an

oven at 650oC for around 12 hours. The calcine was then removed from the oven, weighed and stored for

smelting. Figure 3 shows the flow sheet of the gravity circuit prior to the upgrade.


Figure 3: Gravity circuit flow sheet prior to upgrade

The gold recovery from the gravity circuit averaged 17% over the 18 month period (Oct-08 to Mar-10) prior to

the gravity circuit upgrade. The gravity recovery was shown to be dependent on the ore sources being

processed. As proportions of MLU ore in the blend fluctuated and changes in open pit ore bodies being

processed, it was shown that the gravity recovery also fluctuated. An example of this is shown in Table 1

which shows reconciled monthly gravity and leach recoveries in the 9 month prior to the gravity circuit

upgrade.

Jul-09 Aug-09 Sep-09 Oct-09 Nov-09 Dec-09 Jan-10 Feb-10 Mar-10

Leach Recovery (%) 63.3 71.1 68.1 65.5 65.3 61.9 67.3 66.8 64.9

Gravity Recovery (%) 23.9 10.5 16.5 13.9 14.3 24.9 20.8 13.2 18.9

Overall Recovery (%) 87.3 81.6 84.6 79.4 79.6 86.8 88.1 80.0 83.8

Table 1: Plant Recoveries

Issues arising from the gravity circuit

Gravity Recovery Gravity recovery averaged 18.9% of the feed gold to the processing plant for the 9 month period Jul-09 to

Mar-10. Gravity recoverable gold (GRG) testwork on the ore sources being treated indicated that higher

gravity recovery should have be achieved through the processing plant. The predicted gravity recovery

versus actual recoveries is shown in Figure 4.

Smelting

Mineral Cone Tailings

Calcining

Mineral Cone

Final Concentrate

Cyclone Feed Hopper

Magnetic Separator

Knelson Concentrator

Concentrate Hopper

Gravity C t t

Magnetic Material

Gravity Feed

Gravity Screens

Cyclone Underflow

Iron Reduced Gravity Feed

OversizeSecondary Mills

Knelson Tailings


0%

10%

20%

30%

40%

50%

60%

70%

80%

90%

100%

Recovery

Month

Gravity Recovery Overall Recovery Predicted Gravity Recovery

Figure 4: Predicted gravity recovery versus actual gravity and overall recovery

The cyclone underflow assay during this period averaged 30g/t indicating that there was a high probability of

coarse gold that could be recovered through the gravity circuit. The overall plant recovery averaged 83.8%

which was significantly lower than historical plant gold recoveries of around 86%. This was a direct result of

the proportion of Mercury ore in the mill feed blend.

Mineral Cone Performance Significant gold losses were experienced during operation of the Mineral Cone. Gold particles that were too

small, uneven shaped or locked in gangue material were lost from the Mineral Cone through the Mineral

Cone tailings which reported back to the grinding circuit. The performance of the Mineral Cone was strongly

influenced by the gold room operator with several different settings which could be adjusted by the operator.

These process variables included: water flow, water pressure, Mineral Cone angle, and feed flow rate to the

Mineral Cone. All of these variables required re-adjustment whilst operating the Mineral Cone which

depended on the quality of the concentrate being fed onto the Mineral Cone and therefore could not be

unconditionally set. Furthermore, each gold room operator had slightly different set points and operating

techniques depending on their individual experience and knowledge of the Mineral Cone which resulted in

fluctuations in gold recovered from the Mineral Cone.

The low recovery of the Mineral Cone was observed in the 2-hourly leach feed samples. There was a

significant increase in the leach feed assay whilst the Mineral Cone was operating as shown in Figure 5. This


data is the average leach feed assay for August 2009 indicating an extra 16% of gold reporting to the

leaching circuit that should be recovered in the gravity circuit.

0.00

0.50

1.00

1.50

2.00

2.50

3.00

3.50

0500

0700

0900

1100

1300

1500

1700

1900

2100

2300

0100

0300

Lea

ch F

eed

Ass

ay (

g/t

)

Time

Spike in Leach feed assay indicating poor Mineral Cone performace

Figure 5: Two-Hourly average leach feed assay in August 2009

Occupational Health and Safety Sulphide minerals such as Arsenopyrite, Pyrite and Pyrrhotite are present at varying concentrations in the

majority of ore sources processed at Marvel Loch. Due to their higher specific gravity, sulphide minerals are

recovered in the Knelson concentrators. As it was not possible to remove all sulphide minerals when

processing the Knelson concentrate through the Mineral Cone, a significant amount of sulphide minerals

reported to the final gravity gold concentrate. This can be seen in Figure 6. Steel chips and other magnetic

material not removed by the magnetic separator would also be recovered in the Knelson concentrate and the

Mineral Cone concentrate as shown in Figure 7.

The presence of sulphide minerals, particularly Arsenopyrtite, in the gravity concentrate made smelting of the

concentrate problematic and resulted in contaminated gold doré bars. To overcome this issue, the

concentrate was oxidised by calcining prior to smelting. During the calcination process, toxic fumes

containing sulphur dioxide and arsenic trioxide were generated. The majority of these toxic fumes were

removed through the gold room scrubber system but the risk of exposure to plant personnel could not be

completely eliminated. Furthermore, the elevated levels of arsenic in the gold room and surrounding process

plant areas also depended on the gold room scrubber system operating efficiently.


Figure 6: Gravity cone concentrates containing free

gold and sulphides

Figure 7: Gravity cone concentrates containing free

gold and steel

The presence of sulphides, nickel (from grinding media) and arsenic caused a high concentration of

contaminants in the gold doré bars produced at Marvel Loch. Different flux mixtures were tested but the level

of arsenic and nickel contaminants remained above the maximum levels permitted by the Perth Mint gold

refinery and this resulted financial penalties imposed on St. Barbara Ltd.

During operation of the Mineral Cone, there was a risk of injury from exposure to rotating parts because the

Mineral Cone was not fully enclosed or guarded. Manual adjustment of feed water, spray water and rotating

angle were regularly required during operation to ensure optimum performance.

Significant manual handling of gold concentrate was required during filter pressing and handling the

concentrate into and out of the calcine oven. There were numerous occurrences of concentrate spillage

which resulted in a safety hazard as well as a potential loss of gold concentrate.

Security As gold concentrate was filtered and manually transferred to the calcine oven there was an increased risk of

theft of concentrate which typically contained a significant proportion of free gold. Industry standard gold

room procedures were in place, such as the two person policy and CCTV security monitoring, to help

mitigate the risk of gold theft. However, these procedures could not completely remove this security risk.

Metallurgical Accounting Daily estimation of gravity gold recovery was difficult resulting in errors in the daily reporting of grades,

recoveries and ounces produced. These could be reconciled once mint returns were received but this left it

too late to identify fluctuations in mill head grade, plant performance or possible gold theft. In an attempt to

alleviate this issue, several methods of estimating the percentage of gold in the concentrate were tested, but

none were deemed effective for metallurgical accounting purposes.


Time Management The length of time to operate the Mineral Cone on a daily basis was significant. For instance, it could take up

to eight hours for all of the Knelson concentrate to be fed into the Mineral Cone and longer if there were any

operational or maintenance issues. As a consequence, the gold room operators who also operated the

elution circuit and other sections of the gravity circuit, had less time to manage those areas. This led to

increased downtime and below average performance in these areas.

NEW CIRCUIT TESTWORK, EQUIPMENT AND DESIGN A number of studies have been completed on the gravity circuit at Marvel Loch. AMIRA project P420B

conducted by Laplante in September 2003 indicated that by replacing the Mineral Cone with an intensive

cyanidation process, a gravity recovery of 45% could be achieved.

A further gravity circuit review in April 2007 indicated deficiencies in optimising Knelson concentrator feed

flow rate and the performance of the Mineral Cone. Test work was performed on the Mineral Cone tailings

indicating that the Mineral Cone recovery was between 60-70%.

Further modelling and simulation work completed on the gravity circuit showed that increasing the mass

recovery from the Knelson concentrators would increase the gravity recovery from 17% to 21% and

replacing the Mineral Cone with intensive cyanidation would further increase the gravity recovery above

25%.

Intensive Cyanidation The proposed upgrade was to replace the Mineral Cone with an intensive cyanidation unit. Intensive

cyanidation involved the use of relatively high concentration cyanide solution in favourable leaching

conditions to leach gold from a gravity concentrate to produce a gold cyanide solution suitable for

electrowinning.

Arsenopyrite is not dissolved by cyanide removing the requirement for calcining prior to smelting thus

removing the exposure of personnel to sulphur dioxide and arsenic trioxide fumes. Intensive cyanidation

units are fully automated, removing manual handling and physical contact with gold concentrate and thereby

improving safety and security.

Two units were identified as possible secondary treatment options: the Gekko InLine Leach Reactor (ILR)

and the Consep Acacia Reactor.

Gekko ILR The Gekko ILR was trialled by the previous mine owner - Sons of Gwalia in 2003 as a means of improving

gravity recovery. The Gecko ILR consisted of a concentrate hopper to collect and dewater concentrates from

the Knelson concentrators. The main reactor works on the principal of a laboratory bottle roll, with a

horizontal drum rotating at low speeds with baffles and aeration inside the drum to maximise leaching

performance. The solids remain inside the drum while the solution is circulated from a holding tank through


the drum and back to the holding tank. The barren solids are removed back to the grinding circuit and the

pregnant electrolyte is filtered and gold recovered in the electrowinning circuit.

A trial of an ILR 100 was conducted onsite for a number of tests on the Mineral Cone tailings and Knelson

concentrate batches. The results from the test work are summarised in Table 2.

Test no. 1 2 3 4 5 6 7 8 9 10 Average

*

ILR Feed (g/t) 2126 2897 3216 1626 2669 2263 6406 2897 3216 1626 2971

ILR Tails (g/t) 1134 764 111 141 763 356 366 764 111 141 398

ILR Recovery (%) 46.7 73.6 96.5 91.3 71.4 84.3 94.3 73.6 96.5 91.3 86.6

Table 2: Summary of ILR Trial Results (* tests 2-10 only)

The gravity recovery from the Mineral Cone tail was shown to be 46.7% (Test 1). The average ILR recovery

of the Knelson concentrate was 86.6% (tests 2 to 10). ILR recoveries ranged between 73.6% and 96.5%

from the Knelson concentrate. The highest recovery was achieved on one particular sample following 22

hours of leach time. The results from this trial showed how poorly the Mineral Cone was performing and how

intensive cyanidation could significantly increase the recovery.

Although the ILR demonstrated a significant improvement in gravity recovery, the capital request for

installation of an ILR 2000 was not approved at the time of the trial in 2003.

Consep Acacia Reactor The Consep Acacia Reactor was developed by Anglogold Australia at Union Reefs gold plant in 1998

(Watson and Steward, 2002). The main component of the Acacia Reactor is the fluidised bed reactor where

the leaching of gold occurs. It has the same concentrate storage hopper and similar electrolyte storage tank

to the ILR. The Knelson concentrate is transferred from the storage hopper into the fluidised bed reactor,

where electrolyte solution is pumped up through the bottom of the reactor to fluidise the solids and promote

mixing of the electrolyte and concentrate. Once the leaching cycle is finished the barrens solids are rinsed

and discharged back to the milling circuit and the pregnant electrolyte is transferred to the electrowinning

circuit.

During a gravity circuit review by Consep it was determined that the current gravity circuit was operating

below optimum performance. Inefficient secondary treatment of the gravity concentrate was seen as a major

contributor to the mediocre performance. Following the audit report a survey and trial of a pilot Acacia

reactor was conducted.

The pilot Acacia Reactor CS50 trial was conducted to examine the amenability of Knelson concentrates to

the intensive cyanidation process. The first test was conducted on tailings from the Mineral Cone to

determine the quantity of recoverable gold being recycled back to the leach circuit. Subsequent tests were


conducted on samples of untreated Knelson concentrates. Table 3 summaries the findings from the pilot

Acacia Reactor test work (Wren, 2008).

Mineral

Cone

Acacia Reactor

Test 1 Test 2 Test 3 Test 4

Acacia Feed (g/t) 8061 2567 8061 6194 5857

Acacia Tails (g/t) 2567 4.37 40.64 15.7 20.3

Recovery (%) 68 89 to 99.8 99 to 99.5 98 to 99.7 99.7

Operating time (hr) 3 5 to16 2 to16 4 to 24 23

Table 3: Summary of Acacia Pilot Plant Trial

The test work demonstrated that Mineral Cone’s gold recovery was 68% compared to 89% – 99% recovery

from the Acacia Reactor. The lowest Acacia Reactor recovery (Test 1) was 89% which was achieved using

the tailings from mineral cone after five hours of intensive cyanidation. Test 2, 3 and 4 were conducted on

the Knelson concentrates and achieved significant leach recovery results.

New Circuit Flow Sheet With the promising results of the intensive cyanidation trials, a gravity circuit flow sheet was developed to

replace the Mineral Cone with intensive cyanidation. Due to the financial constraints of adding an extra

electrowinning module, Marvel Loch processing management decided to use the existing electrowinning

circuit on site to recover the gold from the pregnant electrolyte, which introduced a number of new issues.

Firstly, gold doré from gravity production was not separated from leach circuit production. Instead, the

pregnant liquor assay and volume was used to calculate the gravity production making gravity recovery

calculations reliant on aqua regia assays rather than on bullion assays. Secondly, as the intensive cyanide

operation relied on the existing electrowinning operation and combined the pregnant liquors, it increased the

gold concentration thus requiring longer electrowinning times and as the previous electrowinning cycle

dictates the elution timing, there was a risk of delaying the next elution cycle.

These issues could be reduced by cleaning out the electrowinning cells more frequently, increasing the

efficiency of the electrowinning process and reducing the electrowinning time. The final gravity circuit flow

sheet is shown in Figure 8.


Smelting

Acacia Tails

Acacia Reactor

Cyclone Feed Hopper

Magnetic Separator

Knelson Concentrator

Gravity Concentrates

Magnetic Material

Gravity Feed

Gravity Screens

Cyclone Underflow

Iron Reduced Gravity Feed

OversizeSecondary Mills

Knelson Tails

Electrowinning

Electrolyte

Figure 8: Final gravity circuit flow sheet

Selection of Intensive Cyanidation Unit There are some significant differences between the two intensive cyanidation units that were tested. The

capital cost of the two units were similar leading us to investigate the unit characteristics and operating costs

of the two units to determine the best option. These are summarised in Table 4 below.

Gekko ILR Consep Acacia Reactor Pros Lower operating cost Less installed power All gold particle sizes treated Ability to be upgraded to larger batch unit

Pros Higher Recovery Robust leaching chemistry Simple de-sliming system Heating provides additional leaching flexibility No moving parts

Cons Produces more pregnant liquor Clarifying pregnant liquor more difficult Larger footprint Rotating equipment needs guarding Higher maintenance costs Higher risk of downtime

Cons Fluidised bed not flexible to feed variations System not expandable Gold fines lost in de-sliming process Leach reagents can effect electrowinning

Table 4: Pros and Cons for Intensive Cyanidation Units

The decision made by Marvel Loch processing management was to choose the Acacia Reactor due to its

higher gold recovery, simplistic design and less moving parts compared to the Gekko ILR.


OTHER GRAVITY CIRCUIT UPGRADES There were a number of other issues identified in the gravity circuit which required further improvement.

These included the inefficiency of the gravity feed screens, electrowinning and acid digestion with mild steel

wool, and relatively low availability of the gravity circuit resulting in poor gravity circuit performance.

Gravity Screen Changes

The gravity screens that fed the gravity circuit were proved to be inefficient. There was a significant portion of

undersize reporting back to the grinding circuit thereby reducing the flow to the gravity circuit and causing the

Knelson concentrators to be under fed. The screen panels were made from hard polyurethane that would

easily peg causing a majority of the feed that fed the screens to just pass over the top of the screens. Plant

operators would have to pressure clean and de-peg the screens daily to ensure adequate efficiency but

would become pegged again after a few hours of operation. The size of the screen panels (1480 x 600 mm)

were heavy and cumbersome, which created manual handling issues during change outs. Due to the

continuous pegging and scale build up, the screen panels had to be changed out every three months.

Due to the efficiency and manual handling issues with the gravity feed screens, different panels were

investigated to improve the efficiency and reduce the maintenance to the screen panels. A softer

polyurethane panel (2.5 x 18 mm slotted) was supplied having increased elasticity that would not allow rocky

material to peg up the screens. As the screens were clear there was no scale build up below the screen so

the efficiency of the screen was very high, as can be seen in Figure 9. The size of the screens was

significantly smaller (484 x 305 mm), allowing easier installation and only requiring replacement of a few

panels when worn which reduced the operating costs of the gravity screens. The new screen panels were

installed in mid-2009 for a capital cost of $12,000 and has resulted in a significant improvement to the gravity

screening efficiency.


0

10

20

30

40

50

60

70

80

90

100

100 1,000 10,000

Per

cen

t Pas

sin

g (%

)

Size (um)

Underflow Overflow Feed

Figure 9: Gravity screen efficiency

Gold Room Modifications

During the process of upgrading the gravity circuit, upgrades to the gold room processes also occurred. The

main upgrade involved converting the mild steel wool cathodes to stainless steel mesh cathodes.

The previous process using the mild steel wool cathodes involved operators pulling mild steel cathodes

contained inside baskets. The baskets would be broken down and the loaded steel mesh removed and

placed into large bins. Hydrochloric acid (30%) would be added to the bins to dissolve the steel wool. During

the acid digest process the gold room would have to be evacuated resulting in operators being unable to

continue other tasks, such as operating the Mineral Cone. There was a risk of exposure to hydrochloric gas

and other fumes and there was also a risk of fire due to the heat generated from the acid digest process.

The introduction of stainless steel mesh cathodes removed the requirement for acid digest in the gold room.

There was no longer a requirement to have baskets with the wool inside, just a barrier to prevent the

cathodes and anodes from touching. Modifying this process reduced the time that was required to remove

gold from the electrowinning cells. As a result, the gold room operators could pull cathodes more often,

increasing the efficiency of the circuit and reducing the electrowinning times.

Knelson Concentrator Spares A major issue with the gravity circuit was the availability of the two Knelson concentrators. If one

concentrator was offline there was a significant increase in the circulating load of gravity gold through the

secondary grinding circuit which would result in extra feed to the leaching circuit, reducing the gravity

recovery. Significant time was spent collating all of the spare Knelson concentrator parts and ordering critical


spares that were not stocked on-site. This resulted in a significant improvement in the Knelson concentrator

availability which consequently reduced the cyclone underflow and leach feed assays.

GRAVITY CIRCUIT AND ACACIA REACTOR OPERATIONS

Installation & Commissioning of the Acacia Reactor

The Marvel Loch processing management decided to install the Acacia Reactor as part of the existing gold

room to minimise the effect the installation had on normal gold room operations and to reduce the possibility

of gold theft. Modification to the gold room and supply of services occurred over several months, with the

final steps being the installation of the Acacia Reactor and rebuilding the gold room surrounding the Acacia

Reactor.

Commissioning of the Acacia Reactor occurred in April 2010 and went very smoothly with only a few minor

issues arising. The main issue was rearranging the program sequence to allow the discharge of the barren

solids before the discharge of the pregnant electrolyte due to the constraints of the current elution setup and

the impact of campaign milling. Another issue was rearranging the direction of the sampling valve to prevent

bypassing during normal operations.

During the first four days of the Acacia Reactor operation, one Knelson concentrator was directed to the

Mineral Cone and the second Knelson concentrator was directed to the Acacia Reactor. During that trial

period gold recovered from the Acacia Reactor was very promising. Table 5 compares the gravity gold

recovered from the Mineral Cone and the Acacia Reactor.

9 April 10 April 11 April 12 April

Gold recovered from Mineral Cone (g) 1 984 1 404 1 212 1 429

Gold recovered from Acacia Reactor (g) 5 440 6 396 5808 6 092

Gold in plant feed (g) 21 553 1 831 17189 16 491

Table 5: Gold recovered from the Acacia Reactor

After four days of running both secondary gravity recovery options, the second Knelson concentrator was

redirected to the Acacia Reactor. This resulted in a significant increase in gold recovered with the Acacia

Reactor in the circuit. Figure 10 shows the significant difference in the gold recovered through the gravity

circuit from before the change, during commissioning and at full production. Gravity recovery increased from

around 20% to over 50%.


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%)

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Mineral Cone Acacia Reactor

4-day peiod where Mineral Cone and Acacia Reactor were operated in parallel during Acacia commissioning

Figure 10: Increase in gravity gold recovery from the Acacia Reactor

Recovery figures for the barren solids samples taken from the Acacia Reactor average 96%. There are

issues with taking an accurate sample of the Acacia Reactor barren solids due to the nugget effect resulting

in some large variances in barren solids assay but overall it shows a very good performance as shown in

Table 6.

Acacia Batch No. 153 154 155 156 157 158

Feed Grade (g/t) 4 969 3 884 3 638 4 933 3 202 3 285

Tails Grade (g/t) 238 43 60 401 23 183

Recovery (%) 95.2 98.9 98.4 91.9 99.3 94.4

Table 6: Typical Acacia Reactor Recoveries

Figure 11 shows the leaching profile during an Acacia Reactor batch. It demonstrates the favourable

leaching conditions created by the Acacia Reactor producing fast leach kinetics. This makes sampling during

the batch reliable in giving a final estimation of the gold recovered allowing a daily grade and recovery figure

to be estimated before the Acacia Reactor batch is complete.


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Figure 11: Leaching profile of an Acacia Reactor batch

Overall Effect on Gravity Circuit

Since the installation of the Acacia Reactor there has been a reduction in the mass of barren gravity

concentrate returning to the grinding and leaching circuits. There has also been an increase in gold recovery

through the gravity circuit as shown in Table 7.

Jan-10 Feb-10 Mar-10 Apr-10 May-10 Jun-10 Jul-10

Leach recovery (%) 67.3 66.7 64.8 50.1 42.9 44.5 42.8

Gravity recovery (%) 20.8 13.2 18.9 41.5 49.7 50.3 52.0

Overall recovery (%) 88.1 79.9 83.7 91.6 92.6 94.8 94.8

Table 7: Increase in Gravity Recovery

The increase in overall recovery partially coincided with Mercury ore being removed from the mill feed blend.

The blend percentage of Mercury ore in the blend was 10% over the first two months the Acacia Reactor

was in production (April and May 2010). The increase in gravity gold production over these two months is

significantly higher than the changes in Mercury ore blend percentage, indicating that the Acacia Reactor

had significant impact on the overall plant recovery.

Figure 12 shows the increase in overall plant recovery from 84% up to 92% once the Acacia Reactor was

installed. The gravity recovery increased from 17% to over 45%. Reviewing recent plant data, the overall

plant gold recovery is 86.3% indicating the improvement to recovery as a result of the Acacia Reactor is

between three and five percent.


The GRG test work on the ore sources being treated indicated that the gravity recovery with the upgraded

gravity circuit was similar to that predicted by the test work.

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Before UpgradeAfter Upgrade

Figure 12: Overall plant and gravity recoveries

There has been a significant decrease in the variability of the assay in the cyclone underflow stream. Prior to

the upgrade the cyclone under flow averaged 30 g/t, after the upgrade it decreased to around 18 g/t. This

represents a significant amount of gold being removed from the circulating load of the secondary grinding

circuit. Figure 13 shows the reduction in cyclone underflow assay before and after the upgrade.

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Figure 13: Cyclone underflow assay before and after upgrade

There has also been a significant decrease in the amount of gold reporting to the leaching circuit, with a

decrease in the gold losses to tails as shown in Figure 14. The tails grade has reduced from 0.40g/t to

0.24g/t after the upgrade. Some of the reduction of tails grade may be attributed to the removal of Mercury

ore from the blend. For one month either side of the upgrade, the tails grade dropped from 0.32g/t to 0.23g/t

indicating there was a significant reduction due the upgrade of the gravity circuit.

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Figure 14: Tails grade before and after upgrade

Impact on Leaching and Elution Circuits

Reagent Costs With the reduction of gold reporting to the leaching circuit there has been a reduction in reagent costs

(particularly in cyanide) required for leaching as shown in Figure 15. Removing the two outliers one from

before and one after the installation, cyanide consumption has reduced from 0.89 kg/t to 0.78 kg/t

representing an annual saving of around $310,000.


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Sep‐09 Oct‐09 Nov‐09 Dec‐09 Jan‐10 Feb‐10 Mar‐10 Apr‐10 May‐10 Jun‐10 Jul‐10 Aug‐10 Sep‐10 Oct‐10 Nov‐10 Dec‐10

Cyanide consumption (kg/t)

Month

Before upgrade

After upgrade

Figure 15: Cyanide consumption for leaching before and after upgrade

There has also been an operating cost reduction associated with consumables required for the elution

circuit. Costs of activated carbon, caustic soda and hydrochloric acid have all been reduced due to less gold

reporting to the leaching circuit as a result of less elution strips required after the gravity circuit upgrade.

Overall Operating Costs With the reduction of cyanide consumption and other operating costs there has been a significant reduction

in overall processing operating costs as shown in Figure 16. The plant operating cost before the upgrade of

the gravity circuit was $29.02 per tonne of ore treated compared with after the upgrade the operating cost

reduced to $27.39 per tonne representing an annual saving of $1.96 million. There have been other process

improvements to the Marvel Loch plant during this period resulting in lower operating costs so the reduction

in the processing costs cannot all be attributed to the Acacia Reactor alone.


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Sep‐09 Oct‐09 Nov‐09 Dec‐09 Jan‐10 Feb‐10 Mar‐10 Apr‐10 May‐10 Jun‐10 Jul‐10 Aug‐10 Sep‐10 Oct‐10 Nov‐10 Dec‐10

Processing Costs ($/t)

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Before upgradeAfter upgrade

Figure 16: Plant operating costs before and after upgrade

With the cost of the upgrades to the gravity circuit and the associated reduction in operating cost the

payback period for the installation is estimated to be between three and six months.

ASSOCIATED BENEFITS OF THE UPGRADE

Occupational Health and Safety

With the removal of the requirement to calcine gravity concentrate, no sulphur dioxide or arsenic trioxide

fumes can be produced in the gold room. The automated process of the Acacia Reactor has significantly

reduced manual handling requirements and spillages requiring physical labour to clean up. There has also

been a reduction in moving parts reducing the associated hazards and possible pinch points.

Security

Due to the reduction in free gold being handled by the gold room operators, there is no physical contact with

the gravity concentrate resulting in decreased opportunities for gold theft within the gold room.

Labour

Labour hours have been significantly reduced throughout the gravity circuit as a result of the above

mentioned gravity circuit upgrades. Prior to the gravity circuit upgrade, it took on average three to six hours

to process the Knelson concentrate over the Mineral Cone but with intensive cyanidation the labour hour

requirement is reduced to less than one hour. Adding in the improvements to the maintenance of the gravity


screens, reduction in time to replenish the electrowinning cells and other process improvements, this allows

the gold room operators more time to concentrate on optimising other sections of the process plant.

Metallurgical Accounting

The benefits to the metallurgical accounting system onsite have been quite significant. There is now

estimation of gold recovered through the gravity circuit once each Acacia Reactor batch has been

completed. An initial sample is taken daily four hours after start up to give an estimation as to what the final

Acacia Reactor assay might be to allow metallurgical performance to be progressively monitored and the

final grade and recovery calculations follow once that batch has finished and assay is confirmed. Having an

assayed value and volume reading has allowed increased precision for gold in circuit calculations and this

makes the daily reporting and weekly reconciliation process more accurate. Calculation of total gravity

recovery is easier with a definite amount from the gravity circuit, although there is no reconciliation from gold

doré concentrations from the gravity circuit.

FURTHER IMPROVEMENTS Further improvements to the gravity circuit will occur with further optimisation of the Knelson operating

parameters and optimisation of the gravity feed flow from the cyclone underflow. Determining size by size

gold recovery and modelling through the gravity circuits will indicate if further optimisation is possible and if

there are any major losses of fine gold to the leaching circuit.

Possible installation of a separate electrowinning cell and electrolyte tank to alleviate pressure and allow

further flexibility from the current elution and electrowinning capacities is considered. Consideration has also

been made to reduce the Acacia reactor cycle time from 24 hours to 12 hours (i.e. 2 cycles per day) to

process a larger mass of primary gravity concentrate to further increase gravity recovery and overall plant

gold recovery.

Another area identified for improvement is the removal of the magnetic separator from the gravity circuit.

Magnetic material currently recovered by the magnetic separator would be processed through the gravity

circuit instead of being recycled back to the grinding circuit. This will also reduce the maintenance and

operating costs of this unit.

CONCLUSION The benefits of good circuit design and continuing operational improvements can not only be shown to

reduce operating costs of a processing plant but result in other operational benefits such as removing

occupational health and safety hazards, improving metallurgical accounting, increase gold room security and

considerably reduce manual labour hours of plant operators to allow them to focus on obtaining overall

increased plant performance.

Specific benefits from the gravity circuit upgrade of the Marvel Loch processing plant include a significant

increase in gravity recovery, an increase in overall plant recovery and a reduction in plant operating costs.

Other benefits that have been realised through this project have been the removal of toxic fumes (sulphur


dioxide and arsenic trioxide) and hydrochloric acid in the gold room, an improvement in gold room security,

the removal of manual handling and equipment hazards, the removal of labour intensive tasks thus

improving overall efficiency of the gravity, elution and gold room circuits.

ACKNOWLEDGEMENTS The authors wish to thank the management of St. Barbara - Marvel Loch gold mine for their support and

permission to publish this paper.

The authors also acknowledge the team from Consep for their assistance during design and commissioning,

Mr. Augy Wilangkara (former Senior Metallurgist at Marvel Loch) for drafting the original proposal for the

Acacia Reactor purchase and installation, and the Marvel Loch processing team for all their efforts during the

recent upgrades of the gravity circuit.

REFERENCES Laplante, A, 2003. AMIRA Project 420B - Marvel Loch Site Visit Report. A.J. Parker Centre

St Barbara Limited 2008. Southern Cross, [online]. Available from <http://www.stbarbara.com.au/our-

operations/southern-cross/> [Accessed 15th February 2011]

Watson, B and Steward, G, 2002. Gravity Leaching With the ConSep ACACIA Reactor — Results From

AngloGold Union Reef. In Proceedings from Metallurgical Plant Design and Operating Strategies

Wren, D 2008. P2689 Testwork Report Consep Acacia CS50 Pilot Plant, Consep, Perth


8 - 9 August 2011 Perth, WA41

In-Line Pressure Jig Preconcentration Plant at the Pirquitas Mine

A H Gray1, G Delemontex2, N Grigg3 and T Yeomans4

1. Technical Director, Gekko Systems Pty Ltd, 321 Learmonth Road, Ballarat Vic 3350. Email: [email protected]

2. Senior Process Engineer, Gekko Systems Pty Ltd, 321 Learmonth Road, Ballarat Vic 3350. Email: [email protected]

3. MAusIMM, General Manager – South America, Gekko Systems SA, Augusto, Leguia Norte 100, Of. 613, Las Condes, Santiago, Chile. Email: [email protected]

4. Director of Metallurgy, Silver Standard Resources Inc, 999 West Hastings Street, Suite 1400, Vancouver BC, V6C 2W2, Canada. Email: [email protected]

ABSTRACT Gekko Systems Pty Ltd engaged in the test work, design, manufacturing and commissioning of the pre-

concentration plant at Silver Standard Resources Inc’s Pirquitas mine in the north western Jujuy province of

Argentina. Testing of the ore at coarse feeds of between 2 and 12mm returned probable recoveries of up to

a total of 95% silver into 50% of the mass fed to the InLine Pressure Jigs (IPJ’s). The plant has been

designed and built based on the test results to include a preparation screen to bypass the fines that are

naturally higher in grade and repulping of the coarse which is then pumped to three parallel trains of 2-stage

IPJ roughing-scavenging. The target 50% yield to concentrate passes over a dewatering screen and the

solids are returned back to the mill feed conveyor. The IPJ tails are transferred to a dewatering screen and

solid rejects are stockpiled off the end of a transfer conveyor. All dewatered underflow products are

processed via hydrocyclones to recover dirty water for use back into the IPJ circuit and return entrained fines

to the fines bypass line.

With the sulphide orebody being processed, the results from operations to date have shown that plant

performance is matching the original laboratory sulphide ore testwork data. The objective of the pre-

concentration step is to both reject gangue and achieve grades between 300-400g/t Ag to the float circuit as

flotation recoveries are maximised within this range. The plant currently achieves overall upgrades of

approximately 150% of feed grade with the IPJ only circuit producing an average 180% increase from

calculated feed grade.

INTRODUCTION Gekko Systems Pty Ltd (Gekko) were requested by Silver Standard Resources Inc (Silver Standard) in 2006

to investigate the use of the InLine Pressure Jig (IPJ) to pre-concentrate the silver, zinc, tin ore from their

Pirquitas deposit. Silver Standard had previously carried out jigging test work to pre-concentrate the ore and

found it to be very successful (Hatch, 2006).

The following sections of this paper describe the test work program, results, plant design, manufacture and

commissioning of an InLine Pressure Jig based pre-concentration plant. Recent plant performance data is

compared to the test work data.



BACKGROUND

Pirquitas Mine Silver Standard Resources Inc (Silver Standard) acquired the Pirquitas mine in October 2004 when it

completed the purchase of 100% of Sunshine Argentina Inc.

The Pirquitas mine is a silver, zinc and tin deposit located in the extreme north west of Argentina in the

province of Jujuy as shown in Fig. 1. The project is located in mountainous terrain at altitudes ranging

between 4000m to 4520m above sea level.

Fig. 1: Location of Pirquitas Project (Google Maps)

According to the Feasibility study (Hatch, 2006) eighty to ninety percent of the potentially economic material

is made up of siliceous non-sulfide gangue composed of intergrown quartz and feldspar grains in a

micaceous matrix. Five to twenty percent of the mineralization is made up of sulfides, which is mostly pyrite,

and is commonly associated with silver, zinc and tin. From mineralogical studies, the silver was reported as

being present primarily as silver sulphides and sulphosalts, containing high concentrations of silver, with

most of the silver mineral grains in the range of 20 to 100µm in size. The zinc occurs primarily as medium to

coarse grained sphalerite. Grain sizes up to 800µm were reported with 30 to 150µm the more typical size

range. Tin occurs as aggregates of cassiterite crystals ranging from 5 to 15 μm, often intergrown with pyrite.



In 1997 and 1998 gravity concentration test work was carried out and indicated pre-concentrating the

Pirquitas ore by jigging, after crushing to ½ inch was expected to recover approximately 50% of the crushed

ore as mill feed, while only losing three to five percent of the silver and tin.

The benefits of pre-concentration to the Pirquitas project included keeping equipment smaller due to the

issues with construction and operation at high altitudes, avoidance of smelter penalties when lower mine

grades are processed and more metal production from the plant.

After the 2006 feasibility study by Hatch Engineering (Hatch, 2006), it was expected the ore would be

processed by sequential crushing, screening and pre-concentration jigging (gravity). The jig concentrate and

fines would be treated by grinding, selective silver flotation and sulphide/zinc flotation from which the zinc

can be floated to a saleable grade concentrate. The sulphide/zinc tailings are then treated by gravity and

flotation circuits for tin recovery to a saleable grade tin concentrate (Fig. 2).

Fig. 2: Overall Pirquitas Flowsheet (MacRae and McCrea, 2008)

Pirquitas Proven and Probable mineral reserves (MacRae and McCrea, 2008) were estimated in house to be

30.4 million tonnes grading 199.6g Ag/t, 0.82% Zn and 0.22% Sn including the historical jig tailings. At a

processing (crushing) rate of 6000tpd, planned mine life is 15 years.

Silver Standard commissioned the milling/flotation circuit on December 1st, 2009 at 4000tpd. The

commissioning of the pre-concentration IPJ circuit in the 3rd quarter 2010 has resulted in an increase in

processing rate to 6000tpd. Forecast (Silver Standard Resources Inc, 2011) silver production for 2011 is 8.5

million ounces of silver.



InLine Pressure Jig The InLine Pressure Jig (IPJ) has been described in numerous papers (Gray and Hughes, 2008) and a brief

summary follows:

The IPJ is unique in its design and use of jigging concepts in that the unit combines a circular bed

with a moveable sieve action.

The screen is pulsed vertically by a hydraulically driven shaft with the length of the stroke and speed

of the up and down stroke varied to suit the application.

Screen aperture, ragging dimension and ragging material can also be altered for the application.

Inside the IPJ, the particles are kept submerged in the slurry thus eliminating the loss of hydrophobic

fine particles at the air/slurry interface of conventional jigs.

The submerged slurry also acts as a pseudo heavy media suspension above the jig bed greatly

assisting the separation performance of the IPJ.

Separation of valuable minerals from gangue particles occurs based on relative density as well as

particle size and shape.

High specific gravity particles are drawn into the concentrate hutch during the suction stroke of the

bed and are continuously discharged while the lighter gangue is continuously discharged over the

tailboard to the outer cone.

An overview is shown in Fig. 3.

Fig. 3: IPJ cross sectional view.



TEST WORK PROGRAM AND RESULTS

Extensive gravity characterisation testwork was conducted at Gekko’s laboratory in Ballarat. The test work

flow sheet is presented in Fig. 4. Key tests conducted during the program include size distribution analyses,

single pass tabling tests of a +1.7mm -3.35mm size fraction and dense media “Viking Cone” testing of

+3.35mm -13.32mm and +3.35mm -9.5mm size fractions.

* All assays to include: Ag and Zn (ICP) + Sn (XRF) as per previous testwork

Level #40, Level #26 & Level #17 Samples

Each samples placed into low

temperature (50oC) oven

Each sample prepared to 100% Passing 13.2 mm using lab jaw

crusher

Each sample Screened at 3.35 and 1.7 mm

-1.7 mm fraction dried weighed. Sample sent for assay

+ 1.7 mm to - 3.35 mm fraction +3.35 mm to -13.2 mm fraction

Single Pass Tabling Test Dense Media Viking Cone Test

Sample split out for size distribution and analysis.

Reserve +3.35 mm to -13.2 mm fraction

Approx 36 kg of each sample was split out and prepared to

100% passing 9.5 mm using lab jaw crusher

Each sample Screened at 3.35 and 1.7 mm

-1.7 mm materialDry, weigh and assay

+3.35 mm to -9.5 mm material

Dense Media Viking Cone Test

+ 1.7 mm to - 3.35 mm fraction

Single Pass Tabling Test

Fig. 4: Flowsheet for Pirquitas continuous gravity recovery test work



Three sulphide ore samples (#40, #26 and #17) of varying grades (Table 1) were tested for Ag, Zn and Sn

recovery using Gekko’s Viking Cone dense media separator. The Viking cone is similar to an Ericsson cone

and uses a dynamic process to determine the sink float characteristics of an ore. The test is particularly

suited to determining the recovery and mass yield that can be obtained using the IPJ at coarse sizes. The

mass yield to concentrate from these tests is used to specify the required mass pull from the IPJ.

Table 1: Pirquitas test work sample characteristics (concentrations in ppm)

Sample #40 #26 #17

Location Level 5, ¾” crush Level 2, coarse rock Level 2, ¾” crush

Silver 250 157 337

Zinc 404 754 2909

Tin 13326 5592 12209

Sulphur 70638 75258 75413

The test results for the +3.35 and -13.5mm secondary crushed material were particularly good. Both silver

and tin reported to the concentrate at mass yields less than 50%.

The results for Ag and Sn recovery in Figs. 5 and 6, show that recoveries in excess of 95% are achievable

for all three samples, at a mass yield of approximately 50%. Zinc recovery was lower as show in Fig. 7

indicating poorer liberation of the zinc minerals at the crush size tested.

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Fig. 5: Silver recovery versus mass yield from Gekko test work



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Fig. 6: Tin recovery versus mass yield from Gekko test work

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Fig. 7: Zinc recovery versus mass yield from Gekko test work

The test work also showed a natural concentration of silver in the fines (-1.7mm) after crushing to minus

12mm with its grade typically twice that of the head grade.

Overall the test work indicated treating plus 1.7mm minus 12mm will result in 95% recovery of silver into 50%

mass.



PLANT DESIGN

The main design parameters for the plant after communication with Silver Standard were 95% recovery of

silver and tin into 50% of the mass from the +2mm to -12mm stream. Throughput to the IPJ’s was specified

as 214 tph.

The recovery needed to be maximised and mass pull to concentrate was too high for a single pass jigging

stage so a two stage rougher-scavenger jigging circuit was used. This type of circuit was used very

successfully at the former Lihir Gold Ballarat Goldfields Project in Australia (Gray and Hughes, 2007). The

final design consisted of two stages of three IPJ2400 model InLine Pressure Jigs connected in parallel in a

rougher scavenger configuration as indicated in the Process Flow Diagram shown in Fig. 8. The throughput

rate for the IPJ2400 was de-rated from a nominal 100tph to 75tph due to the high mass pull to concentrate

required in this application and resulted in the use of three parallel trains of IPJ’s.

Fig. 8: Simplified PFD of Gekko pre-concentration process at Pirquitas

Crushed run-of-mine ore is fed to an existing fine ore screen. A diversion gate in the feed chute to the screen

allows the feed to be bypassed around the pre-concentration process. The fine ore screen splits the feed at

2mm, producing material in the size range, +2mm to -12mm, as feed to the Gekko plant, at a rate of 214 tph.

The dry feed is mixed with water recycled from the IPJ products to approximately 50% solids and is pumped

to the Rougher InLine Pressure Jigs (Stream 1 – Fig. 8).

The IPJ Feed pump (Warman 6/6 gravel pump), feeds a distributor vessel on top of the IPJ platform. The

pump’s speed is controlled by a variable speed drive to ensure that the pipe velocity of the slurry is greater

than the settling velocities of the coarse particles (>4 m/s).

The IPJs use an internal screen with 18mm apertures and the ragging used is 30mm in diameter and has a

specific gravity (SG) of 3.2. The large ragging size was selected to allow the jigging bed to be open and free,

which is necessary for the high mass pull.



The tails from the rougher IPJs (Stream 3) report to the in-series scavenger IPJs. The scavenger IPJs tails

(Stream 5) flow to the tails dewatering screen. The dewatered tail particles (+1 mm) pass to the existing

conveyor, while the water and fine solids pass to a combined screen undersize/cyclone feed sump. The

concentrates from the rougher (Stream 2) and scavenger IPJs combined (Stream 4) flow to the concentrate

dewatering screen and the water and fine solids also pass to the combined screen undersize/cyclone feed

sump.

The water and minus 1mm particles are pumped to a cyclone cluster designed to remove any liberated fines

from the recycled water. The cycloned water overflow is sent back to the gravity circuit’s 40m3 dirty water

tank for re-use in the IPJ’s. The cyclone underflow is directed to a sump along with a bleed of dirty water to

reduce the build up of ultra-fines in the circuit. A pump transfers this fine material directly to the ball mill

discharge hopper where it is then pumped to the cyclones in this grinding circuit.

The de-watered IPJ concentrates report to the ball mill feed conveyor.

The de-watered IPJ tailings report to the rejects stockpile for disposal.

The scope of Gekko’s supply for the project was expanded from just the IPJ supply to frames, feed pump,

product and tailings dewatering screens, water recycle circuit, PLC control and SCADA system and motor

control centres. This enabled Gekko to control the way the circuit was engineered to ensure it matched the

requirements of the IPJ’s and allowed Silver Standard to have one supplier answerable for the performance

of the circuit from pumping to de-watering. It also allowed the circuit to be installed independently of the

main grinding circuit.

The overall pre-concentration plant layout as modelled by Pro-E 3D modelling software is shown in Fig. 9.

Fig. 9: 3D Model of Pirquitas two stage gravity concentration plant



PLANT MANUFACTURE AND INSTALLATION The plant was fabricated, trial assembled and factory commissioned at the Gekko’s factory in Ballarat,

Australia in 26 weeks. Upon successful completion of factory commissioning (see Fig. 10), the plant was

marked, disassembled, packed into containers and shipped to site.

Fig. 10: Pirquitas two stage gravity concentration plant assembled at Gekko, Ballarat

After arrival at the Pirquitas mine site, a local team of construction personnel mechanically installed the

majority of the hardware, including frames, equipment and pipe work, following a Gekko designed modular

system that allowed them to construct in a Meccano-set style approach with little interaction from Gekko.

Electrical instrumentation and drives were then installed under supervision and assistance from Gekko

electricians.

Gekko provided services throughout the installation and commissioning period that included personnel with

specific expertise in mechanical, electrical, process control, metallurgy and training. This enabled Gekko to

provide the technical expertise to resolve issues rapidly and provide the essential knowledge sharing and

relationships to ensure the project will be a long term success.

Despite language barriers, Pirquitas mine staff and Gekko worked together and provided the required

resources from both sides with language skills that ensured focus and progress of the project was

maintained at all times.

PLANT PERFORMANCE During commissioning, throughput to the IPJ circuit was reduced to 130 tph to ensure that each bottleneck

was identified and rectified prior to the next staged increase. It also allowed time for the crushing circuit to

be upgraded to the design 6000tpd capacity. On several occasions the plant was taken to rates of up to

230-240tph for short periods or 165-170tph feed to IPJs, but these were for periods too short to monitor

recovery performance.

The initial lower than designed crushing circuit throughput resulted in the opening up of the screens in the

crushing circuit to increase crushing rates and resulted in material up to 25mm reporting to the circuit.



Although the +12.7mm material was less than 1% of the feed to the IPJ’s, occasionally stray rocks up to

100mm would enter the circuit and lead to immediate blockages and often significant delays in operation.

This lead to the installation of a 25mm scalping screen on top of the 2mm banana screen and the blockage

issues were resolved. Ultimately the mine should be targeting 20mm if not 15mm aperture screens to

ensure that all feed has equal opportunity to report to concentrate and to minimise the hang up of coarse 20-

25mm in the IPJ beds. The majority of these larger rocks were also noted to be heavily mineralised and

should have gone to concentrate, but could not due to the internal IPJ screen size.

The IPJ 3.2 SG ragging used in the first fill was soon replaced with 5.5 SG ceramic and 7.0 SG steel as the

heavily mineralised coarse rocks mentioned previously were found to displace the ragging out of the inner

rings of the IPJ bed. Once the ragging was displaced there was no restriction of flow of feed to concentrate

and the concentrate grade was subsequently diluted. In the end the beds all had a mixed bed of these

ragging types, but they generally went from high densities in the inner rings to the lighter in the outer rings.

The ragging size was generally 30mm for the 3.2 and 5.5 ragging, but the steel was commercially available

ball bearing and grinding media at 25mm.

During November 2010, operational data was collected from the plant to provide an understanding of the

performance of the process. A typical mass balance is shown in Fig. 11 and indicates the tonnage, silver

grade and recovery expectations of the plant.

180 tph 180 g/t

32400 g

130 tph 132 g/t 260 tph 180 g/t

17086 g 46800 g

50 tph 304 g/t 50 tph 296 g/t 80 tph 30 g/t

15314 g 86% Rec 14686 g 2400 g

100 tph 300 g/t 80 tph 180 g/t

93% Rec 30000 g 14400 g

180 tph 247 g/t

44400 g

180 tph 300 g/t

54000 g

Total Feed to Mill

Jig Plant Feed

Target Feed

Total Plant Feed

Banana Screen Feed

Fines Concentrate Tails

Jig Feed to Mill Auxiliary Feeder

Fig. 11: Mass balance around IPJ circuit



Due to the lower tonnes produced from the crushing plant than designed, passing all of the ore through the

IPJs would regularly produce insufficient feed to the grinding circuit and cause starved conditions that would

produce finer grinds and in turn lower silver recoveries in the flotation circuit. In addition, the bypass of -2mm

to the ball mill discharge increased the “efficiency” of the ball mill which also enhanced the grind produced.

Recommendations were made to reduce the layers of ragging in the IPJ beds to provide higher concentrate

production, however this was not supported by site as focus was to produce high grades from the IPJ circuit

and proceed slowly on process changes to enable an improved understanding of the IPJ variables to be

made.

The splits or yields of the various products namely the fines bypassed directly to grinding circuit, concentrate

stream and tails produced; are shown in Fig. 12. The actual grade to the grinding circuit is diluted by the

auxiliary feed (as shown in the mass balance in Fig. 11 and depicted in Fig. 12) used to make up the tonnes

for the appropriate ball mill load.

0%

10%

20%

30%

40%

50%

60%

70%

80%

90%

100%

0.0

50.0

100.0

150.0

200.0

250.0

300.0

350.0

400.0

Tonnage (tph)

Date

Pirquitas Jig Tonnes Production

Calc. Conc. Fines Bypass Auxilary Feeder Jig Reject % Metal to Mill

Fig. 12: Tonnes distribution in feed to the Pirquitas milling circuit

Fig. 13 shows the budget of silver metal production and the minimum target grade of the feed to the grinding

circuit by the horizontal line. The vertical bars indicate the metal feed per hour to the grinding circuit

inclusive of the diluted feed from the auxiliary feeder with recoveries in the 90-95% region (excluding

auxiliary feeder material). The O to X shows the increase in grade to the grinding circuit which depicts an

average of 153% grade increase (inclusive of auxiliary feeder), however the grade improvement across the

IPJs alone shows an average 201% increase prior to auxiliary feeder dilution.



0

100

200

300

400

500

0

10000

20000

30000

40000

50000

60000

70000

80000

90000

100000

Mill Feed Grade (g/t)

Metal To

Mill (g/hr)

Date

Pirquitas Jig Plant Daily Actual Metal vs Budget

Actual Metal to Mill Budget ‐Metal and Grade Mill Grade Feed Grade

Fig. 13: Silver grade and metal feed rate to the Pirquitas milling circuit

The higher grinding circuit feed grade for Pirquitas was found to be very important for flotation recovery.

Grinding circuit feed grades maintained in the 300-400g/t Ag range compared to the average 180g/t Ag mine

production grade were found to produce superior flotation recoveries generally providing a 10-20%

improvement in performance and at higher than planned flotation concentrate grades. This somewhat

unexpected benefit reduced smelter penalties and improved overall economics.

The actual results are very encouraging as the data follows the main trend and stays within the variances

seen in the Gekko testwork results (Fig. 14).



0%

10%

20%

30%

40%

50%

60%

70%

80%

90%

100%

0% 10% 20% 30% 40% 50% 60% 70% 80% 90% 100%

Silver Recovery (%)

% Mass Yield

Pirquitas ‐ Actual Performance vs Testwork Data

Test Work Maximum Minimum Plant Data

Fig. 14: Silver recovery versus mass yield for test work versus plant data

Fig. 15 depicts the distribution of silver metal at the Pirquitas project. The auxiliary feeder input is at the

mined grade without any beneficiation, added to the post IPJ processing and the minus 2mm stream and fed

directly to the grinding circuit. The fines were found to be naturally high in grade but are only generally

around 20% of the feed material and are bypassed directly to the ball mill discharge hopper. The +2mm

material is fed to the IPJ circuit and split approximately 50:50 to concentrate and tail and upgrades the

concentrate by approximately 180% of jig feed grade. Tails are sent to the rejects pile at grades typically

between 10 and 30g/t Ag.



0%

10%

20%

30%

40%

50%

60%

70%

80%

90%

100%

Metal Deportation (%)

Date

Pirquitas Jig Plant Metal Deportation

Conc Fines Aux. Feeder Tails

Fig. 15: Metal distribution at the Pirquitas plant

Apart from the lower than planned tonnage mentioned previously, the pre-concentration circuit operated as

planned producing a higher grade to the grinding circuit for greater silver production and higher recoveries.

There were some challenges with the pre-concentration circuit experienced during commissioning and

optimisation. These included:

Excessively large rock entering the plant that contributed to low availability due to blockages and

excessive wear to sections of the plant’s pipe work. This issue was resolved by retrofitting a

scalping screen strapped to the banana screen. Although not ideal it did perform OK and eliminated

the blockage problems that were being experienced. However due to spatial limitations, the

modification restricted flow and hence feed tonnes that could be fed to the banana screen.

Eventually a properly installed safety scalping screen will need to be installed into this circuit to

remove this bottleneck.

Pumping of coarse gravel without any fines proved to be challenging as most pump models take into

account the fines to assist in the transport of the coarser material. Higher than expected pipe

velocities were required to maintain proper operation of the gravel pumps, but the higher velocities

lead to accelerated wear. The high level of instrumentation and control in the plant logic system

allowed a modified operational philosophy to operate the pumps according to this essential criterion.

Severe wear was mainly found in the suction of the IPJ feed pump and extended life in this area was



achieved with design changes to the flow pattern of the material to the pump’s suction inlet and

improved material selection to line these parts.

Controlling the split of concentrate yield on high mass pull circuits is largely attributed to machine

variables (i.e. screen size, ragging density and size, ragging depth) as opposed to traditional thinking

where operating variables (i.e. stroke length and pulse frequency) play the larger role. In “through-

bed” jigs, the machine variables make up 80% of the yield characteristics and the operating variables

are used to fine tune the yield. This change in concept is difficult to transfer to operators that have

experience in traditional jigging culture. This does not always provide the flexibility that some

operations would like, but generally in relatively steady state plants these machine variables will not

alter significantly once the right operating parameters are established.

CONCLUSIONS The performance of Silver Standard Resources Inc’s silver pre-concentration plant utilising InLine Pressure

Jigs has verified the scale-up from test work to plant performance. The challenges of dealing with truncated

coarse feeds have been overcome by mine and Gekko personnel and resulted in a robust operating plant

that is exceeding its targets.

ACKNOWLEDGEMENTS The authors would like to acknowledge Silver Standard Resources Inc and Gekko Systems Pty Ltd for their

permission to publish this paper.

REFERENCES Gray, S, Hughes, T, 2007. A Focus on Gravity and Flotation Concentration and Intensive Leaching Rewrites

Conventional Milling Circuit Design and Improves Environmental and Cost Outcomes, paper presented to World Gold

Conference 2007, Cairns, 22-24 October 2007.

Gray, S, and Hughes, T, 2008. Improvements in the InLine Pressure Jig expands its applications and ease of use for

gold, silver, sulphide and diamond recovery, in Proceedings Metallurgical Plant Design and Operating Strategies 2008,

(Australasian Institute of Mining and Metallurgy: Perth).

Hatch Engineering and Mine Development Associates, 2006, 43-101 Silver Standard Resources Pirquitas Silver, and Tin

Project, 9 May 2006.

MacRae, PS and McCrea, JS, 2008. Silver Standard Resources Inc, Technical Report on Minas Pirquitas, Silver, Tin and

Zinc Project, Jujuy Province, Argentina, 29 September 2008, Available from www.silverstandard.com [Accessed: 01 May

2011]

Silver Standard Resources Inc, 2011. Silver Standard Reports Fourth Quarter and Year-End 2010 Results, 1 March

2011, Press Release available from www.silverstandard.com [Accessed 01 May 2011]



Energy Efficiency Opportunities in Milling – Improving

Comminution Circuit Efficiency

F Musa1, M Stewart2 and G Weiss3

1. Consultant, Energetics Pty Ltd, Unit 2, 313 Boundary Street, Spring Hill Qld 4000.


2. MAusIMM, Group General Manager, Energetics, Level 7, 132 Arthur Street,

North Sydney NSW 2060. Email: [email protected]

3. Principal Consultant, Energetic, Level 7, 132 Arthur Street, North Sydney NSW 2060.


ABSTRACT

One of the main energy efficiency opportunities identified from the majority of mine site energy

audits is to optimize the comminution circuit to reduce energy consumption through improved

throughput and product size distribution. Typical analysis reviews comminution process

efficiency/inefficiency i.e. utilization of grinding mills, process bottlenecks and product vs. target

size, to identify instances of overgrinding or undergrinding. From this, it is possible to establish the

potential improvement that could be gained through formal plant optimization and what the

optimization program would entail. It can also sometimes reveal immediate opportunities for process

improvement. This paper presents steps that can be taken to optimize the efficiency of comminution

circuits.

INTRODUCTION

Energy use is a critical issue for the resources industry. In Australia, the cost and availability of

energy influences the growth of the resources industry, the extent of processing carried out at

individual sites and the long-term profitability of the sector. The minerals industry consumed about

30% of all electricity generated in Australia in 2009 (ABARE, 2010).

Where they are used, the comminution processes that liberate the valuable minerals are usually the

largest consumers of electricity. As an example, Fig.1 illustrates a breakdown of energy use by

activity at an open pit gold mine producing on average 122,000 ounces of gold per annum

(Energetics, 2009). It can be seen that diesel is dominantly used in mining operations mainly for haul

truck and excavator activities, and the electricity is predominately consumed in ore processing,

particularly grinding.



Fig.1 Energy Use by Activity

The aim of comminution processes is to reduce particle size and it may involve one or several stages

of size reduction. Each stage has a specific energy demand to achieve the required size reduction. The

energy needed at each stage increases as the product size decreases. Energy requirements for

comminution also depend on ore type and ore grade. Indicative energy requirements and comminution

efficiency for three types of comminution devices are given in Fig. 2. Efficiency for a crusher was

estimated using the Whiten crusher model (Andersen and Napier-Munn, 1988) and efficiency for ball

mills and SAG mills was estimated using the Perfect Mixing Model (Musa and Morrison, 2008).

Fig. 2 Average energy requirements and comminution efficiency for AG/SAG mills, ball mills and crusher.



The increased cost of energy, supply constraints and the need to manage greenhouse gas emissions are

all driving the industry to reduce energy demand and improve energy efficiency. Given the

importance of the comminution processes to production rate, yield and operating costs through energy

use, there have been recent efforts by mines to improve the energy performance of the comminution

processes. This paper explores the impact of comminution on energy use and offers suggestions on

how to improve the performance of the comminution process, and in particular their energy

performance.

OPERATION OF COMMINUTION CIRCUITS

As indicated above, comminution typically involves a number of stages – crushing, grinding and

separation. Ideally, comminution circuits are designed so that all of the individual unit operations are

fully utilised. In practice, this is seldom the case. When ore characteristics change, the operating

characteristics of the circuit change as well, making part of the circuit operate less efficiently than the

design case, which then forces other unit operations to run at less than their rated capacity. The result

is a constraint on the ore processing rate.

Fig. 3 shows the fluctuation in throughput of the ball mill at three different sites for a 12 months

period. This is compared with the nominal baseline capacity of the mills estimated using the Bond

equations as shown below. Periods of reduced throughput are clearly evident.

0

100

200

300

400

500

600

Jan Feb Mar Apr May Jun Jul Aug Sep Oct Nov Dec

Site A Site B Site C

Baseline for site A

Baseline for site B

Baseline for site C

Tonn

age,

tph

Fig. 3 Tonnages through ball mill at three different sites

The throughput was estimated using the equation:

)/ tkWhE

kWdemandpowermillT



The specific energy, E was calculated using the Bond equation:

8080

1010

FPWE i

where E is specific energy (kWh/t) Wi is the Bond work index (kWh/t) P80 is the 80% passing product size (µm) F80 is the 80% passing feed size (µm)

Kawatra and Eisele, 2005 suggested that the two main causes of energy loss in grinding circuits are

underutilisation of equipment due to process bottlenecks and over-grinding due to material being

retained in the circuit after it has reached its target size.

Most plants have a good idea where their operational bottlenecks are. Still identifying the cause of the

bottlenecks is not an easy task. Plant reviews involving intensive examinations of trend graphs and

information from the control room are typically used to identify grinding circuit bottlenecks. During

plant surveys, information such as circuit throughput, mill throughput, mill operating hours and mill

power are collected. There are a number of computer models which can be used to identify process

bottlenecks but a simpler method of plotting power used by different comminution units can provide a

quick indication of the bottleneck of the process. One example is shown in Fig. 4. The SAG mill

power draw is nearly constant while the ball mill power draw varies substantially thus meaning that

the SAG mill is the throughput bottleneck in this instance.

Fig. 4 SAG mill power vs Ball mill power

Identifying process bottlenecks is just the first step. The next is to determine the cause of the

bottleneck. Comminution circuit bottlenecks could be caused by a number of factors. A typical

comminution circuit consists of a crusher, a SAG mill followed by a ball mill and a nest of

hydrocyclones as shown in Fig. 5.



Fig. 5 A typical comminution circuit

In identifying the causes of a bottleneck in comminution circuits, the approaches taken include:

Evaluation of ore characteristics. The ore hardness or the work index can be used as a

measure. Changes in ore hardness will have a significant impact on the grindability of the ore

and hence the efficiency of grinding mills.

Assessment of mill charges. This includes particle size distribution of charge, filling level and

ball size. Typical SAG mill charge level is 30-35% with a ball charge of 6-15%. For ball

mills, charge level is typically between 45-50% with a ball charge of 32-35%.

Assessment of mill circulating load. The typical circulating load is about 250%. Higher

circulating load means that the ore is not being efficiently ground in the mill and needs to be

reprocessed in the mill for a second or sometimes more times to achieve the required

breakage. In SAG mills, materials of size 25mm to 50mm (sometimes called the critical size)

often do not get broken and hence tend to stay in the mill. A large amount of these critical

size materials restricts the amount of new feed coming into the mill, and hence limits

capacity.

Evaluation of hydrocyclones. Poor performance of hydrocyclones means poor separation of

particles. Often this increases the circulating load of the mill. It is important to ensure cyclone

feed density and cyclone dimensions are appropriate for the duty.

Overgrinding. Use the Bond equation to assess if there is overgrinding at the plant.



ENERGY OPTIMISATION OF COMMINUTION CIRCUITS

Within the constraints of throughput, a variety of factors will influence the energy performance of the

comminution circuit. Mine engineers and metallurgists recognise the importance of tight control of

feed to the comminution circuit. Comminution capacity and efficiency are strongly influenced by the

run-of-mine fragmentation distribution, which in turn is influenced by the blasting. There are two

important parameters of blasting on fragmentation, the size distribution of blasted fragments and

crack generation that occurs within fragments (Workman and Eloranta, 2003). In the process of

optimizing blasting it is very important to know that the fragmentation distribution is adequate.

Consideration must also be given to how blasting will precondition individual fragments by internal

fracturing. While the first factor is now measurable directly, the second must be evaluated through

study of production and energy consumption. As an example consider ore that is being blasted with a

heavy ANFO having effective energy of 3.30 MJ/kg. The 80% passing size of the blasted ore is 31.75

cm. The ore passes through primary crushing, primary and secondary grinding. The final product is

80% passing 0.007 cm. The work index for the ore is 17.43 kWh/t. The specific energy is calculated

using Bond equation as follows:

FP

WW i

1010

where W is the specific energy (kWh/t), Wi is the work index (kWh/t), P is the 80% passing size of

the product (µm) and F is the 80% passing size of the feed (µm). Table 1 shows the feed and product

size, the calculated total energy, and the energy cost for each unit operation. The explosive cost is

$2.30 per kg. Electrical energy cost is assumed to be $0.80 per kWh.

Table 1 – Energy and cost calculation by unit operation

Operation Feed 80%

passing size (cm)

Product 80%

passing size (cm)

Specific Energy

(kWh/t)

Energy cost

($/t)

Explosives 31.75 0.31 0.67

Primary crushing 31.75 10.2 0.24 0.03

Primary grinding 10.2 1.5 0.88 0.11

Secondary grinding 1.5 0.007 19.41 2.35

Totals 20.83 3.15

It can be seen that grinding has the highest work input. Studies have shown that enormous amounts of

energy can be saved by optimising the size distribution of blasted rock fragments for specific

downstream operations i.e. comminution. The energy consumed can be changed in three ways:



(1) reduce the feed size to the primary crusher which means that less energy will be required

to crush the ore to the same product size,

(2) a decrease in work index, Wi related to additional macro-fracturing and micro-fracturing

within individual fragments, and

(3) an increased percentage of undersize that bypasses stages of crushing thereby decreasing

the percentage of total tons crushed (this is enhanced if there is a mill by-pass route available

in the process flowsheet).

These energy efficiency opportunities can potentially be achieved using mine-to-mill optimisation

study such as JKMRC Blast Fragmentation Model and CRC Mining’s Integrated Extraction Model

(Sustainability and Extraction Efficiency Software).

There are a few considerations that could be taken to optimise comminution circuits:

Load balancing in the grinding circuit. In order to fully utilise all of the available capacity in

the grinding circuit, it is necessary to shift the grinding load to the point in the circuit where

extra capacity is available. When operations use several stages of size reduction, there may be

scope to increase overall throughput by adjusting the transfer point between comminution

stages.

Check mill charge level. If charge level is lower than the typical safe working level, it is

possible to increase tonnage by utilizing this free volume. SAG mill throughput can be

increased by increasing the ball charge.

Control of SAG mill feed. The main problem in controlling AG/SAG mills is the tendency to

build-up the critical size faction in the mill. Pre-treatment of mill feed should be done to

minimise the presence of this fraction. For material of 50-90 mm size, a pebble crusher can be

used to crush these scats before returning them to the mill. High Pressure Grinding Rollers

(HPGR) can be included in the grinding circuit to take care of the presence of the critical size

material. It has also been claimed that the final product was more amenable to flotation and

leaching circuits. Another alternative is to close the SAG circuit with hydrocyclones. This

will increase mill filling level and power draw however it is important to check the product

size.

Check ball size and number. Increasing the ball size and decreasing their number would

increase the kinetic energy of breakage but the frequency of impact would decrease. On the

other hand, increasing ball numbers increases the breakage rate and therefore increases the

throughput rate. If product from a SAG mill is reduced to finer sizes from installation of

cyclone in the AG/SAG circuit, the ball size in the ball mill could be reduced for better

grinding. This can help to reduce mill circulating load and lower the cyclone underflow

percent solids.



Check classifier performance. For hydrocyclones, it is important to ensure that the feed is

diluted and the cyclone is running at a sufficiently high pressure. This can be done by

increasing water addition to the feed, increasing the spigot and adjusting the pressure.

Better separation. A common objective in optimisation is to match product sizing criteria as

accurately as possible to comminution equipment capability (Napier-Munn et al., 1996). That

is, too coarse a product results in a low circulating load and poor energy utilization, while too

fine a classification cut-size results in a large circulating load and inefficient production of

excess fines (overgrinding).

Reducing overgrinding. In some grinding circuits, overgrinding is a particular concern. This is

because not only energy is wasted for unnecessary grinding, but it can also produce fines that

are much smaller than can be processed efficiently. For instance, a grinding circuit has a

target size of 80% passing 45 µm but a size analysis of the operating product gives a product

size of 80% passing 40 µm then using the Bond equation will reveal the estimated energy

wasted in overgrinding. Given the average work index of the ore is 12.5 kWh/t, the energy

wasted can be calculates as follows:

tkWhW /13.145

10

40

10*5.12

For a plant processing 1 million tonnes of ore per annum and electricity cost of $0.80 per

kWh, about $900,000 is wasted due to overgrinding.

Pre-concentration. One opportunity is to consider gangue rejection from a circuit by ore

sorting techniques.

There are a number of approaches that can be used to carry out circuit optimisation. Some of the

examples above have used the conventional Bond method using single point size – F80 and P80 (Bond,

1961).This is reasonable provided the feed and product size distributions are parallel on a log/log plot.

Another new approach is to look at energy in terms of energy to produce new material in the fine

sizes. Typically this is in the minus 75 micron region (Musa and Morrison, 2008). At the expense of

more detailed data collection, computer modeling and simulation using JKSimMet software can be

used for this optimization work. It typically involves complete plant surveys followed by mass

balancing and model fitting the survey data to create a model of the circuit in JKSimMet. The model

is then used to assess and compare options for optimisation of the plant operation.

CONCLUSIONS

A number of opportunities to optimize the energy efficiencies of comminution circuits have been

described. The main focus of optimization is to eliminate process bottlenecks by increasing process

throughput. Bottlenecks are eliminated through:

better control and balancing of the circuit,



circuit redesign and changing flow paths for example removing undersize from feedstreams,

and

integrating mill requirements into mine operations and blasting.

In this paper, the significant of comminution circuits in the national energy consumption profile have

been outlined. Tips on how energy consumption can be reduced are given. When operations get it

wrong, the value in reducing energy use is not trivial when quantified as avoidable operating costs.

REFERENCES

Andersen, J and Napier-Munn, T J, 1988. Power prediction for cone crushers. Third Mill Operators

Conference. Cobar, NSW (AusIMM).

Australian Bureau of Agricultural and Resource Economics (ABARE), 2010. Energy in Australia

2010. ISSN 1833-038

Bond, F C, 1961. Crushing and grinding calculations part 1, in British Chemical Engineering. 6(6) pp

378-385.

Energetics, 2009. Internal Report. Energetics Pty Ltd. Australia.

Kawatra, S K and Eisele, T C, 2005. Optimization of comminution circuit throughput and product size

distribution by simulation and control, in Final Technical Report, Michigan Technological University.

Musa, F and Morrison, R, 2008. A more sustainable approach in assessing comminution efficiency, in

Comminution 2008. CDROM. MEI: Falmouth.

Napier-Munn, T J, Morrell, S, Morrison, R D and Kojovic, T, 1996. Mineral comminution circuits –

Their operation and optimization, in JKMRC Monograph Series in Mining and Mineral Processing 2.

Workman, L and Eloranta, J, 2003. The effects of blasting and crushing and grinding efficiency and

energy consumption. www.elorantaassoc.com/Blasteffect_Lyall.pdf



Non-Contact Acoustic Measurement of Dynamic In-Mill Processes for SAG/AG Mills

R A Pax1

1. MAusIMM, Principal, RAP Innovation and Development, PO Box 559, Indooroopilly Qld 4068. Email: [email protected]

ABSTRACT

Heavy machinery such as AG and SAG mills generate a significant amount of noise due to the processes

that are occurring inside them. From an environmental perspective, the level of noise generated is a

significant occupational safety hazard. However from a technical perspective the noise presents an

opportunity to determine some details of the processes occurring inside the mills. The noise generated is

not random and can be attributed to a range of events that occur inside these mills, consequently non-

contact acoustic determinations of in-mill behaviour provide powerful insights into the operation of the

primary comminution tumbling mills.

Traditionally, the use of perfect mixing has been assumed when modelling the operation of AG/SAG mills. If

this were true then the acoustic emissions from AG/SAG mills would be uniform along the length of the mill.

Experimentally this has been found not to be the case. From another perspective it should not be the case

because of the different conditions existing at the feed end compared to the discharge end. This paper

presents the results determined by non-contact acoustics and their analysis in terms of understanding of the

mechanisms occurring inside a mill.

INTRODUCTION The environments associated with the breakage of rocks inside comminution equipment, in particular SAG

and AG mills, are inherently hostile that preclude the incorporation of direct sensing elements inside the

equipment. Traditionally, the success of the breakage events that are the main purpose of SAG/AG mills,

has been monitored by measuring the output particle size distribution and comparing it to the input rock size

distribution, for a variety of mill operating conditions. These operating conditions include mill speed, internal

load of rocks and media, water, rock and media addition rates as well as the mill design criteria that include

mill diameter and length, lifter and discharge grate design and the pulp lifters. The mill power draw is also

closely monitored. Operationally, the approach to mill operation described above is implemented with

circuit/mill surveys and implicit with this approach is the delay in obtaining valuable information from the

laboratory analyses (Napier-Munn et al., 1996).

The input/output approach described above does not detail the internal mechanisms that are causing the

breakage and transport of rocks through the mill. These need to be visualised by inference from the use of

appropriate experimental design and mathematical models. The models have tended to be empirical in

nature and consequently are challenged to incorporate detailed process mechanisms. Recently, discrete

particle models have been used to investigate in-mill processes (Cleary et al., 2001).



Measurements techniques to determine internal mill conditions have generally consisted of instrumentation

to measure strain of the mill shell or liner bolts (Moys, 1984) or sound emissions from the mill. Direct strain

measurements infer charge behaviour from the interaction of charge with the mill shell components. Sound

emission determinations using microphones measure the dynamic responses of all components of a mill.

However the response from the charge inside the mill together with its interaction with the mill shell

dominates the sound emissions from an AG/SAG mill.

Acoustic emissions, ‘noise’, from industrial equipment is common place and is often considered an

environmental hazard if excessive. From another perspective the acoustic emissions are a window into the

dynamic processes that are occurring inside the equipment. Static processes are usually of little interest for

mineral processing, but have importance for equipment designers. Since acoustic emissions are mechanical

waves in a medium, they may be detected using strain measuring equipment or using a non-contact

methodology such as microphones. This paper is concerned with the use of non-contact acoustics which

deploys microphones external to a mill to obtain internal mill operational information in real time.

NON CONTACT ACOUSTIC MEASUREMENT METHODOLOGY

Microphones are essentially very sensitive pressure transducers that convert dynamic mechanical energy to

an electrical signal that can then be processed and stored in a variety of ways. The mechanical energy,

originating from processes occurring inside the mill, manifests itself in an acoustic signal as frequency

components within the bandwidth of the transducer. Ideally the individual mechanical frequency

components should be faithfully reproduced in the generated electrical signal over the entire bandwidth of

interest, so that no new frequency components are generated, nor the relative magnitudes of the

components altered. Good microphone design requires an essentially flat frequency response with minimal

or no frequency generation, although a non-flat frequency response can be compensated with appropriate

calibration.

The dynamic range of a microphone is also very important so that the least intense of the frequency

components of interest can be monitored accurately at the same time as the most intense components. By

virtue of a microphone’s manufacture, some directional sensitivity is available with a given microphone, so

that acoustics signals from behind the microphone or beside it are attenuated when compared to signals

originating from directly in front of the microphone. This property is used to advantage to provide localised

information from a SAG/AG mill.

Previous work (La Rosa et al., 2008) have tended to use a single “microphone” in a specific location,

normally located close to the mill near the expected charge toe location. Sometimes the microphone is

located further away from the mill. There are distinct advantages in using microphones close to the mill shell

which include localisation of information as well as an inherent removal of background signals since the mill

noise is the dominant signal. Background signals will usually be significantly weaker than the noise being

generated by the mill.



Localisation of acoustic signals has the advantage of providing information that is specific to the part of the

charge that is accessible for that microphone. The directional response of the microphone is also useful so

that the signals emanating from other locations are attenuated. By appropriately locating a number of

microphones around a mill, a consistent visualisation of internal charge behaviour of SAG/AG mills can be

obtained (Pax, 2001). The reception areas inside a mill are shown in Fig. 1 for microphones located in a

quadrant array.

Toe array

Mill rotation

T2

T3

T4

Q1

Q4

Q3

90°

270°

0° 180°Q2

T1

Fig. 1 – Reception regions for four microphones located around a SAG mill

For SAG and AG mills the noise generating mechanism is principally due to steel balls hitting each other and

the mill liner, as well as collision events involving the ore load. These events are of interest in the operation

of the mill, since they directly influence the wear of the mill components and the breakage of rocks. To

enable analysis of the signals obtained from each location a set of descriptors has been developed, which

are used to characterise the signal in terms of a linear noise level, a steel signal, and a characterisation of

hitting events. The signals are analysed as a function of time and frequency content. For certain events that

occur inside the mill, the frequency content of the signal is time dependent, which can be used to advantage

for appropriate discrimination.

The base signals are then used in combination to provide specific information at locations around a mill.

Environmental noise is traditionally characterised by a logarithmic decibel scale since our ears respond to

noise logarithmically. From an instrumentation perspective this scale is of limited use, because it reduces the

dynamic range for the signals of interest, consequently a linear scale is used instead. This has also the

advantage that superposition principles can be used to interpret the measured noise data and thus isolate

the processes occurring inside a mill.



Mill Rotation

M4

M8

M5M1

M6

M7

M3

M2

ShoulderArray

ToeArray

Fig. 2 – Microphone positions for a toe and shoulder array with angular definition in degrees.

Arrays of microphones have been used in this work to obtain information about the internal working of

SAG/AG mills. The preferred configuration is four microphones in a shoulder array and four microphones in

a toe array as shown in Fig. 2. These microphones are 30 degrees apart within one array and allow

localised discrimination to help identify events inside the mill. Four microphones located every 90º starting at

0º is termed a quadrant array. The attenuation of the sound generated inside the mill, due to the mill shell

and liners, is the same irrespective of the microphone’s location.

In this paper, either a quadrant array, or a toe array or partial shoulder and toe arrays is used on three

different mills. All mills process base metal ores.

RESULTS AND DISCUSSION

Fig. 3 shows the locations of the non contact acoustic measurements for SAG mill A. Locations one to six

are on the toe side of the mill, seven to 12 are on the shoulder side of the mill, whilst thirteen is directly

underneath the mill. Apart from measurement number 13, all odd numbered measurements are located just

below the centre line of the mill and the even numbered measurements were approximately 30º towards the

bottom of the mill.



1

28

7 Top view

13

1

7 9

3 5

11

Feed End

Discharge

End

10

2

8

4 6

12

0˚

90˚

180˚

Fig. 3 – Diagram showing the locations of measurements taken around the mill A.

Fig. 4 – Noise levels at various locations around mill A. The number sequence (1-13) refers to the locations

specified in Fig. 3.



Fig. 5 – Steel signals at various locations around mill A. The number sequence (1-13) refers to the locations

specified in Fig. 3.

Fig. 6 – Mill A: Mill shell steel hits (MSSH) at various locations around mill A. The number sequence (1-13) refers

to the locations specified in Fig. 3.



Fig. 7 – Transient information of the (MSSH signal) at various locations around mill A. The number sequence

(1-13) refers to the locations specified in Fig. 3.

Fig. 4 shows the noise level at locations specified in Fig. 3. The noisiest location is directly underneath the

mill. The noise level from the toe region (measurements 1 to 6) is louder than that from the shoulder side of

the mill as would be expected. On the toe side, the noise level from the microphones located below the

horizontal (2,4,6) through the mill (≈210º) are noisier, indicating that more activity is occurring at these

locations. The expected activity is ball hits on the mill shell as well as ball-rock and ball-ball interactions. On

the shoulder side some discrimination between the two vertical positions is also evident.

Fig. 5 shows the steel signal for the various locations, this signal has been labeled Hits3 in the graph. The

steel signal is derived from the collected noise data and is unique to any steel interactions. For a rubber

lined AG mill, the steel signal has been observed to drop to zero as would be expected for this type of mill.

Overall better discrimination is now possible between the various locations around the mill. In particular,

significantly more steel activity occurs at the lower microphones on the toe side of the mill, indicating that

that is where the steel balls are interacting with the mill shell and other steel balls.

In contrast, at the shoulder locations, it is the higher location (≈0º) that has the greater steel activity. This

situation indicates that the charge is more dynamic at the higher location, which would be expected as the

balls and rocks are released from the liners of the mill. What is surprising is that this situation is occurring at

(≈0º), which is low for the shoulder position of a SAG mill. In fact, this particular mill has a packing problem,

evidenced by a visual inspection inside the mill.

The variation that is evident in each of the measurements shown in Fig. 5 indicates the dynamic nature of

the charge behaviour in the mill. The time interval for each datum corresponds to 80 mS, and so precludes

the witnessing of individual events in the mill and instead shows clusters of events. Some very large events

(transients) are evident at microphone position 4 and 6. Measurement thirteen corresponds to underneath

the mill which shows a large amount of steel activity including transients. The characteristics of this signal



correspond to an indication of the looseness of the charge which also influences the ability of the charge to

act as an anvil so the breakage can occur by direct impact. A significant steel signal at this location also

corresponds to poor dilution of the steel balls by interspersed rocks. For this mill the rock to ball ratio was

approximately 0.5 by volume.

Also evident in Fig. 5 is the longitudinal variation in the steel signal, indicating that the processes are altering

along the length of the mill. It is often assumed for the analysis of AG/SAG grinding mills that they are

perfect mixers, i.e. that the charge distribution, in the mill is uniform and that the same sub processes apply

though out the mill. That the dynamic character of the acoustic signals, in particular the steel signals, are

varying longitudinally indicates that the assumption of perfect mixing is not true. Further analysis is

warranted.

By incorporating time domain information together with the steel signal, a composite signal can be derived

that describes the events that constitute steel balls hitting the mill shell. This signal is called the mill shell

steel hits (MSSH) signal and is shown for all locations in Fig. 6. Although of primary utility in the toe region,

the MSSH signal has meaning at all locations because of the mobility of steel balls in the charge. Fig. 6

shows that the lower toe microphones at ≈210º, have significantly more activity than the corresponding

microphones at ≈180º, providing an indication of the ball trajectories. Increased activity is evident along the

length of the mill, further supporting that the mill is not perfectly mixed. The shoulder microphones show

significantly lesser activity than the toe microphones, consistent with their location and the interpretation of

the signal.

Further analysis of the MSSH signal is possible by extracting the number of transients that occur in the

signal per unit time. The result, MSSH tr (transient), is shown in Fig. 7 and illustrates the value of this

analysis by the significantly increased dynamic range and the discrimination between microphone locations.

The MSSH transients indicate the dynamic activity of the charge and steel balls inside the mill and further

show the lack of perfect mixing inside the mill.

The analysis of Figs. 4 to 7 has shown the utility of analysing the noise signals at various locations around a

SAG mill, and by illustration the excellent discrimination that can be achieved with appropriate signal

processing techniques that reveal the components that are inherent in the measurement signals.

Tests were also conducted on a fixed speed AG mill with an aspect ratio (D/L) of approximately unity and a

mill speed of 74.6 % critical. The liners/lifters were made of steel and so significant steel-based signals were

obtained. Mounted on this mill is a quadrant configuration of microphones. This configuration was mounted

on an overhead gantry that allowed the quadrant configuration to be swept between the feed end and

discharge end in approximately 2 minutes.



0 5 10 152000

3000

4000

pow

er (

kW)

0 5 10 15140

150

160

170

load

(T

)

0 5 10 15200

400

600

(fee

d (T

/hr)

0 5 10 1515

20

25

(wat

er (

m3 /h

r)

0 5 10 150

0.2

0.4

Equ

ality

0 5 10 150

0.5

1

Time (min)

Dis

ch

F

eed

Fig. 8 – Mill B data. The process variability was kept to a minimum during the test work. The last two graphs

show the equality parameter and the position of the microphone arrays along the length of the mill.

0 5 10 150

5

10

15

90o m

icro

phon

e

0 5 10 150

20

40

180o m

icro

phon

e

0 5 10 152

4

6

8

270o m

icro

phon

e

0 5 10 150

5

10

0o mic

roph

one

0 5 10 150

0.5

1

Time (min)

Dis

ch

F

eed

Fig. 9 – Mill B. Results of the noise level data as a function of the position of the microphones along the length

of the mill.



0 5 10 150

20

40

90o m

icro

phon

e0 5 10 15

0

10

20

180o m

icro

phon

e

0 5 10 150

50

270o m

icro

phon

e

0 5 10 150

20

40

0o mic

roph

one

0 5 10 150

0.5

1

Time (min)

Dis

ch

F

eed

Fig. 10 – Mill B. Large hitting events (abnormal signal) along the length of the mill.

0 5 10 154

6

8

10

90o m

icro

phon

e

0 5 10 150

10

20

180o m

icro

phon

e

0 5 10 155

10

15

270o m

icro

phon

e

0 5 10 154

6

8

10

0o mic

roph

one

0 5 10 150

0.5

1

Time (min)

Dis

ch

F

eed

Fig. 11 – Mill B. Steel signal variation along the length of the mill.



Tests were conducted under reasonably constant process conditions to explore the longitudinal variations

that occur in this mill. Fig. 8 shows the process conditions for the duration of the 15 minutes test time. The

test time was kept short to ensure that the mill was operating under uniform conditions. The load of the mill,

as determined by load cells, the power draw, the ore feed rate and water addition rate were kept reasonably

constant with the exception of a disturbance in the ore feed rate at 0.5 minutes.

Fig. 8 also shows the position of the microphone quadrant at any time during the test, the value of 1 referring

to the microphones being located at the feed end of the mill and 0 to the discharge end of the mill. The same

plot is used as the last graph for each of Figs. 9-11.

It has been experimentally observed that the noise level indicated by microphones located at different

locations around a mill tend to become similar when water holdup inside the mill becomes significant such as

pooling events. Consequently a combined signal was derived using all four microphones in the quadrant

configuration and was labeled the equality signal. Fig. 8 also shows the equality measurement which varies

between zero and one. Of particular interest is that the equality output is higher at the feed end of the mill

compared to the discharge end of the mill. Recognising that the four microphones could give similar outputs

when there is either liquid or fine ore material present, then the higher signal at the feed end could represent

increased water (possibly because there is more material present as well) or increased fines material. The

latter is unlikely because it is at the discharge that most fines are expected. Consequently, the data of Fig. 8

indicates that there is a water gradient in the mill from the feed end to the discharge end.

The rise in signal at 5 minutes is also of interest. Considering that a material disturbance was injected into

the mill at 0.5 minutes perhaps this rise is due to that disturbance. ie. the average transport time for material

is about 4.5 minutes through the mill. Comparing the width of the disturbance, the dispersion is of order 3

which indicates that some mixing is happening along the length of the mill.

Fig. 9 shows the noise level (RMS) data for each microphone throughout the test. Although there are small

fluctuations in the top most microphone (90º) as the microphone is located from the feed end to the

discharge end, this microphone has the least variation in the signal, suggesting that not too much material is

being thrown from the shoulder of the charge to the toe region. This is also consistent with the minimal

variation in the up microphone at 0º. Having said that, the feed end is slightly noisier (at 90º) than the

discharge end as would be expected. The down side microphone (180º) shows the most variation, with

increased noise and variation in noise at the discharge end. Since the acoustic emissions are due to the

number of particles and the size of the particles (and also effective surface area) the increased RMS signal

at the discharge end could be attributed to these mechanisms. The underneath microphone at 270º shows

a similar variation along the length of the mill, supporting the notion of particle number and size being the

determining measurement this signal provides on this occasion.

Fig. 10 shows the “Abnormals” data for each microphone. The “Abnormals” signal is a measure of the

number of load hitting events. Of note is the low level of activity at each microphone, supporting again the

hypothesis that not much material is being thrown around in this mill. The only “Abnormals” for the down

side that do occur are at the feed end where the large rocks are being fed.



Fig. 11 shows the Hits3 (steel signal) data. This signal shows some very interesting variation along the

length of the mill. Of note is that the down side microphone (180º) does not show any variation along the

length of the mill which indicates that there is no activity within the reception angle for that microphone.

However the top microphone (90º) shows that there is significantly more steel signal at the feed end of the

mill than at the discharge end, consistent with large rocks interacting with the shell somewhere close by. In

fact the significant steel signal reduces to the discharge value approximately half way along the mill. The up

side microphone tells us where, by its similar steel signal values also at the feed end. The underneath

microphone tells the other story. The charge contact with the steel shell is greatest in the discharge half of

the mill.

All this data would seem to indicate that the majority of the feed rock is reduced to small particle sizes by the

middle of the mill. There would seem to be a water holdup in the mill, with the greatest retention at the feed

end. Again, longitudinal variation in the signals is clearly evident.

Fig. 12 – Comparison of an acoustically determined load of a mill (C) with a load cell located at the discharge

end of the mill as function of feed rate to the mill.

A composite experiment on a SAG mill was conducted using a toe array located at the discharge end of a

mill. Fig. 12 shows mill load variations effected by a sequence of the ore feed rate changes. This mill has an

aspect ratio of 2. An acoustic load measurement has been derived from the signal outputs from all four

microphones of the toe array. The results of acoustic load measurement (upper noisy trace) are compared

to the load cell measurement. The lower noisy trace corresponds to the feed rate changes. As expected,

there is a good correlation between all trends shown. However, the additional features of the acoustic load at

2.5x104 (units are 0.08 s) are of interest. As the feed rate is reduced (at constant feed sizing), the size

distribution of the rock particles in the mill is made finer and conversely coarsens when the mill feed rate is

again increased until a new equilibrium is reached. Overall good agreement in the trends has been achieved

with the acoustic estimator.

To further understand the results of Fig. 12, it should be noted that the load cell and the acoustic

measurements are both made at the discharge end of the mill. As a consequence the time of transport of

the load disturbance would cause the ore feed rate signal and load measurements of Fig. 12 to not exactly



coincide in time. The difference in the two signals could be used to describe the longitudinal behaviour of

the mill load. The changes in mill load were affected by the changes in mill feed rate. Consequently the

acoustic load signal trends correspond well with load cell measurement while the mill is unloading. However

when the mill is loading, the load takes some time to distribute longitudinally. The time from minimum ore

feed rate to the structure in the acoustic load is approximately 20 minutes (TD, Fig. 12) which corresponds to

the typical residence time of the large rocks in this SAG mill. The complicated structure of the acoustic load

signal is thought to be due to residence time effects in the distribution of particle sizes in the mill with new

feed coming into the mill.

CONCLUSIONS Using microphone arrays to listen to AG/SAG mills acoustic emissions, a study has been conducted of their

dynamic operating behaviour. From the measurements it has been determined that AG/SAG mills are not

perfect mixers, as is often assumed.

Non-contact acoustic measurement has been shown to be of use in discerning the internal processes that

occur inside AG/SAG mills. Good discrimination between differently located microphones around AG/SAG

mills has been achieved with the developed steel, MSSH and transient analysis signals, consistent with

expected internal mill processes.

Using arrays of microphones, the trajectories of particles inside a mill can be determined, so as to provide

direct feedback of liner/lifter designs. Incomplete mixing along the length of the mill suggests different lifter

designs might be appropriate longitudinally. The signals derived from the acoustic emissions from AG/SAG

mills can also be used within control schemes to adjust ore and media feed rates, mill speed and water

addition to ensure that the internal mill processes are adjusted to maintain desirable operation including the

mill load.

REFERENCES Cleary, P W, Morrison, R D, Morrell, S, DEM validation for a scale model SAG Mill, in Proceedings International

Conference on Autogenous and Semi-Autogenous Grinding Technology (SAG 2001), 2, (eds: D Barratt et al)

University of British Columbia, Vancouver BC, Canada, 2001, pp. 191-206.

La Rosa, D, Cantarutti, A, Wortley, M, Oskocak, T, 2008, The use of acoustics to improve load estimation in the

Cannington AG mill, in Proceedings of METPLANT Conference, 2008, pp 105-115, Australian Institute of Mining

and Metallurgy, Melbourne.

Moys, M H, 1984, The measurement of parameters describing the dynamic behaviour of the load in a grinding mill, in

Proceedings International Conference in Mineral Science and Technology 1984, pp 205-219, The Council for

Mineral Technology, Randburg, South Africa.

Napier-Munn, T J, Morrell, S, Morrison, R D, Kojovic, T, 1996, Mineral comminution Circuits, Julius Kruttschnitt Mineral

Research Centre, University of Queensland, Brisbane.

Pax R, 2001, Non-contact acoustic measurement of in-mill variables of SAG mills, in Proceedings International

Conference on Autogenous and Semi-Autogenous Grinding Technology (SAG 2001), 2, (eds: D Barratt et al)

University of British Columbia, Vancouver BC, Canada, 2001, pp 386-391.



IsaMill™ Design Improvements and Operational Performance at Anglo Platinum

C Rule1 and H de Waal2

1. Head of Concentrators Technology, Anglo American Platinum, 55 Marshall Street, Marshall Town, Johannesburg, South Africa. Email: [email protected]

2. Consulting Metallurgist, Xstrata Technology, Level 4, 307 Queens Street, Brisbane Qld 4000. Email: [email protected]

ABSTRACT In 2003, Anglo Platinum, in a joint development with Xstrata Technology, installed the world’s first 10 000

litre IsaMillTM in a concentrate regrind duty at the Western Limb Tailings Re-treatment Plant. The success of

that installation was the enabling event for Anglo Platinum to proceed with a substantial investment in

horizontal stirred milling technology. Since 2006 an additional three IsaMillsTM in concentrate regrind duties

and a further 18 IsaMillsTM in the more technically challenging coarse grinding mainstream applications,

have been commissioned in group Concentrator operations - bringing the total number of IsaMillsTM installed

in Anglo Platinum plants to 22 .

A collaborative approach between Anglo Platinum and Xstrata Technology towards improving milling

efficiency and reducing operating costs, through internal mill component wear optimization and operating

recipe development, has resulted in further improvements in the overall success of IsaMillsTM in the flow

sheets of many Anglo Platinum operations. The addition of IsaMillsTM in the Anglo Platinum flow sheets has

improved plant PGM recoveries by as much as 5%.

This paper explores the improvements made to the IsaMillsTM flow sheet and mill internal design and shares

some of the operating experience with IsaMillsTM technology in Anglo Platinum

INTRODUCTION IsaMillTM technology has been applied to a number of different ore types and process applications in an

ongoing drive to liberate valuable minerals from complex ore bodies.

The need for ultrafine grinding arose at McArthur River mine in the Northern Territory of Australia where

finely grained lead/zinc concentrates required a grind size P80 of 7 microns to produce a saleable product.

The high energy intensity and inert grinding benefits of this unique stirred milling technology further

contributed to producing a suitable final concentrate (Barns et al., 2006).

Anglo Platinum pursued stirred milling technology as an economically viable option to improve liberation

following successful scoping work on bench scale initially using a 4 liter bench scale IsaMillTM. Thereafter a

series of off-site and onsite pilot scale tests were conducted which ultimately led to the installation of a large



scale 10 000 liter mill at their Western Limb tailings retreatment project. This mill has subsequently been

deployed in a number of grinding duties with varying coarseness of feed, allowing investigation on the wear

rates of mill internals (Rule C M, 2010).

Further requirements for liberation of finely grained platinum group metals (PGM’s) in the main stream

process flow resulted in IsaMillsTM allocated to a coarser grind duty as tertiary grinding applications,

producing a pre-scavenging flotation product at an approximate P80 of 53 microns (µm).

The coarse grinding applications typically receive feed at F80 of up to 100 µm although coarser feed sizes of

ten – 15 % >150 µm has been observed in certain cases (Rule et al.,2010)

The coarser feed sizes have been shown to impact negatively on wear performance of grinding discs and

wear of shell liners to a lesser extent.

To date Anglo Platinum has installed 17 IsaMillsTM in main stream inert grinding (MIG) applications and 4

units in an ultrafine concentrate regrind duty (UFG).

Since the installation of the first MIG mill a number of improvements have been made to subsequent flow

sheet designs and mill internal components to optimize the component life and value gained from this

technology.

This paper explores some of these changes and highlights the benefits realized and value gained from the

collaborative efforts of Anglo Platinum and Xstrata Technology towards unlocking the true potential of

IsaMillTM technology in the PGM industry.

DESIGN IMPROVEMENTS

The majority of IsaMillsTM in Anglo Platinum have been installed in MIG grinding applications where the

mills receive feed from a secondary or tertiary ball mill. IsaMillTM product will then report to a scavenging

flotation stage to capitalize on the improved liberation state and flotation kinetics of the ground pulp.

Due to the inherent variability of the upstream operating plant; i.e. highly variable mainstream slurry flow

rate and slurry particle size distribution, the need arose to modify the process flow sheet and mill internal

component design parameters to better adapt to this variability. Media handling systems were also improved

to facilitate faster reloading of media after maintenance inspections with a reduction in mechanical wear on

this equipment and an improvement in spillage generation and media accounting.



Constant Flow Concept

The internal classification functionality of the IsaMillTM utilizes a product separator at the discharge end of

the mill to centrifuge coarse material to the outer circumference of the mill and transport this material back

towards the feed end of the mill.

Fig. 1. IsaMillTM internal classification

Fig. 1. depicts the internal configuration of the IsaMillTM detailing the classification function performed by

the product separator (rotor).

Due to the pumping effect of the product separator towards the feed end of the mill a certain degree of

compression is generated towards the feed end of the IsaMillTM. Slurry feed into the mill counteracts this

back pressure to provide a net positive flow through the mill. Conventional PI level control in the IsaMillTM

feed tank cascades the change in feed tank level onto a flow rate set point for the feed pump. A reduction in

volumetric flow rate into the IsaMillTM circuit due to upstream process changes will result in a reduction in

flow rate into the IsaMillTM and subsequent increase in compression towards the feed end of the IsaMillTM.

This concentrates the grinding action towards the front end of the mill with the majority of the grinding

chambers not contacting with a slurry / media mixture, and promoted uneven and accelerated wear rates of

mill grinding discs particularly towards the feed end of the IsaMillTM.

A constant flow concept was subsequently devised to fix the flow rate set point of the IsaMillTM feed pump

and return a portion of the milled product to the feed tank, thus maintaining feed tank level set point while

achieving a constant flow. The original valves in the discharge line was used during the starting and shut

down procedures, these were replaced by ceramic lined variable split range control valves during the

modification. The recycled portion should however be maintained to a minimum and advanced process

controllers were utilized to maintain the recycle portion to a minimum as shown in Fig. 2.



Fig. 2. – IsaMillTM recycle flow control

Fig. 3. – Waterval UG 2 A-Section IsaMillTM Flow Control Impact



The result of the new control system can be seen in Fig. 3. where the stability of feed flow increased

drastically after the modifications to the recycle flow control were made to this mill. The operating recipe

was also altered to achieve the best media distribution and grinding efficiency, thus the reduction in slurry

feed flow rate set point to 180 m3/h. The average flow rate achieved before re-cycle flow control was

between 200 and 250 m3/h, the variability in flow rate during normal plant operation can result in drastic

changes in feed flow rate to as low as 50 m3/h as displayed in Fig. 4. The modification to recycle flow

control has proven to be significant during periods of low feed flow to the IsaMillTM circuit.

Fig. 4. – Mogalakwena South IsaMillTM Feed Flow Rate

All MIG installations in Anglo Platinum have subsequently been fitted with this modification with similar

improvements in flow rate stability realized.

Reduced Diameter Discs (RDD’s)

Case Study One – Mogalakwena South Concentrator disc trial The major wear component inside the MIG IsaMillsTM in Anglo Platinum is the grinding disc which enables

the agitation of media for attrition grinding between agitated media and solid particles in the slurry stream.



The original grinding discs at 1720 mm in diameter showed some degree of wear particularly towards the

feed end of the mill. The high tip speeds and distance between the disc outer circumference and mill shell

resulted in accelerated wear of the discs and shell liner.

The MIG installations in Anglo Platinum have displayed the most aggressive wear rates of all main stream

installations worldwide. Installations in a main stream copper processing plant in Australia have shown much

lower wear rates at similar specific energy consumptions as the Anglo Platinum installations (refer to Fig. 8.)

This could be attributed to the specific mineralogy of the reefs in the Bushveld complex mined for its PGM

content. Many Anglo Platinum operations treat Upper Group 2 (UG2) reef known for its chromite spinel

content, which comprise about 70 % of the Chromitite reef. The angular crystal structure and high specific

gravity of chromite can contribute to higher wear rates on grinding discs if allowed to accumulate inside the

IsaMillTM. Chromite content in the mill can be controlled by ensuring that coarse chromite is treated

separately from the IsaMillTM circuit or that the size fraction of chromite reporting to the IsaMillTM circuit

has been sufficiently reduced to allow treatment through the IsaMillTM. Typical F95 sizes suitable for

treatment through the mill are approximately 110 - 115µm. This grind can normally be achieved with

conventional ball mills as a pre-IsaMilling stage.

Merensky and Plat Reefs also form part of the PGM rich deposits in the Bushveld Complex.

Less aggressive component wear rates have been reported for the mills deployed in a MIG duty on these

reefs although the pyroxenite gangue rock associated with the Merensky reef and the weathering alteration

effect on Plat Reef are known to be some of the hardest ore types to beneficiate.

The main wear dimension on the grinding discs was in thickness wear with much lower wear rates recorded

in disc diameter. Disc replacement requirements are mainly as a result of width or thickness wear where the

steel reinforcing inside the rubber moulded discs are exposed after the majority of rubber has worn away,

thus necessitating the replacement of a disc. .

The first IsaMillsTM where reduced diameter discs were installed was at Mogalakwena South’s C Section

IsaMillTM. This mill was installed in a tertiary grinding application receiving feed from the secondary ball

mill product stream. Slurry entered the IsaMillTM circuit at a P80 of 75 µm producing a product at P80 53

µm.



Fig. 5. – Grinding disc wear rate comparison

Fig. 5. show the reduction in disc thickness wear after changing to reduced diameter discs from normal

discs. An immediate reduction in wear rates can be observed in grinding discs 1 and 2.



Fig. 6. – Mogalakwena South C Section Final Grind

Of additional interest is the effect of slurry feed particle size distribution on disc thickness wear. The

IsaMillTM was employed in secondary regrind duty for a period when the secondary ball mill was off line for

major repairs. During this period the feed particle size distribution to the IsaMillTM circuit increased due to

the absence of the secondary ball mill in the circuit, with a consequential increase in disc wear rates.

Fig. 6. displays the reduction in final grind during the period while the secondary ball mill was off line. No

process data was available for IsaMillTM feed particle size distribution but the final tailings grind clearly

shows a reduction in grind during the period in question. The operating recipe of the IsaMillTM remained the

same with no changes to grinding media type or size and the same specific energy (kWh/ton) throughout.

Once the secondary ball mill was returned to duty the disc width wear rates reduced although not to previous

rates observed when originally switching to reduced diameter discs. This could be attributed to a change in

the pebble crushing circuit preceding the grinding circuit in C Section, where the cut point of final product

from the crushing circuit was increased to capitalise on the expected additional grinding capacity with the re-

introduction of the secondary ball mill into the circuit.



Case Study 2 – Western Limb Tailings Re-treatment Concentrator reduced diameter disc performance

The IsaMillTM at the Western Limb Tailings Re-treatment (WLTR) Concentrator was originally installed as a

concentrate regrind mill. The WLTR plant processes PGM tailings material from historical dumps in the

greater Anglo Platinum Rustenburg concentrator operations area.

This mill was later converted from using sand as grinding media to ceramic beads to allow treatment of

coarser slurries in main stream grinding applications.

Initial disc width wear while in concentrate regrind duty was between 0.19 – 0.02 µm/kWday. See Fig. 7.

Fig. 7. – WLTR Concentrator comparative disc wear rates



Disc wear rates increased rapidly to 2.44 µm/kWday when the mill was deployed as a primary grinding unit

treating ore in a parallel side stream to the primary ball mill. Reduced diameter discs were then installed on

discs one and two with normal diameter discs remaining in the rest of the mill. Disc width wear rates reduced

on the first two positions with an expected increase in wear rate on disc three due to a shift in the media

distribution profile towards the discharge end of the mill. In essence the discs after the first two discs started

working a bit harder and this was illustrated in the increase in wear rates on disc three.

Comparative wear data

Fig. 8. – Comparison between PGM and Copper reef disc wear rates

Disc width wear rates in the majority of MIG installations in Anglo Platinum have shown significant

differences when compared with “softer” ore types. A typical MIG installation on a copper ore as trended in

Fig. 8. shows the difference in wear rates between this IsaMillTM and Mogalakwena C Section IsaMillTM.

Both these mills operate on very similar specific energy consumptions with similar operating recipes.



IsaChargerTM hydraulic media transfer

Fig. 9. – IsaChargerTM unit

The IsaMillTM require that the complete grinding media charge be removed from the mill prior to an internal

inspection. A media hopper is situated underneath the mill shell to contain media scuttled from the mill.

The media hopper is also used to store new media required to replenish the mill charge as the media is

consumed as part of the grinding process. A screw feeder was originally used to transfer ceramic media from

the media hopper into the IsaMillTM feed tank, from where it is pumped into the IsaMillTM as a slurry/media

mixture.

The screw feeder consisted of numerous moving mechanical parts like gland seals, greased bearings etc.

which required routine maintenance interventions to ensure high equipment availability was maintained.

Xstrata Technology subsequently developed a hydraulic transfer device called “IsaCharger” to transfer

media from the media hopper to the feed tank. The IsaChargerTM consists of no moving mechanical parts and

utilizes a custom built high pressure venturi type device to transfer media by means of a high powered jet of

water from underneath the media hopper into the mill feed tank. See Fig. 9.

Most IsaMillTM installations in Anglo platinum have been fitted with IsaChargersTM with reported

improvements in equipment on-line and similar media recharge times as achieved by the original screw

feeders.



To date no maintenance interventions have been required on any of these units installed on the Anglo

Platinum IsaMillsTM. Some of the units have been in operation for up to 6 months.

OPERATIONAL PERFORMANCE

Case Study – Amandelbult Concentrator UG 2 # 2 Plant

Fig. 10. – Amandelbult UG 2 IsaMillsTM showing MIG and UFG concentrate regrind mills.

Table 1 – IsaMillTM installations at Amandelbult

Plant Number and duty Total Installed Power (kW)

Commissioning Date

Merensky 2 Parallel Tertiary MIG M 10 000

6 000 April 2009

UG 2 no. 1 1 Tertiary MIG M10 000 3 000 April 2009

UG 2 no. 2 1 Tertiary MIG M 10 000 3 000 March 2009

UG 2 no. 2 1 UFG, M 3 000 1 500 March 2009

Table 1 list the IsaMillsTM installed at Amandelbult.

The Merensky IsaMillTM circuit comprises two 10 000 liter mills treating Merensky reef exclusively while

the two UG 2 plants combined contribute approximately 57 % of the total production throughput at

Amandelbult.

The importance of UG 2 reef as a contributor towards the total PGM ounces produced at Amandelbult is

apparent and optimizing the PGM recovery for this reef type is vital.



The UG 2 # 2 plant IsaMillTM ties into the plant flow sheet as a tertiary regrind mill on the silicate stream.

Chromite is removed from the process stream by means of cyclones after the primary roughing flotation

stage. The chromite stream reports to a dedicated tumbling mill regrind circuit while the silicate stream is

reground in a secondary ball mill prior to reporting to the IsaMillTM circuit. Fig. 11. illustrate the basic flow

sheet of this circuit.

Fig. 11. – Amandelbult UG 2 # 2 Plant basic flow diagram.

This case study on the Amandelbult UG 2 # 2 plant explores the performance of the IsaMillTM and quantifies

the benefit realized after commissioning of the IsaMillsTM

UG 2 Plant Mineralogy Table 2 details the mineral association in composite samples from the plant pre-IsaMill installation.

A significant portion of PGM’s are enclosed in gangue as a result of incomplete liberation, in particular the

> 53 µm fractions in the final tailings sample comprise 57 % of mineral deportment in tailings in that size

fraction.

The liberation issues on the final tailings streams could be addressed through main stream inert grinding with

ultrafine concentrate regrinding of attached mineral particles in the sub 25 µm size ranges.



Table 2 – Mineral Association table for typical Amandelbult UG 2 process samples

Association Feed Concentrate Tailings Tailings <10 µm

Tailings >10 µm

Tailings >53 µm

Liberated 49.2 53.1 31.3 82.3 18.5 2.4

Enclosed in BMS* 23.6 15.8 4.7 4.1 9.7 1.6

Attached to BMS 7.9 12.7 0.3 0.4 0.6 -

PGM/BMS/Silicate 5.6 6 7.7 - 5.6 15.5

Enclosed in Silicate 7.5 8.4 36.0 2.7 44.3 57.0

Attached to Silicate 0.6 2.7 9.3 3.5 13.9 6.8

Enclosed in Oxide 4.8 1.3 7.6 4.3 6.0 11.7

Attached to Oxide 0.8 - 3.1 2.7 1.4 5.0

TOTAL 100.0 100.0 100.0 100.0 100.0 100.0

Midlings 7.2 14.6 7.0 8.3 14.8 -

Locked 43.6 32.1 61.7 9.4 66.7 97.6

*BMS – Base metal sulphides

Grind Performance The grind performance achieved by the MIG IsaMillTM on the UG 2 # 2 plant showed significant reduction in

the coarser size fractions (>150 µm) with a marked improvement in IsaMillTM product towards the end of

2010. This improvement was mainly due to an increase in mill power draw from 1500 kW to 2000 kW as

illustrated in Fig. 12.

Fig. 13. illustrate the increase in finer size fractions passing 53 µm primarily as a result of the increased

power draw. This result is vitally important for the UG 2 operations as a significant proportion of PGM’s

was locked in the > 53 µm fraction of the final tailings plant composite samples. Ref. Table 2.



Fig. 12. – Amandelbult UG 2 coarse fraction size reduction

Fig. 13. – Amandelbult UG 2 fine fraction size reduction



Recovery benefit

Fig. 14. – Amandelbult UG 2 fine fraction size reduction

Fig. 14. shows the historical tailings grades achieved in the two UG 2 plants as well as the contribution of

these two process streams towards the final tailings grade of the concentrator complex. The IsaMillsTM were

commissioned in March and April 2009 and after a period of grinding and flotation circuit optimization, the

tailings grades shows a marked improvement with tailings grades averaging between 0.5 and 0.6 g/t PGM

and gold.

The reduction in final tailings grades on the other Anglo Platinum operations where IsaMillsTM have been

installed has shown similar trends as in the Amandelbult UG 2 # 2 plant.

CONCLUSIONS

IsaMillTM horizontal stirred mills have been successfully introduced in the majority of Anglo Platinum’s

operations in South Africa. Further advances made in wear component design and circuit layout have further

improved the results obtained from these mills from a component wear and grinding efficiency perspectives.

As an illustration to the importance of PGM recovery in the Concentrator Operations in Anglo Platinum, a

one % increase in PGM recovery in the concentrators equates to approximately 75 million US$ at recent

market prices and exchange rates. The following extract from the Anglo Platinum Annual report for 2009

puts further emphasises on the magnitude of the recovery improvements directly contributed to IsaMillsTM

UG 2 # 2 IsaMillTM

Commissioned



“Early indications during the latter part of 2009 were promising; for example, platinum metal recoveries

for the last quarter of 2009 at Rustenburg increased substantially post IsaMill™ commissioning by in excess

of 3 percentage points.”

ACKNOWLEDGEMENTS

The authors would like to acknowledge Anglo American Platinum for permission to publish data from their

concentrator operations in this paper.

REFERENCES

Barns K E, Curry D C, Pease J D, Rule C M and Young M F, 2006. Transforming flowsheet design with

inert grinding – the IsaMillTM, paper presented to CMP conference 2006.

Rule C M, 2010. Stirred Milling – new comminution technology in the PGM industry, paper presented to

the 4th International Platinum Conference, platinum in transformation “Boom or Bust”, 2010, (The

South African Institute of Mining and Metallurgy)

Rule C M, Knopjes, L, Jones, R.A, 2010. The introduction of main stream inert grinding, or “MIG”

IsaMilling technology at Anglo Platinum, paper presented to Comminution 2008 conference, 2008.

Anglo Platinum Annual Report, 2009, page 85.



The Influence of Liner Wear on Milling Efficiency

P Toor1, T Perkins2, M S Powell3 and J Franke4

1. MAusIMM, Analytical Metallurgist, Scanalyse Pty Ltd. Email: [email protected] 2. Research Assistant, JKMRC University of Queensland, [email protected] 3. FAusIMM, Chair in Sustainable Comminution, University of Queensland, JKMRC.

Email: [email protected] 4. Chief Technical Officer, Scanalyse Pty Ltd. Email: [email protected]

ABSTRACT

Due to the high energy consumption of milling, it is desirable to improve the efficiency of this process. One

alternative to reduce energy and material consumption while optimising production yield in mill operation is

through balanced design and selection of liners. Over-design of liners leads to increased life but at a cost to

grinding performance, a relationship which to date has not been studied closely. This paper provides a

description of this process of liner life cycle optimization and presents initial results from the first of two life

cycles being studied for a 32 foot (ft) SAG mill. The use of commercially proven high resolution laser

scanning based liner shape information provided by the MillMapper® software correlated with well

controlled site surveys at various points in the liner life was used as the basis for comparison. The survey

data was modelled in JKSimMet to provide direct comparisons of mill performance. This potentially

provides a methodology for the design of liners that maintain a favourable shape for the majority of liner life,

yielding desirable production parameters such as lower specific energy consumption and higher throughput,

while at the same time maintaining practical liner life cycles.

INTRODUCTION A key area of interest in minerals processing has been research into the establishment of grinding processes,

and more recently, of representative grinding simulations. This is due to grinding being an important unit

operation in a processing plant, both in terms of energy consumption and overall performance. It is not

uncommon for grinding circuits to constitute up to 40% of the plant power usage and operating cost in a

processing plant (Herbst et al, 2003).

The interest in research into and simulation of grinding and comminution in general has been triggered by

many current issues such as those listed by Norgate et al, (2010):

The increasing cost of fuel and materials. The comminution of ore does not only consume

significant amount of energy but also large amounts of steel due to wear of grinding media and

liners.

Climate change is now a most pressing global environmental issue and has increased the need to

make comminution more energy efficient beyond purely economic arguments. Milling uses about



90 % of comminution or 40 % of plant energy. In the case of copper ore, the grinding stages form

the largest contribution to the total greenhouse gas emissions for the production of copper

concentrates.

The general trend of falling grade of ore bodies which are more complex and finer grained. This will

require finer grinding to liberate the valuable material to achieve separation and concentration and

inherently make comminution processes more energy intensive.

The ever increasing global demand for industrial and consumer goods means that the production of

primary metals can be expected to increase well into the future.

Many research and industry practitioners are working on addressing the energy issue in comminution with

new equipment designs or novel circuit solutions using existing equipment. These approaches can overcome

the inherent limitations in the existing grinding equipment designs and circuits that cause high energy

consumption. Whilst they are expected to provide solutions for new projects, it is important to note that

existing plants will continue to run with traditional grinding circuits because of the capital investment costs

already incurred. It will not be commercially viable to refit more than a small percentage of existing plants

with the new equipment and circuit designs. Therefore solutions to existing traditional grinding circuits are

needed if progress is to be made in the industry as a whole.

One possible way to reduce energy and material consumption in milling is through appropriate design and

selection of liners. Liners play an important role in grinding due to their strong influence on load motion and

behaviour (Makokha et al, 2006), (Powell et al, 1993). Generally the milling efficiency depends on the

behaviour of the load inside the mill which governs the nature of the ore presentation to breakage (Makokha

et al, 2006), yet liner shape is dynamic as it is continually changing because of wear.

It is a common practice in the minerals processing industry to attempt to overdesign mill liners with the aim

of achieving a maximum possible liner life. Overdesigning liners is a trade-off between liner life and

grinding performance which to date has not been studied closely. A common observation over the life of a

shell liner life cycle is an unplanned but welcome throughput increase and specific energy (kWh/t) decrease

as wear progresses. Without the use of appropriate liner measurement tools to track and quantify wear, an

intentional increase in throughput and decrease in power draw with loss of liner profile as the conclusively

identified root cause is difficult to validate.

The authors’ own observations of a number of mills show an increase as great as ten percent (10%) in

throughput from new to old liner, Fig. 1 shows an example for a large SAG mill. It illustrates that there is a

trend of increased throughput over shell liner life. Since mill speed increases as shown were used to

compensate for the decrease in lift as the liners wear, it is likely that the highlighted increase in throughput

over shell liner life is due to both an increase in internal mill volume as the liner wears, as well as an

optimum liner shape being achieved during liner life.



Fig. 1- Increase in throughput over liner life for a large SAG mill.

Unfortunately however, the 32 ft SAG mill that was studied for this project does not feature recent recorded

evidence of a throughput increase trend over shell liner life. Throughput and power draw curves for the 32 ft

mill over three life cycles are shown in Fig. 2. The ore supply to this mill is the current operational

bottleneck in the circuit and is the main reason for the observed steady production tonnage and for not

utilising mill speed to compensate for loss in lift caused by liner wear. This in turn makes it more difficult to

identify optimum mill and production parameters during liner life from production data alone. Though the

system contains more statistical noise than hoped for it is believed that an optimum liner shape and running

conditions can still be recommended due to the robustness of the modelling techniques. The processing

methodology in this case is to identify an optimum throughput window not based on observed throughput

behaviour, but by normalizing the throughput data in JKSimMet to select a window of optimum breakage

rate over the liner life cycle. Modelling techniques may also include the use of the Discrete Element Method

(DEM) Simulations. Variation in the feed accounts for much of the associated noise in a processing plant.

The JKSimMet SAG model can be used to normalise the different conditions to a constant base case of ore

competence, feed size, mill filling, etc., thus compensating for the large variation in feed witnessed over a

liner life. DEM simulations will be used in a similar manner to be able to make one to one comparisons

between liner shapes. It should be noted in this study only data from the 32 ft SAG mill is analysed.



Fig. 2- Throughput over liner life for 32 ft case study SAG mill.

The project work forming the basis for this paper is scheduled for two full liner life cycles. The first liner

life cycle was completed in December 2010; the second life cycle was completed in June 2011. Data from

the second life cycle will be weighted more heavily due to improvements to sample points made at site thus

increasing the confidence in the second data set. During the first life cycle changes to the design of the

discharge grates were made adding to the noise of the data set. This paper presents the preliminary results

from the first life cycle with data from the second life cycle not complete at the time of writing.

Four grinding surveys were conducted during the liner life alongside MillMapper® laser scanning of the mill

liners to obtain liner shape, ball size distribution, liner volume and liner mass (suited for load cell calibration)

and ball/ore charge volume information. The survey protocol is given in Table 1.

Table 1-Survey Protocol.

Date Survey Notes Liners Relined

05/07/2010 Survey 1 (Worn) Data collection at the end of previous liner life prior to reline.

N/A

08/07/2010 Survey 2 (New) Data collection immediately after reline.

Shell, Half Set Discharge Grates

1/10/2010 Survey 3 (Mid Life) Data collection at approximately mid life.

Half Set Discharge Grates, Inner Discharge Head, Feed End Outer

24/11/2010 Survey 4 (Highly Worn) Data collection 2 weeks prior to reline. N/A

06/12/2010 Survey 5 (Fully Worn) Data collection prior to reline. N/A



PROBLEM STATEMENT - LOW EARLY LINER LIFE THROUGHPUT The hypothesis of this work is that the increase in throughput in SAG milling over a shell liner life cycle

witnessed in industry is due to both an increase in internal mill volume due to liner wear, and secondly, to an

optimum liner shape being achieved during liner life.

Increase in internal mill volume over liner life cycles is easily tracked from the MillMapper® software as it is

a standard survey output. The net mill volume for each survey is calculated by accounting for the exact

laser scanned 3D liner shape at the wear stage in question for feed, shell, and discharge liners. Traditional

mill volume estimations are based on manual single point distance measurements that can not account for 3D

shape changes. MillMapper® technology can track changes in mill volume, arising from the wear of shell

lifters, as well as feed and discharge cones at a high degree of accuracy.

In the case of the large SAG mill referred to in Fig. 1, the mill’s initial volume is 757m3 when new liners are

installed, with the volume on average increasing to 792 m3 at the end of shell liner life over three liner lives

for the 2010 calendar year, which equates to an average increase of 4.4%. However, historical data

demonstrates the period from new to old liner stage yields on average an increase in throughput in the order

of 10% (Table 2). Notably, this increase in throughput is achieved for the 6 observed subsequent shell life

cycles to varying degrees, and therefore is a systematic trend rather than an outlier. The reline schedule of

this mill is for the all the shell and feed end to be relined simultaneously. Only the discharge liner is relined

in an intermediate sequence. This routine imposes less distorting influence on the assessment of throughput

over shell liner life cycles. The reline schedule for the large SAG mill is summarised in Fig. 3.

Fig. 3: Reline Schedule for large SAG mill



Assuming constant transport rates and percentage fill levels for new and completely worn liners, the increase

of only 4-5% in mill volume does not account for the observed 10% increase in throughput. Lacking other

conclusive evidence that could explain any potential change in transport rate or fill level, the remaining 5-6%

in observed increased throughput may be attributed to an increase in efficiency of liner shape as it wears.

The data suggests that new liners as used are most inefficient and at the other end of the scale, worn liners

are most efficient. This paper outlines a robust technique to capture and quantify this optimum liner shape

for improved milling efficiency. This can inform the design of liners that maintain a favourable shape for the

majority of liner life, yielding desirable production parameters such as lower energy consumption and higher

throughput, while at the same time maintaining practical liner life cycles and not compromising product

sizes.

Table 2- Percent Increased Throughput from Old to New Shell Liner

New Shell Liner Very Worn Shell Liner Increased Throughput %

(Worn-New)

Period Tonnage Period Tonnage

23/04/2007-07/05/2007 1939 11/08/2007-25/08/2007 2154 11.1%

15/09/2007-29/09/2007 1912 18/01/2008-1/02/2008 2183 14.2%

14/02/2008-28/02/2008 1972 14/06/2008-28/06/2008 2126 7.8%

07/07/2008-21/07/2008 1939 30/10/2008- 13/11/2008 2267 16.9%

25/11/2008-09/12/2008 2181 19/03/2008-2/04/2009 2242 2.8%

03/04/2009-17/04/2009 2081 12/08/2009-26/08/2009 2268 9.0%

Average Average Average

2004 2207 10.3%

METHODOLOGY This section provides a methodology for the design of liners that maintain a favourable shape for the

majority of liner life, yielding desirable production parameters such as lower energy consumption and higher

throughput, while at the same time maintaining practical liner life cycles.

To aid in determining this optimum liner shape the following tools were used:

Controlled Grinding Surveys;

MillMapper®;

JKSimMet.



It is also possible to add Discrete Element Models (DEM) at four different liner life stages based on the

survey work and MillMapper® scans. Furthermore it is also possible to develop DEM for appropriate “What

If” scenarios aiming at liner redesign.

Below (Fig. 4) is a conceptual methodology for testing the hypothesis.

Fig. 4- Conceptual methodology of project.

Grinding surveys are conducted at various stages in the liner life (4 times), to gain snap shots of milling

efficiency at each stage over two life cycles. In conjunction with the grinding surveys MillMapper® provides

detailed and accurate 3D liner shapes at each stage.

Due to the large statistical noise associated with commercial processing plants, JKSimMet is used to filter

noise in the grinding surveys by normalizing feed and operating conditions and only leaving the breakage

rates as the variables in the steady state models.

JKSimMet model results are used to identify the most productive operating period of the life cycle for the

present liner design before correlating them to the corresponding MillMapper® liner shapes. With knowledge

of optimum liner shape, a new liner design can be proposed which maintains a favourable liner shape for the

majority of life. The functionality of the new liner design can be further tested in a subsequent DEM

simulation.



Conduct controlled grinding surveys in conjunction with accurate laser scanning during the course of

a mill liner life to make direct comparisons between various liner shapes and their impact on mill

performance.

Measure the liner shapes and mill fill volume levels at four (4) times during the life cycle using

Scanalyse’s established laser scanning technique.

Conduct carefully controlled surveys to measure mill performance directly prior to each liner

measurement.

Use mill surveys, mill performance and production data to calculate key mill performance

parameters over the life of the liner. These parameters include:

o Throughput,

o Grind size – fineness of product desired to feed downstream processes, and

o Specific energy consumption.

Use JKSimMet to normalise the data to keep parameters such as feed hardness and cyclone

performance constant across all surveys to highlight the effect of change in liner shape on mill

performance.

LINER SHAPE TRACKING MillMapper® provides high resolution three-dimensional mill liner thickness information inside a mill by

mapping tens of millions of individual survey points. Aside from tracking liner shape and wear, other

variables such as liner weight, net mill volume, charge volume, ball size distribution, and discharge grate

open area are also quantified. After completion of a condition monitoring series, high wear zones are

automatically identified and used for reline forecasting, as illustrated in the liner thickness plots of Fig. 5.

More detailed descriptions of the MillMapper® methodology and industry case studies have been provided in

Franke et al, 2006 and Franke, 2008. Variables such as effective lifter height and face angle which

significantly affect the trajectory of the charge are measured and provide the basis for representative DEM

modelling. Critically this can be done not just for new liner shapes, but for worn shapes as well, which is the

only way to identify actual grinding behaviour at times other than the very initial period of a liner life cycle.



Fig. 5-3D Thickness model as produced by MillMapper®.

Fig. 6 - Highest wearing profile progression at a central mill section over liner life

Fig. 6 shows the stages of the liner wear when undertaking the surveys. Each line represents the liner profile

measured at a given survey date with the darkest line being the most recent (06/12/2010 reline) and the

lightest being the most dated (08/07/2010 new liner). Peak performance periods therefore can be correlated



to a distinct liner shape. Whilst Fig. 6 shows the specific 2D profile of the highest wearing zone on the shell

liners, the identification of optimum new liner shapes needs to be considered in 3D when assessing all liner

areas. This is relevant for all cases where liner wear varies significantly along the mill axis which is true for

the vast majority of SAG and AG mills in operation (Franke, 2008).

PLANT SURVEYS The SAG circuit being studied is shown in Fig. 11. The SAG is in closed circuit with cyclones and the feed

material is typically minus 75mm. This material is a combination of crushed RoM material and a recycle

stream of crushed SAG mill pebbles. The SAG mill discharges onto a trommel, the coarse material from the

trommel is then passed over a single deck vibrating screen before being conveyed to the pebble stockpile.

The pebble stockpile material is fed to one or both of the cone crushers which discharge onto the SAG feed

belt.

The SAG trommel undersize flows into a common sump shared with the ball mill. The material in this sump

is pumped to a cluster of cyclones for classification. Portions of the cyclone underflow stream are used to

feed the flash flotation cell and the gravity concentrators, with the remainder of the material being fed into

the SAG, ball and verti-mills. The gravity concentrate reports directly to the gold room and the cyclone

overflow forms the feed to the flotation circuit.

In order to properly sample the mill product, samples were taken from the SAG trommel undersize, the SAG

screen undersize, and the SAG screen oversize. The full feed to the mill was sampled through a SAG feed

belt cut and sampling of the cyclone underflow streams. The feed and product samples from the recycle cone

crushers were sampled directly at the conclusion of the first survey. The cyclone overflow was sampled to

provide circuit product information and allow an estimation of the effect of the reline on the total circuit

performance to be determined.

Samples were taken around all relevant points with the goal of fully specifying the inputs and outputs of the

SAG mill, Recycle Crushers, and cyclones. A summary of the sample points is given below.

1. SAG Feed;

2. Recycle Crusher Product;

3. Recycle Crusher Feed;

4. SAG Pebble Recycle;

5. SAG Trommel Undersize;

6. SAG Screen Undersize;

7. Cyclone Underflow;



8. Cyclone Overflow;

9. Cyclone Feed.

Of the above samples the SAG Trommel undersize (US) sample is inarguably the most crucial sample point.

The sample from this point was initially shown to be non-representative. To test the repeatability of the SAG

trommel sample point, three samples were taken immediately following each other while the circuit was

relatively stable. The samples were sized and produced the results as shown in Fig. 7.

Fig. 7- Size distributions of repeat SAG trommel undersize samples. Original sample point

It can be seen there is a large amount of variability in the sample point. The modifications required to

improve the sample point were not possible to implement without a major shutdown, thus this refit was not

completed until the second life cycle.

The original sample was taken through the door in the trommel casing shown in Fig. 8. And collected with a

very long arm sample cutter. The capacity of the sample cutter was enough to ensure that overflowing was

not likely, but the reach of the sample cutter was not enough to allow the entire stream to be sampled. The

highest flow section of the stream was not reachable, introducing a strong possibility of bias in the sample.



Fig. 8- SAG trommel undersize door sample point

During a reline shutdown, the opportunity was taken to improve the sample points by cutting three hatches

on the opposite side of the trommel casing closer to the high flow stream from the trommel. The hatches

were spaced such that the entire length of the stream could be sampled with a short arm sample cutter with

each hatch allowing for one third of the sample/stream to be cut. The new sample point is shown in Fig. 9.

Fig. 9- New SAG trommel undersize point with multiple hatches.



The new sample point allowed for a representative sample to be taken. As it can be seen in Fig. 10, the

improvements resulted in the distributions in repeat sampling being virtually identical. Additionally, the

distribution is different to the previous reproducibility test, indicating a likely bias in the original sample.

The new sample point installation will allow more reliable and accurate modelling for the second life cycle.

Fig. 10- Size distributions of repeat SAG trommel undersize samples after installation of improved sample point

JKSIMMET MODELLING

JKSimMet is a steady state simulation program for comminution and classification circuits which is widely

used in industry. The JKSimMet circuit flow diagram used in this project is given in Fig. 11.

Fig. 11- JKSimMet Model Flow Diagram



JKSimMet was used in this project to compare breakage rates over the liner life. This was done by

performing well controlled grinding surveys at four stages of liner life and creating subsequent models. To

make direct comparisons between surveys and gauge the effect of change in liner shape, the data was

normalized. This was done by holding other variables constant such as the ‘A’ and ‘b’ ore impact breakage

parameters from the JK drop weight test. By doing this, any change in the breakage rates and or mass

transfer can be attributed to the change in liner shape rather than changes in feed or operating conditions.

Table 3 presents the model parameters for the surveys. Actual differences in the data from those listed in this

table are as follows.

The diameter varies throughout the liner life and this value was taken from the MillMapper® outputs.

The fine classifying size for the discharge function was defined to be approximately 1.1 mm,

because this was the best overall fit from for all the surveys and should have very little variance

between surveys.

The pebble port fraction varied over the period analysed, so has to change in the models to allow for

this impact on mill hold-up. The fraction used in the model does not represent the true fraction of the

open area made up by the pebble ports – a discrepancy being addressed in separate work on

upgrading the SAG mill model – so is a relative number.

Current values were chosen based on experience and were allowed to vary slightly around their starting

points to determine the stability of the solution. All these best fit values were then locked into the simulation

to calculate the discharge functions for each survey.

Table 3 Mill Data



Fig. 12 SAG Discharge Functions.

Fig. 12 shows the various discharge functions calculated for each survey. The discharge functions

demonstrate a large amount of variance across the liner life. However, the variance shown is predominately

due to a change in the grate design. The fraction of the grate open area made up of pebble ports changed

during the period over which the surveys were conducted. Half the grates were replaced with a new design

between the 05/07/2010 and the 08/07/2010 survey, and the other half were changed before the 01/10/2010

survey. This reduced the fraction of the grate open area made up of pebble ports from 100% during the

05/07/2010 survey, to approximately 72% during the 08/07/2010 survey, to 43% for the rest of the surveys.

This caused the grouping seen in Fig. 12, where the October, November, and December surveys are grouped

together due to the same grate design. The October survey shows a faster discharge rate, opposing the

expected reaction to grate wear. However, the change in the discharge functions due to wear still can be

commented on both grate designs separately.

The progression in discharge function due to the change in grate design can be clearly seen i.e. a step change

from a discharge rate of 56hr-1 to 48 hr-1 for the minus 1mm size fraction. This step change is due to a

reduction in the grate open area fraction from 0.124 to 0.091.When considering the discharge functions for

the new designs (Surveys 3, 4 and 5) it can be seen that there is little difference. This is due to the fact the

grate open area fraction was equal for all three surveys at 0.059. Thus the discharge rates stayed relatively

constant during the three surveys, as expected.



Surveys of the mill during the second liner life will remove the uncertainty about the variance in discharge

function by providing a series of surveys on a mill with the same grate design throughout, and will also allow

the effects of grate wear on mill performance to be quantified.

Fig. 13: Breakage rates of all surveys.

The calculated breakage rates (Fig. 13) will be a useful measure of the change in mill performance with liner

wear but the relationship is too preliminary to draw meaningful conclusions at this stage. There are

differences seen in the breakage rates across the surveys from new and worn but these are susceptible to a

number of sources of potential bias. The largest potential source of bias is the sub-optimal nature of the mill

discharge sample point described earlier. A new sample point has been installed, which has greatly improved

the quality of the sample, thus reducing the potential bias in the samples for the second life cycle and future

surveys. The improvement in the mill discharge sample point will enable the collection of more accurate

data, and will add to the confidence ascribing the change in mill performance to liner wear can be quantified.

At this stage, while the outcomes of the JKSimMet modelling indicate a possible trend across liner life, the

effect is small and it is not clear if it is due to noise in the data.

A preliminary analysis based on the first liner life cycle is possible by comparing Survey 1 (fully worn) and

Survey 2 (new liner) as these surveys offer the greatest step change in liner life and wear. Thus the adverse

effects of plant noise, inconsistent operating conditions and sampling errors are mitigated most when

comparing these data sets. Considering that these surveys were conducted only days apart - one immediately



before reline and the other immediately after - also add to the confidence when comparing survey and

modelling results.

Fig. 14 shows the breakage rates for the two liner conditions very worn and new. The plot illustrates that

there is an increased breakage rate for the size classes below 30 mm for the worn liner. Conversely, the new

liner exhibits an increased breakage rates for particles greater than 50 mm with 40 mm being the

approximate cross over point. This result can be explained by new liners imparting more lift to the charge

thereby causing comparatively high impact breakage of the coarse particles whereas the worn liner causes a

greater cascading motion of the charge. Thus a larger amount of breakage would occur through attrition and

abrasion of the finer material for the worn liner case. The depressed rates for both sets of data at coarse end

is due to the small top-size of the feed material, lacking material in the largest size range, the rates in that

range have little meaning.

Fig. 14- Breakage rates of worn and new liners.

A summary of process variables is given in Table 4 which shows that throughput decreased by 62 tph

(7.76%) and power increased by 559 kW (9.28%) after the reline in accordance with the hypothesis and

problem statement described earlier.



Table 4 Summary of Key Process Variables

Process Variable Very Worn New

Fresh Feed tph 799 737

Recycle Crusher tph 126 67

Total Throughput (tph) 925 804

Power (kW) 6019 6588

Mill Filling (%) 19.3 19.5

When comparing the associated SAG product size distributions (Fig. 15) it is evident that the size

distributions are quite similar with the SAG product of the new liner being finer overall. This being the case

the increased breakage rates shown by the worn liner for the minus 30 particles is most likely due to a greater

rate of production of fines rather than a finer product being produced. Thus the increased overall throughput

can be explained by a greater rate of fines being produced by the worn liner as compared to the new liner.

The increase in power associated with the new liner with the same mill filling is possibly due to the change

in liner profile.

Fig. 15- Product size distribution new (8/07/2010) versus worn (8/07/2010) liner.

FURTHER WORK Further work required in this project is the analysis and creation of JKSimMet models for the second life

cycle. This life cycle will take advantage of the improved sample point and the consistent grate design

throughout the liner life. It is envisaged that the second data set will provide more detailed insight to the

influence of liner wear over the liner life cycle as opposed to solely considering new versus worn liner

shapes.



The subsequent step to quantifying the influence of liner wear on the relevant production parameters will be

to utilize this information in order to improve new liner design. This will be achieved by matching

MillMapper® liner shapes with periods of peak performance after normalizing the influence of all other

relevant parameters using JKSimMet. This provides a methodology for the design of new liner shapes that

maintain a favourable shape for the majority of liner life, yielding desirable production parameters such as

lower energy consumption and higher throughput, while at the same time maintaining desired product sizes

and practical liner life cycles.

Once the desired new liner shapes are established, it is possible to utilise DEM modelling to explore the

effect of the design changes. In the context of this project, the four liner shapes which were captured with

the corresponding grinding surveys can be input into respective DEM models. This provides four distinct

liner shapes producing varying charge motion behaviour and associated output statistics on power draw,

transport of particles, residence times, particle size segregation, impact energies and other variables. These

simulations may be used to provide further verification of the liner shape identified as the optimum new liner

design.

Fig. 16 – MillMapper® thickness model imported into DEM platform.

The DEM also provides an ideal environment for comparative assessments of ‘What If’ scenarios of liner

redesign. The work presented will form the basis for such DEM models.



CONCLUSIONS A methodology for quantifying the influence of liner wear has been provided. The methodology could be

utilised for optimising liner design and reline scheduling. Results from the first life cycle have been

presented with JKSimMet modelling and indicate a possible trend across liner life. An analysis of fully worn

and new liners was carried out as this represented the greatest step change in the mill operating conditions.

Initial results for the mill indicate the worn mill is more apt at producing fines at a higher rate resulting in

increased throughput. The increased throughput is achieved at a lower power, possibly due to a decrease in

high impact energy collisions occurring in a worn mill due to the reduction in lift. This is based on findings

from JKSimMet breakage rates which show that the worn mill has increased breakage rates for particles

below 30 mm.

Results from the second life cycle were not available at time of publication. Data from the second life cycle

will be weighted more heavily due to improvements to sample points made at site thus increasing the

confidence in the second data set. During the first life cycle changes to the design of the discharge grates

were made adding to the noise of the data set.

ACKNOWLEDGEMENTS

The authors would like to acknowledge the support and co-operation of site personnel. Financial support for

this project was provided by AusIndustry through the Climate Ready grant scheme.

REFERENCES Franke, J, Lichti, D D, and Stewart, M P, 2006. MillMapper: A new tool for grinding mill thickness gauging, in Proceedings International Autogenous and Semiautogenous Grinding Technology 2006, Volume III, pp III-75–III-87 (Department of Mining Engineering, University of British Columbia: Vancouver, Canada) Franke, J, 2008. MillMapper experiences – a mill condition monitoring and operational improvement case study, in Proceedings MetPlant 2008, (The Australasian Institute of Mining and Metallurgy: Carlton) Herbst, J.A, Lo, Y.A, Flintoff, B, (2003). Size Reduction and Liberation, in Principles of Minerals Processing (ed: K.N Han, M.C. Fuerstenau) (pp. 61-115). Society for Mining, Metallurgy, and Exploration, INC (SME). Makokha, A.B, Moys M.H, Bwalya M.M, Kimera, K. (2006). A new approach to optimising the life and performance of worn liners in ball mills: Experimental study and DEM simulation. Minerals Engineering , 1439-1445. Norgate, T, Haque, N. (2010). Energy and greenhouse gas impacts of mining and mineral processing operations. Journal of Cleaner Production , 266-274. Powell, M.S, Verneulen L.A, (1993). The influence of liner design on the rate of production of fines in a rotary mill. Minerals Engineering , 169-183.


Energy Efficiency Assessments in Design – Seminal Decisions and Effective Processes

M Allen1, N Rosaguti2, B Innes3 and M Stewart4

1. Senior Consultant and Regional Manager, Energetics Pty Ltd, Level 4, 172 St Georges Terrace, Perth WA 6000. Email: [email protected]

2. Senior Consultant and Group General Manager, Energetics Pty Ltd, Level 7, 132 Arthur Street, North Sydney NSW 2060. Email: [email protected]

3. Senior Consultant and Regional Manager, Energetics Pty Ltd, Level 4, 172 St Georges Terrace, Perth WA 6000. Email: [email protected]

4. MAusIMM, Senior Consultant and Group General Manager, Energetics Pty Ltd, Level 7, 132 Arthur Street, North Sydney NSW 2060. Email: [email protected]

ABSTRACT

Energetics has been in the business of climate change for more than twenty years. This experience has only served to cement the understanding that decisions taken during design define the environmental and economic performance of any project. The 80-20 rule definitely applies, with 80% of a project’s performance being set through the first 20% of design decisions.

This in depth knowledge is used for determining energy price forecasts, future carbon markets and carbon costs, current and future energy technologies, as well as giving a thorough understanding of the processes and technologies in place in the mining industry to work out what the most carbon and energy sensitive decisions during design are.

Project selection and design is a gated process where decisions are taken in line with the information available at that point in the evolution of the project. Design decisions are, by their nature, dominated by uncertainty. In order to ensure that robust decisions are taken, it is necessary to ensure that the correct information is brought to bear at the correct decision point. This then defines and supplies the necessary carbon and energy information for design decisions to be correctly made.

One of the most significant risks facing design engineers and decision makers in the mining industry is the future cost of energy. Part of this risk is associated with a future cost of carbon, however, a significant part of this risk is the cost, and variability in this cost, of the energy itself. This paper investigates the information required to support good decision taking during the design process.

INTRODUCTION

Energetics has been in the business of climate change for more than twenty years. This experience has only served to cement the understanding that decisions taken during design define the environmental and economic performance of any project. The 80-20 rule definitely applies, with 80% of a project’s performance being set through the first 20% of design decisions.


In this paper, the background in depth of knowledge developed over these years on energy price forecasts, future carbon markets and carbon costs, current and future energy technologies, as well as a thorough understanding of the processes and technologies in place in the mining industry, is used to identify the most carbon and energy sensitive decisions during design.

Project selection and design is a gated process where decisions are taken in line with the information available at that point in the evolution of the project. Design decisions are, by their nature, dominated by uncertainty. In order to ensure that robust decisions are taken, it is necessary to ensure that the correct definition and the necessary carbon and energy information for all design decisions is available and brought to bear at the correct decision point. One of the most significant risks facing design engineers and decision makers in the mining industry is the future cost of energy. Part of this risk is associated with a future cost of carbon; however, a significant part of this risk is the cost, and variability in this cost, of the energy itself. This paper investigates the information required to support good decision taking during the design process.

TRADITIONAL DESIGN PROCESSES

Traditionally, the design process for a minerals processing plant consists of a number of distinct steps, each with a different level of confidence in the inputs and requirements for different degrees of certainty in the outputs. Projects typically move from scoping study to prefeasiblity study (PFS) through to feasibility study, definitive feasibility study (DFS), bankable feasibility study (BFS), front end engineering design (FEED), detailed design and finally construction. The inputs to each level of study and the typical outputs are shown in Table 1.

Table 1 - Typical stages of process design

Inputs Design stage Outputs

Drilling results Resource definition Core samples

Core samples Laboratory testwork Lab test results

Mineralogy

Processing options

Results of lab work Scoping study Preliminary design

Processing method

±50% Capex and Opex

Factored costs

Rules of thumb for operating costs

NPV/DCF

Mass balance/Flowsheet options


Inputs Design stage Outputs

Scoping study

Further testwork

Prefeasibility study Finalised (high level) PFDs

Processing method finalised

±30% Capex and Opex

Quotes for major equipment

Minor equipment and other costs factored

NPV/DCF

Mass balance of processing route

Prefeasbility study Feasibility study PFDs and preliminary P&IDs

Preliminary plant layout

Plant utilities scoped

± 10-15% Capex and Opex

Quotes for majority of equipment

Estimates for civils, E/I etc

Finalised mass balance and energy consumption for major equipment

Feasibility study Definitive feasibility study P&IDs finalised

± 5-10% Capex and Opex

All equipment quoted

Detailed estimates for installation costs

Energy consumption for all equipment finalised and electrical design completed

Definitive feasibility study Bankable feasibility study Plant layout with pipe runs routed and costed

< ± 5% Capex and Opex

Final financial models completed and NPV finalised for bank use

Bankable feasibility study Front end engineering design Process and mechanical design of all major equipment

Mechanical drawings of major equipment

Detailed design Detailed design of all equipment

Full equipment drawings, P&IDs, pipe isometrics


The sequence of events given in Table 1 is not prescriptive and some projects may skip one or more steps in the process.

It can be said that 80% of the major decisions for the plant design are made within the first 20% of work. The scoping and prefeasibility stages are where the processing route is defined and the major equipment that will be used is decided. It is these decisions which, to a great extent, lock in the future environmental performance of the project. At these stages of the design process, changes can be more quickly and easily made than at other points of the design process. They will likely involve changes to the model/mass balance and PFDs only.

It is also likely that the data available to make decisions in these early stages of the design development is limited. In general, the data available at scoping study level is limited to laboratory results. Industry averages and typical costs are used for the operating costs such as energy, reagents and mining costs. Energy consumption for the process is usually estimated using industry averages and factors such as typical specific energy draw for similar plants. At the prefeasibility study level, more detailed information on energy consumption for major equipment is calculated but the total power draw of the process is still an unknown. By the end of prefeasibility, the processing route and plant PFDs are virtually locked in. Further stages of design are used to finalise minor equipment and obtain quotes for equipment and installation thereby reducing the uncertainty in the capital and operating costs. However these later stages of design typically do not alter the process flowsheet selected – they merely add detail to it.

The overall energy consumption of the plant is typically calculated at the feasibility/definitive feasibility study level. At this point, the energy consumption for all equipment is calculated or obtained from vendors and the electrical system is finalised. Power supply options are often treated as a separate study and completed around the feasibility study. This should include study of various power supply options and primary fuels, such as diesel vs. natural gas vs. coal and islanded power stations vs. grid connect etc.

INCORPORATION OF FUTURE ENERGY PRICING

As part of the development of the financial model, energy prices are included and under the current world economic climate are almost certain to escalate over time. Estimates of inflation are often used for this escalation and this is often set as a single number (e.g., 2.5% p.a. for the life of the project). Escalating prices using this method is not necessarily the most accurate method of energy price forecasting. In general, energy prices have been escalating far in excess of the inflation rate, and not necessarily in a linear fashion. To illustrate this, a comparison of Australian inflation (Rate Inflation, 2011), (ABS 6401.0, 2010) and world oil prices (US E.I.A., 2011) from 1990 to 2010 is shown in Figure 1. Oil prices have been used in this case as a reference to illustrate the fact that inflation is not always an accurate representation of price escalations. When undertaking a plant design, forward price estimates of the actual energy source should be used.


‐40.00%

‐20.00%

0.00%

20.00%

40.00%

60.00%

Variation of Australian CPI and oil prices

Australian inflation World oil prices

Figure 1 - Year on year variation of Australian CPI and world crude prices

The percentages shown in Fig. 1 are the annual rates of change of both the Australian Consumer Price Index (inflation represents this number) and the world oil pricing. Oil pricing for this data is defined as the average oil price in US dollars per barrel for that year.

It can be seen that inflation in Australia has increased by, on average, 2-5% per year. Oil prices on the other hand varied widely with greater than 30% decreases in price followed by 20-40% increases in price. The overall trend for oil pricing over this period has been upward as shown in Figure 2.

0

20

40

60

80

100

120

Oil price (USD

/bbl)

World crude oil prices ‐ 1990‐2010

Figure 2 - World crude prices – 1990-2010 – presented in dollars of the day

Since 1999, it is clear that the price of crude has escalated at a rate far in excess of inflation and that escalating energy costs in line with given inflation numbers could be seen as significantly understating future energy pricing. On average, since 1999, the oil price has escalated at 35% p.a. There are studies that have concluded that production of cheap oil has peaked. With increases in demand predicted, the actual oil price escalation is likely to continue to be high, year on year.


For operating cost accuracy within the design process, two things are important:

The first is to accurately characterise the future energy consumption of the operation; this includes the mine and processing plant. In addition, the power supply options should be considered early as fuel type and power plant efficiency will have a large effect on project NPV, particularly for off-grid sites. Final numbers will not be available for power consumption of all equipment but major equipment should be known and power consumption of the entire plant can be estimated from industry standards and benchmarking against similar operations.

The second important factor in determining operating cost accuracy is having a valid estimation of the forward pricing of energy sources. This includes plausible predictions of future pricing of all relevant energy types such as oil, coal, diesel, natural gas and grid connected power. Future energy pricing is defined by, among other things, predictions of supply and demand. Agencies such as the International Energy Agency and the US Department of Energy publish historical data on the supply and demand of many major energy sources. Both of these agencies also publish predictions for pricing and supply/demand of all major energy sources. It is important, however, to adjust these figures for economic conditions in Australia or even specific regional areas within Australia.

An example of this concerns natural gas pricing within Western Australia. The world pricing for natural gas has decreased since 2008 but pricing within WA has increased for domestic gas consumers. Suppliers in the north west of WA assert that the only plausible substitute for natural gas, in the short term, is diesel. Therefore, the price has increased such that natural gas is still cheaper than diesel but more expensive than global prices. Given that WA has no ability to receive global LNG, the local market is defined by diesel pricing.

The ideal situation is for sensitivity analyses to be completed on a number of different energy price scenarios to see the overall effect of projected future price variability on NPV. If the project is particularly sensitive to energy pricing (i.e., it is a large energy user or highly dependent on an energy source such as diesel) then methods of reducing energy consumption should be explored more extensively during design, and as early in the design process as possible.

Some examples of potential future scenarios for energy pricing (Energetics Pty Ltd, 2011) are shown in Figs. 3, 4, and 5.


Figure 3 - Australian diesel price forecast

Figure 4 - Gas price forecast - Western Australia

Figure 5 - Electricity price forecast - Western Australia

INCORPORATION OF FUTURE CARBON PRICING

Closely linked to energy pricing is the impact of future carbon pricing on the project. Generally, pricing carbon is not included, or included as a flat rate for the project life. Greenhouse gas (GHG) baselines and assessments are generally not completed as part of


the design so the extent of any carbon price liability is not included in project financial assessments. Typically, GHG baselines are assessed as part of the Environmental Impact Statement process, if at all. This is too late for the project team to be able to react in any significant way to what this baseline means. It is our recommendation that all resources projects consider a future price of carbon as these projects typically operate for a significant period of time and it is extremely likely that there will be international agreement on carbon pricing within the life expectancy of many resources projects. At the same time we do recognise that a future price on carbon is likely to have less of an impact on the project economics than the energy variability detailed above.

In order to incorporate carbon price scenarios into the project economics, the first step is to define the GHG baseline. To do this within the accuracy bounds of the design stages illustrated in Table 1, energy types and consumption of those energy types should be defined as well as sources of direct emissions for the plant. This includes:

Diesel consumption by mine fleet

Power plant size, fuel type and efficiency (for islanded power plants)

Imported power (for grid connect)

Natural gas/LPG consumption in process plant (e.g., for kilns, dryers etc)

CO2, methane and N2O produced by process (e.g., neutralisation using carbonates, reduction using coke etc)

CO2, methane and N2O produced from utility processes (e.g., calcination of lime),

Other greenhouse gases produced by process (e.g., natural gas vents, nitrous oxide production etc).

It is noted that there are industry specific emissions, for example coal seam methane emissions from underground and open cut mines, which also need to be accounted for.

These should be aggregated to give total equivalent CO2 emissions for the operation. Where possible, changes in energy consumption as part of business as usual should be incorporated into the financial model. This is particularly true for items such as energy consumption for mining operations, using the mine plan as the basis. As mines get deeper and trucks have to travel further/more ventilation is required, the change in energy consumption should be calculated and incorporated into business as usual.

When the GHG emissions have been calculated for the length of the financial model, the second step is to incorporate a realistic price of carbon. One way to do this is to consider a number of potential scenarios and determine a likely carbon price for each. These scenarios should be plausible given current world political situations. Examples of these include:

Low price scenario – Emissions trading scheme with low price caps set by government because of political sensitivities (AUD5-10/tonne CO2-e).


Mid price scenario – Medium level CO2 reduction targets set by government and carbon price set by emissions trading scheme with mid level caps on pricing (AUD15-20/tonne CO2-e).

High price scenario – Aggressive CO2 reduction targets set promoting large scale fuel switching projects (e.g., coal power plants to natural gas/renewable energy) (AUD30-40/tonne CO2-e).

The pricing then determines the likely carbon liability for the financial model and is incorporated into NPV calculations.

ENERGY EFFICIENCY ASSESSMENT AND DESIGN REVISION

With energy and carbon costs fully accounted, and forming part of the NPV calculation, the question of how to reduce these costs arises. For energy costs, a desktop energy efficiency assessment can be conducted. In an ideal situation, this assessment should be completed as early as possible in the design process so that changes identified can be incorporated at minimal cost and disruption to the project.

Traditionally, energy efficiency assessments are performed after the plant has been running for a period of time. Energy saving opportunities are identified and then retrofitted to the plant. This is an inefficient use of resources (people and money) as the plant is built - for a particular capital cost, operated for some time at a lower efficiency than it potentially could, then further capital is spent to retrofit an energy efficiency project. With a slight increase to the upfront capital of the operation, energy and operating cost savings could be realised from commissioning onwards.

It is also beneficial to undertake a number of energy efficiency studies along the project life cycle. At the early stages of design (scoping and prefeasibility level), decisions can be made on major pieces of equipment or processing routes. Examples of this include:

Use of in pit crushing and conveying vs. haul truck

Use of high pressure grinding rolls instead of or in circuit with traditional milling equipment

Addition of extra crushing stages to maximise work done in crushing rather than grinding

Including by-pass capacity for the comminution circuit

Selection of hydromet process routes vs pyromet vs integrated hydromet and pyromet process routes

Use of atmospheric leach vs. pressure leach,

Use of heat recovery systems on power generation equipment.

These decisions can be made while the process is still being finalised. Once the process is bedded down and more of the minor equipment is chosen and specified, further assessments can be conducted and decisions can be made on items such as:


Use of variable speed drives on pumps, compressors and fans with variable duties

Optimisation of pumping and pipe runs to ensure pumps are operating close to maximum efficiency

Modification of pipe runs to ensure pipe losses are minimised,

Integration of heat sources and sinks where appropriate.

Finally, at the time the control system is specified and P&IDs drawn up, energy efficient design decisions for plant control can be made including:

Conveyor slowing and/or shutdown signals based on weightometer readings

Soft start technologies

Underground fan telemetry systems,

Expert systems for energy intensive process steps, for example grinding control, any pyromet processes in the plant and electrowinning/refining.

There is a clear advantage in conducting energy efficiency assessments during the design process as information becomes available. Design revisions undertaken at this stage are much more cost effective than those completed once the plant is already built and commissioned. Conducting successful energy efficiency assessments is dependent on having good quality information on site energy use as the design proceeds. This may mean bringing forward some of the detail in the steps around developing the energy balance and electrical sizing of equipment. Detailed mine planning may also need to be brought forward in order to optimise mine fleet and production schedules. The power supply options should also be examined during the first stages of the design process as they will have a large bearing on the overall energy consumption of site. Advancing to this level of detail sooner will involve a higher engineering cost during the early stages of design and potentially a higher plant capital cost overall but the NPV of the project will be improved.

CARBON REDUCTION PROJECTS

Similar to the energy efficiency assessments, carbon efficiency can also be looked at during process design. It is recommended that energy efficiency and carbon efficiency are looked at during the same assessment stages. In fact, all energy efficiency projects will have a carbon benefit that should be quantified and contribute to informing the GHG baseline. However, there are some projects that may result in carbon savings without involving a reduction in energy. Examples of these sorts of projects include:

Alternative processes to neutralise acids or for pH control

Fuel switching by using natural gas instead of diesel or coal

Inclusion of renewable energy projects such as wind-diesel hybrid power generation,

Site planning to reduce disturbance of vegetation and upfront planning of site rehabilitation.


As with the energy efficiency assessments, it is ideal if the carbon reduction assessments are completed earlier in design process. Obviously the carbon baseline should be quantified prior to undertaking the assessment.

GREEN HAZOP

The concept of a “Green HAZOP” is relatively new. It involves a structured process, much like a regular plant HAZOP, that focuses solely on the environmental impact of the operation. Energy consumption and energy efficiency form part of this review. Basically, a multi-disciplinary team conducts a systematic study of the process using guidewords to determine how far the existing process is from being a “green” process. Recommendations are then made to increase the overall sustainability of the operation. (Garcia-Serna, Martinez and Cocero 2007)

The guidewords used in a Green HAZOP are the same as those used in a regular HAZOP. These are shown below in Table 2.

Table 2 - Green HAZOP guidewords

Guideword Application/meaning

No Negation of the design intent

Less Quantitative decrease

More Quantitative increase

Reverse Logical opposite of the intent

Other Complete or partial substitution

The guidewords given in Table 2 are then applied to a number of parameters listed in Table 3 for various areas of the plant. Note, that these parameters apply to both the sustainability and environmental impacts on any decisions made for that plant. It is possible to tailor the process to only the parameters relating to energy or carbon as required.

Table 3 - Parameters used in Green HAZOP

Area Parameter Application Notes

Raw materials and products

Renewable Material Material source renewable or depleting

Energy Energy source renewable or depleting

Diversity Material Minimum number of substances or parts

Energy Minimum number of energy sources

Minimisation Material Minimum use of material flows

Energy Minimum use of energy flows


Area Parameter Application Notes

Space Minimum use of space

Time As quick as possible

Toxicity Material Use substances inherently non hazardous

Process

Over design Material Use the exact amount of materials

Energy Use the exact amount of energy

On demand

Material Material flows should be output pulled rather than input pushed

Energy Energy flows should be output pulled rather than input pushed

Volume Equipment over design

Gaseous or liquid emission

Reduce gaseous emissions (CO2, CO, NOx, SOx)

Efficiency Efficiency

Material Efficient processes in terms of mass balance

Energy Efficient processes in terms of energy balance

Space Efficient processes in terms of space

Time Efficient processes in terms of time – for batch operations and plant startup

Integration Integration and interconnectivity

Material Integrate material flows, avoiding waste

Energy Integrate energy flows, avoiding heat loss and energy quality loss

Controllability

Controllability Material Controllability of material flows and equipment

Energy Controllability of energy flows and equipment

Operability Material Operability of material flows and equipment

Energy Operability of energy flows and equipment

During the HAZOP meetings, the deviation guidewords are applied to the parameters and the causes and consequences of each are brainstormed. An indication of how sustainable the proposed process is can be obtained from this discussion.

The next step in the process is to identify recommendations to improve the sustainability of the process. In the case of energy efficiency or carbon reduction, these are the projects that will result in an energy efficiency increase, a reduction in total energy draw, or a reduction in carbon emissions. The recommendations are then reviewed with respect to identifying those that will be implemented. A cost benefit analysis of each recommendation should be completed and those projects that make sense from an environmental, financial and safety point of view should be implemented.


CASE STUDY

A case study for energy efficient assessments is presented. It concerns an assessment completed on a coal project in Queensland. This plant had completed the prefeasibility study when the energy efficiency workshop and design reviews were held. For the energy efficiency review, a team of 15 people was assembled. This team was made up of representatives from operations, production, maintenance, environmental and management for the company that owned the plant, design engineers from the consulting engineering company and external energy consultants.

The operation was split into a number of broad areas and each was examined in turn. These areas were:

Management

Infrastructure

Layout

Ancillaries

Supply,

Technology.

For each area, energy efficiency opportunities were listed. This included opportunities that already formed part of the design as well as those which should be considered. All opportunities presented were listed. When the opportunities had been identified, they were mapped to a ranking matrix that grouped them into four areas based on impact/benefit and cost/effort as follows:

Just do it – High benefit and low cost

Symbolic – Low benefit and low cost

Strategic – High benefit and high cost,

Leave behind – Low benefit and high cost.

A graphical representation is shown below in Figure 6.


Figure 6 - Energy efficiency project ranking

In this study, a total of 25 energy efficiency opportunities were identified, of which only three were in the leave behind category. Together with the project owners, a list of actions, with project owners was generated so that the projects could be progressed and incorporated into the design. The first step was generally to complete further work on each project to quantify the energy savings and costs and make a determination on whether the project would go ahead or not.

A summary of opportunities identified is given below.

Management and systems – e.g., Energy efficiency targets for equipment and processes - 4 projects in total

Infrastructure – e.g., Eco-villages with optimum orientation, trees for shading, recycled water etc - 8 projects in total

Layout – e.g., overland conveyors instead of haul trucks - 4 projects in total

Ancillaries – e.g., fire water system with staged pressure increase in line with demand - 2 projects in total

Supply – e.g., sub metering for power consumption - 1 project in total

Technology – e.g., alternate fuel types for mobile equipment - 6 projects in total

High cost

High impact

Low impact

Low cost


CONCLUSIONS

In this paper the case has been presented the case for including considerations of energy efficiency and carbon management early in the design process. Why this is necessary has been indicated and it is explained how to incorporate these considerations into the design scheduling. It has been demonstrated that including energy efficiency in design considerations is possible using a case study. Two specific considerations in this process, energy and carbon price forecasting, and the timing of the assessment are noted.

To generate a more accurate view of project NPV during design, use of valid assumptions for energy pricing is a key input, as is the methodology used to generate energy price forecasts. This depends on having accurate information for the energy consumption over time and having accurate indications of price escalation over the project life. A valid pricing scenario for carbon is also a key input to the project financial model. This involves modelling of carbon emissions over the project life and the inclusion of a carbon price forecast.

With energy and carbon pricing quantified, it is recommended that energy efficiency assessments and carbon reduction assessments be conducted on the operation during one or more design stages. This can be done using a facilitated workshop or structured green HAZOP methodology. Regardless of the method used, making changes to the plant or mine design or to equipment specification is easier and cheaper, the earlier it is done in the design process – and much easier and cheaper than retrofitting the plant once operational. Energy and carbon reduction projects should be identified early and a cost/benefit analysis completed for each so that they may be included in the design.

REFERENCES

Australian Bureau of Statistics. “6401.0 Consumer price index.” ABS. December 2010. http://www.abs.gov.au/ausstats/[email protected]/mf/6401.0/ (accessed March 3, 2011).

Energetics Pty Ltd. “Energetics Energy Markets Outlook 2020.” February 2011.

Garcia-Serna, J, J L Martinez, and M J Cocero. “Green HAZOP analysis: incorporating green engineering into design, assessment and implementation of chemical processes.” Green Chemistry, 2007: 111-124.

Rate Inflation. Australia Historical Inflation Rates. March 2011. http://www.rateinflation.com/inflation-rate/australia-historical-inflation-rate.php (accessed March 2011).

U.S. Energy Information Administration. “World crude oil prices.” Energy information administration. 3 March 2011. http://www.eia.gov/dnav/pet/pet_pri_wco_k_w.htm (accessed March 3, 2011).


Filtration Test Work – Extracting the Whole Story for Studies and Design

G Bickert1 and B Länger2

1. MAusIMM, Principal Consultant, GBL Process Pty Ltd, 1/7 Jubilee Avenue, Warriewood NSW 2102. Email: [email protected]

2. MAusIMM, Principal Consultant, GBL Process Pty Ltd, 1/7 Jubilee Avenue, Warriewood NSW 2102. Email: [email protected]

ABSTRACT Australia is currently developing and engineering a range of specialised processing options for rare earth

(RE) (Lynas, 2011, Alkane, 2005), zirconia (Alkane, 2005), lithium (Galaxy Resources, 2008) and other

special minerals and metals required in high-tech applications.

The processing of lithium, zirconium and RE minerals is based on complex leaching processes which require

extensive liquor clarification, solid liquid separation and solids washing. The filtration properties for those

leach liquors, residues and precipitates are not well understood and differ drastically according to the

process route applied. Thus, filtration test work is essential for equipment selection, sizing and design.

Similar types of slurries are filtered in laterite nickel, TiO2 and other chemical oriented processes, while more

experience is available here as those processes are applied widely for a number of decades. The increasing

need for dry stacking of tailings requires advanced tailings filtration with properties differing drastically

depending on tailings properties. More traditional mineral processes such as flotation commonly apply

filtration. All of these and other cases require proper filtration test work particularly evaluation of vacuum

versus pressure filtration and the need for cake pressing and/or air blowing for filter presses. Bench scale

test work and/or pilot plant operation is common to determine filtration rates, achievable cake moisture and

wash efficiencies and sometimes filtrate clarity. While those values are important, there is much more valid

information to be obtained by filtration test work.

This paper discusses the intricacy of filter selection and sizing based on driving forces (vacuum and pressure

air, compression by mechanical means and slurry) and parameters (cake thickness, forces and times) and

different processes utilised in common filtration equipment. It will discuss the in’s and out’s of filtration test

work design, planning, and in particular data gathering, processing, reporting, evaluation and filter sizing

from recent examples of bench and pilot scale filtration test work for above applications together with

methods for predicting filtration performance for example where test work is lacking. One focus will be the

realistic extrapolation of test work results towards different parameters utilised in full scale equipment and

plants. The conclusion of the paper will contain checklist for filtration test work.

INTRODUCTION Filtration (Chen, 1997) is an important unit operation in almost all metallurgical plants and flowsheets.

Flotation concentrates, leach residues and precipitated or crystallized valuable intermediates or products are


dewatered by vacuum or pressure filtration (Townsend, 2003). Clarifier overflows, pregnant liquor or other

filtrates are polished by candle or pressure leaf type filters.

The processing of lithium (Galaxy Resources, 2008), zirconium (Alkane, 2005) and rare earth (RE) minerals

(Lynas, 2011) for example is based on partly new and complex leaching processes which require extensive

liquor clarification, solid liquid separation and solids washing. The filtration properties for those leach liquors,

residues and precipitates are not well understood and differ drastically according to process route applied.

Thus, filtration test work is essential for equipment selection, sizing and design. Similar slurries are filtered in

laterite nickel, TiO2 and other chemical oriented processes, while more experience is available as those

processes are applied widely. The increasing need for dry stacking of tailings requires advanced tailings

filtration with properties differing drastically depending on tailings properties. More traditional mineral

processing for example of base metals commonly applies filtration of flotation concentrate. All of these and

other cases require proper filtration test work particularly evaluation of the need for and quantity of cake

pressing and/or air blowing.

Sizing and designing filtration equipment is more often done by limited bench scale test work rather than pilot

scale test work. The filtration test work might already being done during concept studies, should be part of a

feasibility study (Andersen, 2000) and is sometimes only undertaken during detail design particular in boom

times for highly profitable and thus, fast-tracked mine developments. Test work results for filtration

equipment selection and sizing often includes only a few tests with limited results which do not reflect the full

scale filter conditions or else detailed pilot plant studies with extensive quantity of test results which partly

lack data gathering or focus only on recording filtration sizing parameters. Filtration test work is mostly done

by OEM’s or metallurgical laboratories. Limited scientific work (peer reviewed journal papers) is published on

filtration test work recoverable by searching the keywords “filtration” and “test work” bringing up a little

information from all major scientific journals. Non peer-reviewed information, in particular company

information on web-pages is plentiful, partly very useful and partly advertising only. Kram, 2002 discusses

filtration test work do’s and don’ts and provides detailed requirements regarding production plant and slurry

(solids and liquids) material properties important for filtration test work and discusses filter test scale and

appropriate filter types.

The literature on filtration covering theory, filter selection, sizing and experimental work is extensive (Anlauf

and Sorrentino, 2004, Anlauf and Stahl, 1996, Anthony et al., 2010, Svarovsky, 2000, Sutherland, 2005,

Wakeman and Tarleton, 1999, 2005a and b, Chen, 1997) however practical test work often does not apply

all this available knowledge. Haekkinen et al. (2008) introduce statistical design of experiments to reduce the

number of filtration tests required to optimize pumping, pressing (squeezing) and drying time and pressure.

However, not many practical guidelines for reporting and utilizing filtration test work data for filter selection,

sizing and designs are published.

Selecting, sizing and designing of filtration equipment for common duties, such as base metal concentrates

is straightforward due to extensive experiences, plant examples and historical full scale data, which resides

with OEM’s, engineering firms and consultants. More complicated filtration duties in hydrometallurgical


flowsheets with leaching, precipitation, crystallization or difficult to filter tailings filtration duties require case

by case evaluation of well planned, executed, evaluated and reported filtration test work.

Specifying filtration equipment for plant design often includes only limited data, in particular filtration rate (dry

solids production rate per filter area in kg/m2/h) and required cake moisture. However, this limited data is

only suitable for equipment sizing if the parameter under which the filtration rate were determined are

identical to full scale parameters, including feed density, cake thickness, driving force applied and others.

This paper discusses the intricacy of filter selection and sizing based on driving forces (vacuum and pressure

air, compression by mechanical means and slurry pressure) and parameters (feed density, cake thickness,

dewatering forces and times) and different processes utilised in common filtration equipment. It provides

guidelines and examples for filtration test work design, planning, and in particular data gathering and

processing on recent examples of bench and pilot scale filtration test work for various minerals. This should

lead to a more systematic approach to gather and evaluate specific test work data holistically by covering all

required variable process data for selecting, sizing and design.

Filtration theory is useful in extrapolating test work results towards different parameters utilised in full scale

equipment and plants when test work data is lacking.

CAKE FILTRATION THEORY Following is a brief summary of the filtration theory applicable for filter testing, sizing and optimisation. It is

based on Darcy Law (Anlauf and Stahl, 1996), the basic filtration theory describing the specific filtrate flow as

a function of resistance, cake height and liquid viscosity is given below but it is only valid for cake formation

and cake washing. The second part of this only considers the total cake resistance through specific cake

resistance multiplied by cake thickness plus the filter media resistance:

)R H(r μ

Δp Hμ r

ΔpA1

dtdVq

mcllc

smm

2

3

(Equation 1)

With q: specific Filtrate volume / m3/m2/s V: Filtrate volume accumulated / m3

A: Filter area / m2 rc: specific cake resistance / 1/m2

H: Cake height / m l: liquid viscosity / Pas

pΔ : applied filtration pressure Rm: Filter media resistance / 1/m

The cake resistance rc [1/m2] is used, while α [m/kg] is also commonly used. Both are directly related

according to:

k

HHr

A

cVw c

m1 (Equation 2)

With w: mass of solids per cake height / kg/m α: specific cake resistance / m/kg


Equation 1 can be directly used for cake washing i.e. a steady state one phase flow of fluid through a

particulate bed (the filter cake). Only the liquid viscosity might change while the mother liquor is progressively

replaced by wash liquor, if both liquids have different viscosities.

As the cake height is not known during filtration test work but the filtrate volume can be measured, the cake

height is replaced by the filtrate volume through utilising the filter feed solids concentration which can be

calculated from volume or weight concentration if solids and liquid density and cake porosity are known:

A

V H κ

and lws

lw

v

v

ρ C ερ ε1

ρ C

C - ε - 1 C

κ

- (Equation 3)

With : cake porosity (pore volume/total volume) / - Cw, CV: solids concentration / wt/wt, vol/vol

: density of solids (s) or liquid (l) / kg/m3

For constant pressure filtration, which is the case for vacuum filtration, the equation above results in:

Δp A Rμ

VΔp A 2

κ r μVt ml

2cl

3ms

(Equation 4)

Thus, if filtration test work has recorded the filtrate accumulated over time, plotting the time versus filtrate

volume (or mass) over filtrate volume should give a linear curve. The specific cake resistance (or specific

cake resistance multiplied by viscosity if latter is unknown) can be determined from the slope of the line and

the filter media resistance from the intercept of the y-axis (EN ISO 6145-5:2010). Only one filter test is

required for this, which also could be used to determine the desaturation kinetics.

The filtration time for neglectable media resistance (Rm=0) is a function of filtration parameters such as cake

height, dewatering driving pressure, feed density and specific cake resistance according to:

Δp κ 2h r μ 2

cclt s (Equation 5)

This formula also shows that the cake height increases according to a square root proportionality with

filtration time, so that the specific solids throughput of continuous vacuum filters is proportional to the square

root of the filter speed.

For rotary filters, where the cake formation ratio to cake drying ratio is constant, the rotation speed n is

applicable and the solids throughput rate is dependent on filter speed n according to:

)C - - (1 r μn Δp 2

r μn Δp κ 2

Vclcl sm

s

kg (Equation 6)


With n: filter rotation speed in 1/s sm : dry solids throughput / kg/s

Thus, plotting the specific solids throughput rate as function of square root of filter speed should provide a

linear curve, which is confirmed by Ehrfeld et al. 2008.

The required filter area for cake formation can be calculated from Equation 5 with the cake density according

to:

)1( Δp κ 2h r μ

As

ccls

m 2m (Equation 7)

While extensive work has been undertaken to predict cake desaturation, this is not achieved unless

extensive experimentation and analyses are performed to determine to determine their broad size

distribution and complicated microscopic structure (Anlauf and Sorrentino, 2004). However, the change in

cake saturation with time is proportional to the desaturation driving force and the cake height to the power of

minus two. The driving force for desaturation is the pressure difference between applied pressure and the

capillary entrance pressure at the particular cake saturation.

H)(

pΔ - pΔ

) H H (

pΔ - pΔ

dt

d2

cap(S)

2cE

cap(S)

S

s

1 (Equation 8)

With cap(S)pΔ : capillary entrance pressure at the particular cake saturation in Pa

HcE : equivalent cake height (for filter media)

With the saturation or moisture reducing over time, the moisture drop gradually decreases to zero if no

thermal drying occurs.

As the cake desaturation cannot be sufficiently predicted theoretically for practical filter duties, the achieved

cake moisture for various drying times (for applicable pressure difference and cake thickness) has to be

determined experimentally by one test recording the filtrate volume over time during drying. This could be the

same test used for determining cake and media resistance. For disc filters, the drying time is almost fixed

through disc filter speed and filter design and thus, the correct ratio between cake formation time and

desaturation time (about 0.8 to 2) needs to be applied.

VACUUM BELT, DISC, DRUM AND PAN FILTER TESTING Buchner funnel test work is common to simulate continuous vacuum filters such as horizontal vacuum belt,

rotary vacuum disc, drum and pan filters. The Buchner funnel set-up could be of various designs and

complexity. Bourgeois et al. (1995) proposed a standard benchmark filter test for vacuum filtration, simulating

bottom fed and top fed filtration for coal, but the method was not adopted. A detailed European Norm exists

for the determination of filter cake resistance (EN ISO 6145-5:2010) which is partly considered but still, not

well known. This Norm uses over-pressure filtration (hyperbaric) for determining filtration parameters such as

cake resistance as shown in Fig. 1. (EN ISO 6145-5:2010). Hyperbaric filtration is even used for sizing


vacuum filters (Ehrfeld et al., 2008) which assumes that filtration is similar for similar pressure difference, be

it vacuum or overpressure. All these batch tests simulate the continuous vacuum filter by applying the filter

stages across time (batch filtration), while the full scale continuous filters have similar stages across the

position on the filter (continuous filtration). Thus it is important to considering the different filter stages.

Figure 1: Pressure filter used to determine specific cake and filter media resistance (EN ISO 6145-5:2010)

Continuous Vacuum Filter Stages All continuous vacuum (and hyperbaric) filters apply the following stages while cake washing might or might

not be necessary:

1. Cake formation – one phase (liquid = mother liquor) flow through the filter media to transform a slurry

into a filter cake;

2. Cake washing – one phase (liquid = wash liquor) flow through the filter cake to replace the mother

liquor within the cake with wash liquor; and

3. Cake desaturation – two phase (liquid and gas) flow to replace the liquor within the cake with air to

reduce its cake moisture. This is also called deliquoring or drying.

It is important to record the transition from one stage to the next during test work and use different theoretical

approaches for those stages during test work evaluation and in particular for extrapolation.

Common Vacuum Filter Test work Most commonly, vacuum filtration is tested directly by applying reasonable parameters for:

Slurry feed density;

Cake formation time in ratio to cake desaturation time;


Cake thickness; and

Vacuum pressure applied.

Varying the ratio of cake formation to cake desaturation time is limited for rotary vacuum filters due to their

design of pick-up zone and drying zone and it is mostly fixed by design and not adjustable during operation.

In particular for disc filters, this ratio is only changeable by reducing the cake formation (pick-up) zone in

length and thus in time. Here, no difference is made between disc filters using textile cloth or mesh or

ceramic material as filter media as this applies to both. Drum filters have a slightly bigger ratio variation. In

contrast, the ratio between cake formation and desaturation develops freely on horizontal vacuum filters

such as belt and pan filters and thus, is not a controllable design parameter.

The filtration test work result, in particular specific solids throughput rate / kg/m2/h (commonly referred to as

filtration rate) and cake moisture is valid only for those selected parameters. Extrapolating the results for

different feed densities, cake thicknesses or cake formation to cake desaturation time ratio is only valid

through experience (and applying filtration theory indirectly) or by directly applying the filtration theory. More

commonly, the effect of those parameters is directly tested, thus requiring more test work.

The well known filtration theory allows determining filtration parameters such as the specific cake resistance

(or cake permeability) which then could be applied to extrapolate, predict and optimize filtration.

Nevertheless, it is only applied in very limited practical filtration test work cases directly.

Example for filtration theory and filtration testing for vacuum and hyperbaric filtration The above theoretical considerations will be used for extrapolating filtration and washing from an example of

leach residue where only limited test work could be performed. The following graphs in Fig. 2 show an

example of applying this theory for one vacuum filtration test, where the filtrate volume gathered over time is

first plotted over time and then according to Equation 4 as time over filtrate volume versus filtrate volume.

The filtrate volume can be replaced if preferred by filtrate mass for known liquid density. In this example, the

cake formation time was not recorded and only one filtration test without cake washing was undertaken while

applied filtration parameters (vacuum pressure and cake thickness) were different from the anticipated full

scale parameters.


FIG 2: Test work data (Filtrate over time) plotted on the left

and as per filtration theory Equation 4 on the right as t over V versus V

It is essential to record the cake formation time (transformation from slurry to cake) during bench scale batch

filtration test work and then apply filtration theory individually for cake formation and cake washing and

consider cake drying separately. This is particular important for testing of horizontal vacuum belt filter

applications, as the cake formation to cake desaturation ratio is not fixed but depends on filter feed

conditions and could vary drastically. This is not as critical for filtration test work done properly for rotary filter

applications (i.e. applying the correct and constant ratio between cake formation and desaturation).

The cake formation time in above example was not recorded but it can roughly be estimated from the

following graph (Figure 3) which shows again the t/V over V curve. The t/V curve is linear to the last red

square data point (6 data points) so that the cake is formed between this value and the next, which is in this

example between 70 and 120 s. Visual observation during test work would have been significantly more

accurate.

Here, very limited filtrate data points are available as the filtrate was recorded manually. The first two data

points have to be neglected for data analysis, due to experimental start-up errors. These start-up errors are

typically experienced and most likely due no instantaneous measurement of the filtrate as it exits the filter

media but which has to be, for example, collected within a volumetric flask or on a scale. The following graph

demonstrates that by applying Equation 4 across different sets of data points (or excluding certain data

points on the fringes of cake formation) results in different linear approximations and thus, different values for

specific cake resistance and filter media resistance. Thus, having electronic data recording in short intervals

and recording cake formation time is helpful and improves the accuracy of the test work results and their

application.


Fig 3: Test work data plotted as t over V versus V with linear approximations

(according to Equation 4 to determine specific cake and media resistance)

depending on selection of data points shown with linear formula shown.

However, with those parameters determined from plot above, the required filter area for cake formation can

be universally determined for various vacuum pressures, cake thicknesses and feed solids concentrations.

The following Table 1 gives some examples the of required filter area depending on selected cake height

and vacuum pressure by pure theoretical approach according to Equation 4:

Table 1 – Calculated required cake formation filter area

for different combinations of vacuum pressure and cake height

Cake height m 0.01 0.02 0.015 0.02 0.02

Vacuum pressure kPa 4000 2400 4000 3200 4000

Relative Filter area required for cake formation

% 30% 100% 45% 75% 60%

The feed solids affects the required filter area according to Equation 5 with replaced through Equation 3.

The area increases with more diluted feed while the increase is dependent on solid and liquid density and

cake porosity. The required filter area can be estimated for feed solids which have not been tested if those

material and cake properties are known.

Applying filtration theory therefore eliminates the need to test various feed solids concentration, slurry

volumes or cake formation time, vacuum pressures and cake heights in particular for cases where very

limited slurry for testing is available and can assist in optimising filtration parameters. If sufficient slurry is

available it is still recommended to confirm that the filtration theory which is based on certain assumptions

(in particular homogeneous filter cake and no cake cracking) is valid across the parameters tested by

undertaken filtration tests with the end points of the parameters altered. Filtration theory assumes a

homogeneous filter cake. This is not valid if segregation occurs. If feed solids concentration is reduced


significantly and the slurry changes in regards to sedimentation from hindered to swarm settling (Bickert

2010) then significant segregation suddenly could occur on top fed filters due to gravity settling in particular

for long cake formation times.

Filtration theory allows also to extrapolate expected filtration performance beyond parameters tested or to

predict full scale filter operation at different operating parameters or to predict cake washing kinetics as this

could not be tested experimentally.

Cake washing Once the specific cake resistance and the filter media resistance is determined from one filtration test, the

required filter area for cake washing depending on the wash liquor amount can be determined theoretically.

However, if possible it is still recommended to confirm cake washing kinetics by experiment, while it is

essential to determine wash efficiency. Theory allows the determination of the required filter area for varied

wash amounts and various wash stages such as multi-stage counter current cake washing on horizontal

vacuum belt filters.

Equation 2 can directly be used to calculate the filter area required for cake washing. It depends on wash

liquor flow, viscosity, pressure difference, cake height and filter media resistance. The dependence for those

parameters for neglectable filter media resistance is similar as in the case of cake formation. Thus, the

required filter area for cake washing is two-fold for half the pressure or double the cake thickness.

Cake Desaturation The cake desaturation cannot be sufficiently predicted theoretically for practical filter duties, so that the

achieved cake moisture for various drying times (for applicable pressure difference and cake thickness) has

to be determined experimentally by one test recording the filtrate over time during drying. This could be the

same test used for determining cake and media resistance. For disc filters, the drying time is almost fixed

through disc filter speed and filter design and thus, the correct ratio between cake formation time and

desaturation time (about 0.8 to 2) needs to be applied.

Following graph (Fig. 4) shows the extrapolated moisture drop over time during desaturation time,

commencing after cake formation time.

The appropriate drying time can be determined from Fig. 4 with the knowledge of the required cake moisture.

For example is 60s drying time sufficient to achieve 30% moisture while 220s are required to achieve 28%

cake moisture. For sizing, additional limitations have to be considered which are:

The drying time on rotary filters limitation due to limited ratio between cake formation and drying;

To provide sufficient filter length on top fed (belt and pan) filters to avoid slurry flowing into cake

discharge area and to allow for uneven cake distribution across the filter width.


Fig 4: Test work data for moisture over desaturation time

For cake desaturation, belt and pan filters utilise typically 30 to 70% of the filter area required for cake

formation. This should be above 10 – 15% if multiple stage cake washing is done on the filter.

The total effective filter area required is the sum of filter areas for cake formation, washing and desaturation,

all for a particular cake thickness, feed solids and pressure.

PRESSURE FILTER (CHAMBER, MEMBRANE AND TOWER FILTER PRESS AND CANDLE, LEAF AND KELLY FILTER TESTING

Overview of filtration principles Continuous pressure filters such as hyperbaric disc filters are treated similarly to continuous vacuum disc

filters and the above details apply with the appropriate pressure. Batch operated filters have different duties,

designs, cycles and requirements. The pressure filtration occurs due to different driving forces:

Slurry feeding under pressure (pumping) results in initial filtration and compaction of a cake which is

completely saturated;

Squeezing by mechanical means (‘membrane’ or better diaphragm bladder squeezing through fluid

behind it (mostly either air or water) or flexible chambers through rubber seals and hydraulic plate

compaction). The cake is still fully saturated but at lower moisture if compressible;

Compressed air pressure replaces the liquid in the cake with air and thus, reduces the saturation

and cake moisture, if the air pressure is sufficient to overcome the capillary forces holding the water

within the cake; and

Shear combined with mechanical pressure such as in belt press filters or screw presses results in

cake moisture reduction for compressible cakes which benefit from shearing.


Simplistically, the main pressure filters used in metallurgical plants are:

1. Filter presses (chamber, membrane, tower, tube presses) with following cycles being all or partly

applied,

Feeding and filtration (also referred to as pumping);

Squeezing by membrane pressing or other means (also referred to as pressing);

Cake washing (optional);

Cake drying by compressed air blowing; and

Technical time for opening, cake discharge, cloth washing, closing and other equipment

requirements.

2. Candle, leaf and Kelly type filters applying some of the above cycles; and

3. Belt press filters and screw presses.

Filter presses are predominantly solids focused, producing a dry cake, while the other pressure filters are

predominantly liquid focused, producing a clear filtrate. However, versatile hybrids such as the Fundabac®

candle filter can apply all the above filter press cycles excluding squeezing but with optional steam drying

within the pressure vessel with automatic cake discharge. Thus, filter presses are often specified by using a

“filtration rate” better expressed as specific solids throughput rate (kg/m2/h), while candle, leaf and Kelly type

filters use “filtration rate” which is better expressed as specific filtrate throughput rate (m3/m2/h).

Common Filter Press Test work Small scale filter press test work (Mayer et al., 2009) is commonly undertaken with a range of filter cells of

various sizes, with most OEM’s having their own design system. The important thing here is (Palmer, 2008)

that the laboratory filter reflects the design of the full scale filter as closely as possible with same and

variable chamber depth and drainage area. Palmer (2008) recommends at least 0.1m diameter as the error

increases below this. At bench scale, test work simulating tower filter presses should preferably be done

with a horizontal chamber on bench scale while chamber/membrane filter presses require a vertical chamber

orientation with the latter including a pressing device such as membrane or piston.

The laboratory filter press should be able to simulate all process cycles, including filling / filtration by either

pump or pressurized slurry vessel, squeezing by membrane or other means, cake washing if required and

cake compressed air blowing.

The available filter area should reflect full size, either one-sided or two sided, with the correct drainage area

used for squeezing and air blowing. Alternatively, one sided laboratory test work for two-sided filter presses

simulates only half the cake thickness, while cake squeezing and in particular cake blowing is not simulated

properly.


Most commonly, pressure filtration is tested directly, by applying reasonable times and pressures for feeding

and filtration (pumping), squeezing and compressed air drying. However, the result is sometimes only

reported as ‘filtration rate’ meaning the specific solids throughput rate in kg/m2/h. This throughput rate might

include the technical time or not. Thus, the throughput rate cannot be used directly for filter press sizing,

unless all individual full scale filter press cycle times are the same as in the laboratory test work and the

correct technical time, which depends on the filter press design, is already correctly incorporated in the

throughput rate.

Pilot scale filter presses are more truly scale-downed full scale filters with the pilot scale reflecting the OEM’s

design particulars such as horizontal plate orientation for tower filter presses. Thus, their scale-up is more

straightforward and less risky.

Filter press sizing is based on two requirements:

1. Sufficient filter area for filtering the slurry within the used feeding cycle time, taking into account the

feeding pump characteristics during chamber filling and the filtration characteristics (filtration rate)

during filtration; and

2. Sufficient cake volume which is discharged every cycle to provide the required solids throughput,

with taking the dry cake density into account.

The second requirement is often the limiting factor for filter sizing i.e. considering the filter press as a batch

volume that is discharged certain times an hour and dropping a certain cake volume and weight, resulting in

the total solids throughput. It takes the total cycle time into account, while the filtration cycle is determined by

the first requirement of filtration rate.

Thus, proper reporting of filter press test work requires all data, including filtrate gained during squeezing

and air blowing cycle, the times and pressures applied for feeding, squeezing and drying as well as the final

cake weight, its wet or dry cake density and its compression due to squeezing.

Examples for filtration theory and filtration testing for filter presses An example of test work for filter presses to dry stack tailings is presented. The test work simulated the filter

press process of feeding, squeezing and air blowing for various chamber thicknesses as well as with and

without flocculant. Measuring the filtrate flow over time for all cycles is beneficial to extract further information

out of the test work, in particular for the filtration, squeezing and drying kinetics which than could be used to

optimise cycle times theoretically. The assumption of constant filtration pressure is not valid during the

complete feeding / filtration period, as the feeding pressure is increased initially to a maximum pressure.

From then on the assumption is valid and the filtration theory for constant pressure discussed previously is

applicable. Fig. 5 shows the plot of t over m versus m (in this case we use the filtrate weight in g and the

filtrate time in min and not the filtrate volume) for three tests undertaken, varying only chamber thickness at

25, 41 and 57mm. This requires the recording of filtrate weight (or volume) for the duration of filtration.


Fig 5: Test work data for filtration time over filtrate weight versus filtrate weight for filter press laboratory

scale test work at various chamber thicknesses with and without flocculant.

Shown is the theoretically approximated filtration time (CFT) from 41 to 25 and 57mm

The tests show a linear increase over time for the initial feeding, till the chamber is filled. Test T3 is parallel

(indicating a similar specific cake resistance to the other tests) but is parallely shifted upwards. This indicates

that the filter media resistance is larger for this test, so the test used either a less permeable filter cloth or it

was blocked, the latter being the case here. The thinner the chamber thickness, the earlier (at lower filtrate

weight collected) does the slope of the curve increase drastically, which indicates that the chamber is full.

This is visible for Test T2 in Fig. 5 and allows optimising the filling time, which is about 9min in this example.

Equation 3 provides the relationship between filtration time and cake (here chamber) thickness, so that the

optimal filtration time can be calculated from this for changes in chamber thickness. Doubling the chamber

thickness increases the filtration time by four fold. Table 2 shows the test conditions, applied filtration times

and the optimized filter time for Test T2 based on end of the linear curve in Fig. 5 above and approximating

for Tests T1 and T3. Those approximated values are shown in above Fig. 5 by solid symbols. The optimal

filtration time of test T1 and T3 could thus have approximated purely from filtration theory and so are quite

similar to the actual optimum. The underestimation for the thinnest chamber (T1) is thought to be due to

neglecting the chamber filling time (prior to any filtrate flow).

Table 2 – Filter press test conditions and calculated optimal times

Test Unit T1 T2 T3

Chamber thickness mm 25 41 57

Feeding / filtration time (test) min 6 10 17.5

Optimal feeding / filtration time from Fig. 5 min 4.5 9 17

Optimal feeding / filtration time from theory (extrapolated from T2)

min 3.3 17.4


Thus, filtration theory can be used to estimate filtration times in particular if test work has not been performed

for different chamber thicknesses. Again this implies certain assumptions so that if possible, test work should

at least be used to confirm the major trends.

After feeding/filtration, cake squeezing and air blowing are usually applied. For selecting and designing the

most suitable filter press it is important to know the efficiency of the squeezing and the air blowing cycle for

cake moisture reduction. Thus, it is necessary to record the filtrate collected (best over time) for all cycles

including feeding/filtration, squeezing and the air blowing cycle. This allows the extrapolation backwards in

time of cake moisture at any time during the filtration cycle.

Fig. 6 shows the cake moisture for the test with 41mm chamber during squeezing at two different pressures

and air blowing for this tailing material. Significant moisture reduction is achieved with squeezing at 6 bars

while increasing the squeeze pressure to 15 bar increases the desaturation rate (larger slope downwards).

Thus, the tailings filter cake is compressible. Cake blowing in contrast does not lead to significant moisture

reduction and thus, is not considered useful after squeezing. Air blowing is only effective, if the air pressure

is larger than the capillary entrance pressure within the cake and 6 bar air pressure does obviously not

exceed the capillary entrance pressure of this cake formed of ultrafine tailings with partly plate-shaped

particles.

Such a plot as in Fig. 6 allows the selection of the appropriate squeezing time depending on cake moisture

requirements.

Fig 6 – Cake moisture as function of filtration time (excluding feeding) of squeezing and air blowing.

Fig. 7 is another example of the effect of squeezing and air blowing on the moisture of a fine base metal

concentrate filtered (fed) at up to 6 bar, squeezed for 1 min at 15 bar and air blown at 6 bar for another 6

minutes. The feeding cycle produces a cake within the chamber of 15.3% moisture which is reduced in


moisture by squeezing only to 14.9% as the base metal concentrate is virtually incompressible if compacted

sufficiently during feeding. Longer squeezing times do not reduce the moisture further. The main moisture

reduction occurs during the air blowing with the reduction in moisture leveling off with increasing air blowing

time, which is obvious from the larger downwards slope when starting cake blowing compared to when

starting squeezing.

Fig 7 – Cake moisture as function of total filtration time for feeding, squeezing and air blowing.

Thus, the effectiveness (moisture reduction over time) of cake squeezing and air blowing has to be tested

experimentally while recording the filtrate over time and back-extrapolating for the moisture provides

moisture reduction kinetics and allows to select optimal squeezing / drying time even if only a test with longer

time was undertaken.

The filter press size (filter area, chamber thickness and thus, cake volume) is selected and sized based on

first selecting the best suited chamber thickness based on the test work. Then follows the determination of

the total process cycle time (feeding/filtering, squeezing plus air blowing) plus the technical cycle time

(opening, cake discharge, cloth wash if required and closing) for the particular filter press design selected.

After this the required total filter press volume can be calculated from the solids duty, the total cycle time and

the dry cake density while accounting for the cake volume reduction during squeezing.

Common Candle, Pressure Leaf and Kelly Filter Test work Pressure filters utilising a pressure vessel filled with slurry and filtration elements such as candle, pressure

leaf and Kelly filters are used for liquor clarification of thickener overflow, filtrate from vacuum or pressure

filters. Direct filtration of slurries in these pressure filters can be applied if thickening is problematic or cake

production and liquor clarification within one filter is needed or beneficial. They are also fully enclosed and

allow very corrosive applications with plastic internals.

Determining candle, pressure leaf and Kelly filter sizing is quite similar to that for filter press sizing and is

based on the same two requirements, only this time the feeding pumps usually do not provide a restriction


and the specific liquid filtration rate is relevant as the main filter requirement is commonly production of clear

filtrate. This might require pre-coat, body-feed or the (usually experimental) selection of a suitable filter

media.

Again, the bench scale pressure filter should reflect the design of the full scale filter as closely as possible

and in particular, the shape of the filter element (leaf or candle) and the sealing of the filter cloth, both of

which are important for liquor polishing. For example this is achieved in the TSD, a laboratory candle type

filter simulating the DrM Fundabac® candle type filter (Chen, 1997) which is suitable for liquor clarification,

wet or dry cake discharge, cake washing, cake drying through compressed air and/or steam and automatic

cake discharge by snap-blow (for dry cake discharge) or flush-back (for wet cake discharge). Fig. 8 and 9

show the schematics and pictures of the candle design, bench scale and pilot scale filters.

Fig 8 – Schematic design of the Fundabac® candle filter (left) and candle design

for full and bench scale with cake discharge principal (right).

Fig 9 – Picture of the bench scale TSD candle filter in PVDF (left), cake pictures (centre top with candle

and bottom cake only) and pilot scale Fundabac® unit (right)

As the bench scale filter has the same design compared to the full scale filter, only smaller candle diameter

and length, it reflects the design scale-down well and even cake discharge properties (cake stickiness) can

be tested. The lab scale filter provides a worst case for cake discharge due to the smaller candle and thus,


lower cake weight and tighter radii and so provides only a conservative estimation of appropriateness for

automatic cake discharge.

Similarly, pressure leaf filters should be ideally simulated by pressure vessels with a leaf of similar design

inside. However, filtration test work using either vacuum Buchner funnel or pressure cells for pressure leaf

and candle filter design is better than no test work and are reasonable for determination of cloth selection

and liquid filtration rate but they do not reflect cake discharge properties.

Examples for filtration theory and filtration testing for candle filters, pressure leaf and Kelly filters The main filtration within these filters for cake formation occurs due to slurry pumping into the pressure

vessel containing the candles or leafs. The specific liquid filtration rate during this filtration cycle is described

by Equation 2. If constant pressure is applied, Equation 3 is valid. This is the case for the second part of the

filtration from when the maximal feed pressure is reached. This relationship is shown in Fig. 10 for an

example of Bayer liquor clarification (Bickert and Schumacher, 2008).

Fig 10 – Specific liquid filtration rate as function of filtration time measured in lab scale,

expressed by Equation 3. Typical plant flow and pressure is indicated.

Applications of filtration for liquor clarification in nickel or lithium (ANZPLAN, 2011) processing, precipitate

removal in a particular zirconia process (Alkane, 2005) or lithium process (Galaxy Resources, 2008) or

in rare earth flowsheets (SGS) partly require a pre-coat and/or body feed so the requirements for effective

liquor clarification filtration cannot be predicted easily and test work is required. Fig. 11 shows an example

for a precipitate removal step when using a very tight filter cloth, using pre-coat filter aid or both a pre-coat

and body feed filter aid.

Using pre-coat and body feed results in initially the highest filtration rates, using only pre-coat shows initially

average filtration rates while the test without filter aid and a less permeable filter cloth results in lower

filtration rates from the start. The filtration rates are similar from a filtration time of approximately 20min. The

mlc

l

l

RA

V rμ

pΔ

dt A

dV


higher values for the test without filter aid could be due to the less clear filtrate and thus, a cake formed with

less ultrafines i.e. higher permeability.

The filtration rate could be extrapolated for longer filtration times (and more filtrate cumulated) which is

shown in the left hand graph in Fig. 11 for the test with pre-coat and body feed using Equation 4 and solving

the quadratic equation. This is useful if not enough liquor is available for test work or if the cake formed in the

test was not sufficiently thick to be discharged. According to filtration theory, it takes twice the time to form a

cake that has double the thickness and generates double the filtrate which is what is being experienced

here. As per the examples above, the filtration can be extrapolated and predicted for changed operational

parameters from filtration theory which reduces the required number of tests required. The optimal cake

thickness and overall cycle time can then be selected using additional cycle times for vessel filling, cloth

cleaning, vessel emptying, cake washing, cake discharge and any other cycle steps.

Fig 11 – Specific liquid filtration rate as function of filtration time measured in lab scale without filter and with

filter aid as pre-coat only and pre-coat plus body feed (left) and Equation 3 applied

for predicting filtration and cake formation rate for longer filtration times.

CONCLUSIONS (CHECKLIST FOR TEST WORK AND SIZING) The following checklist summarises the various theoretical and practical evaluations and considerations

discussed in this paper.

General There are a number of general rules which apply to all filtration testing and the acquisition of design data as

follows;

A true reflection of the full scale filtration step is required covering all slurry, solids and liquid

properties, including but not limited to particle size, feed solids content, temperature, residence time

after chemical processes such as precipitation and crystallisation, entrapment of air from flotation

and other items.


Bench or pilot scale filtration equipment should be similar to the full scale equipment in design and

process. Filter cells with cake squeezing and air blowing for example should be used for sizing

membrane filter presses with an air blowing cycle.

Filtration theory can be used to extrapolate filtration results based on specific cake resistance and

filter media resistance where no test work covers those process parameters. It is important to

consider all assumptions, apply the appropriate equation (constant pressure or constant volume

filtration) and account for cake compressibility.

The effect of flocculants, coagulants or filter aids used as pre-coat or body-feed has to be evaluated

experimentally.

Vacuum filtration and hyperbaric continuous pressure filtration It is essential to record the cake formation time and apply different theoretical approaches to cake

formation, cake washing and cake desaturation / drying.

Filtration theory for constant pressure is suitable to predict filtration performance once the specific

cake resistance is measured for both cake formation and cake washing. In particular the effect of

variation of process parameters such as feed solids, cake thickness, filter speed and vacuum

pressure applied can be predicted without further test work. However, filtration test work, if sufficient

sample is available, should be applied to confirm the material behaves according to the assumptions

made in filtration theory.

True to the full scale equipment, bench scale test work for rotary filters such as disc and drum filters

require fixed ratios between cake formation and drying, while the cake formaton time on top fed

filters such as belt and pan filters is dependent on slurry properties.

Pressure filtration by filter presses Filtration test work data for filter presses is required for general equipment sizing due to the batch

nature of the operation of these presses thus needing data on different cycle times in regards to

feeding and filtration, squeezing, cake washing, cake blowing and technical times such as opening,

closing, cloth washing etc. Essential are the detailed cycle times and pressures applied in each of

these steps of the press operation.

‘Filtration rate’ or specific solids throughput rate can be either reported for

o the cake formation time only; or

o incorporating the process times (feeding, squeezing, cake blowing) only; or

o for the complete full scale filter times including technical times (longer feeding due to pump

capacity restrictions, full scale squeezing and cake blowing times and non-process /

technical times).


Thus it is important to specify how the value was derived (and on which basis) and it is best to

provide the cycle time split for each of these activities.

Cake thickness and final cake moisture effect filtration cycles and thus, filtration rate figures should

be stated in combination with the cake moisture achived and the cake thickness used, as the specific

solids throughput (filtration) rate drops significantly with increasing cake thickness and reduced cake

moisture.

Critical for filter press sizing is the total cycle time, the total filter press chamber volume, the dry cake

density – which is not the bulk density of the solids material - (or wet cake density in combination

with cake moisture) and the cake compression, so that test work (and data sheets for press sizing)

should report those together with the specific solids throughput rate to determine the required filter

area.

Pressure filtration by candle, leaf and Kelly filters Cake formation is the main filter cycle and the sizing is similar as for filter presses.

Filtration theory is suitable to extrapolate filtration after the maximum filtration pressure is applied

and constant pressure filtration occurs.

ACKNOWLEDGEMENT The authors would like to thank all companies and organisations for which they could undertake filtration test

work in the past, which allowed us to gain, test and apply the presented considerations and understanding.

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Feasibility Studies, MINPREX 2000 Melbourne, Vic, 11 - 13 September, P. 79-89

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Filtration Congress, Orlando, USA, 22.-25. April, Poster Session 5.1-5.8

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– 135

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27, No10, P. 1080-1084

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Improving the Efficiency of Test Filtration Tasks, in Proceeding of the 10th World Filtration Congress, 14-18 April,

Leizpig Germany, P. I81-85

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December 2002 P. 3-6

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[Accessed 30 March 2011]

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press dewatering, AFSS 22nd Tech. Conf. & Expo, Minneapolis, USA, 4-7th May

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of the 10th World Filtration Congress, 14-18 April, Leizpig Germany, P. I86-90

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Sutherland, K.S., 2005, Solid/Liquid Separation Equipment, Wiley-VCH, Weinheim

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165 – 173

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Advanced Technology, Oxford, UK 1st edition.

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Sensible Cost Cutting for Resource Projects

D Connelly1

1. MAusIMM(CP), Director/Principal Consulting Engineer, Mineral Engineering Technical Services Pty Ltd (METS), PO Box 3211, Perth WA 6832. Email: [email protected]

ABSTRACT As the resources sector becomes tougher with rising costs and cyclical metal prices, mining companies

need to continue to reduce their costs and do more with less. Many new resource projects fail to come in

under or on budget based on Feasibility Studies. Particularly, capital expenditure (CAPEX) proves

difficult to achieve for resource projects. A number of resource projects have failed because of

aggressive plant CAPEX cost cutting resulting in projects with no surge capacity and are unable to

achieve design throughput. In addition, the plants are not operable because of the omissions. In recent

times many businesses have also experienced the "costs" of simply implementing aggressive operating

"cost cutting” measures as a strategy for solving business performance problems. This paper looks at

technology driven, employment related, luxury cuts, department cuts and restructuring cuts. In the past

companies used an incremental approach based on the use of past budget information as an integral

part of the budget construction process going forward. The use of performance reports and Management

Information Systems (MIS) is examined along with the role of continuous improvement in achieving

sensible cost cutting.

INTRODUCTION All companies look for ways to reduce costs and increase profits, and resource companies are no

different. With the scale of mining projects, a cost saving of a few percent could mean millions of dollars

for the company. Budgets are to determine where money is spent and where it could possibly be saved.

This applies to all resource companies whether it is an exploration company looking to build a plant, an

engineering design company or a mining company producing minerals.

Historically, the resources industry has always gone through peaks and troughs. Understandably then

mining companies want to take advantage of the “good times” by maximising output of existing mines or

constructing new ones. Designing a plant to these favorable market conditions during a mining boom

can be a dangerous exercise. If the metal price suddenly drops, a low ore grade or high processing cost

which was previously viable could result in a mine closure. To avoid this, companies must look at cost

cutting measures to increase the viability of their project for a range of commodity prices.

A plant in the design stage has a much larger scope for cutting costs from the CAPEX and operating

expenditure (OPEX) through circuit optimisation or process changes than an existing operation.

However, overly aggressive process cuts to reduce the CAPEX may cause the process to run

inefficiently during operation and could reduce the overall availability of the plant, causing extra costs



and smaller revenue. Process changes and their operability implications must be thoroughly considered

by the design company. Mining companies should not simply cut costs to the design process as this may

cause greater costs to the project in the future (e.g. Murrin Murrin, Browns Project, Ravensthorpe Nickel

project, Hot briquetted Iron Project, Rapu Rapu Copper Zinc Project).

An existing operation which is struggling to be profitable may need to make some significant operating

cuts. Rather than one large cut, which could have serious negative implications to the personnel or

process, a number of small cuts should be implemented to achieve an overall significant saving. These

could be employment related cuts, luxury cuts, department cuts or restructuring cuts. These are

applicable to both engineering design companies and production companies. However, production

companies also have scope to refine the operating costs through power conservation, process

automation and consumables optimisation. It is important that, prior to implementation, all cuts be

assessed to ensure that the negative impacts are minimised. Continuous Improvement (CI) plays an

important role in the cost cutting procedure. Streamlining information communication through the use of

Management Information Systems (MIS) is just one aspect which could be included for continuous

improvement.

BUDGETING AND BUDGET CUTS Creating a budget should be the first stage of any project. It is a process of predicting and controlling the

expenditure over the life of any given project. Budgets are the foundation of an organisation’s financial

success. The importance of creating a budget is that it forces an organisation to consider the

expectation for its products and services with the required resources to meet that expectation. In

addition, budgets can transform an organization’s higher priorities into the appropriate resources

required to achieve those priorities. The potential problems could be highlighted in a sufficient time to

acquire the corrective actions to be performed. A baseline can be created against which the actual

results can be compared.

Due to the variability of the resources industry and the difficulty in predicting future costs, the budget is a

document which should be continually monitored. When a planned budget is overshot, the consequence

can be varied from mere frustration to anger or even litigation if it involves a new engineering design and

construction project with other companies. Overspending an allocated budget will result in the reputation

of the company being marred and good business relationships being severed. (See Fig 1 which

highlights impact of changing budget)



FIG 1- Cost of Not Getting Budget Right

Strategic Budget Cuts Strategic budget cutting is a common organisational policy. Applying the right combination of budget cutting

and strategic growth is a fundamental input for the long term success of any resource organisation. It is vital

for managers to approach budget cuts practically so as not to affect the organisation’s capability.

There are many ways to perform a budget cut, yet extreme cutbacks can eventually affect the growth of an

organisation. Many organisations today have come to understand how an overly enthusiastic implementation

of budget cuts can have the unintended effect of inhibiting their revenues. Profits were increased in the short

term due to the severe budget cut, but the overall revenues were then decreased due to the lack of growth.

In some cases, budget cuts have caused the subsequent declines in customer service and product quality.

There are a number of types of budget cuts and these include luxury cuts, employment status related cuts,

technology driven cuts, department cuts and restructuring cuts.

Luxury Cuts An approach to reduce the budget is to focus on the costs of supplies and services along with employee

related expenses. Travel and other related benefits spending can be minimised for instance, by employees

only being offered economy class when travelling instead of business. Also, many organisations often spend

a good deal of money on office spaces to impress their clients and competitors (e.g. use of hire cars, video

conference instead of travel, mobile phone calls, alcohol, dining expenses).



Employment Status Related Cuts By employing consultants and independent contractors, organisations will be able to minimise the overhead

salaries spent on permanent employees. Although independent contractor arrangements can make a

significant saving to an organization, it has some downsides that need to be taken into consideration. An

unintentional effect of changing a number of former employees into independent status may create

unfortunate tax complications in an organisation. Some organisations offer year- end bonuses instead of

pay-rises to avoid a fixed commitment. There is a growing trend to utilize temporary workers to support some

parts of an organisation’s operations such as during the shutdown period of an operating plant. Certainly,

temporary workers will never be as fully invested in the company as would permanent employees and this

could result in less efficiency and productivity.

Technology Driven Cuts With the advancement of technology, organisations have started to find ways to reduce the current workload.

The application of technology will help to reduce the amount of employees needed and other associated

costs to perform the tasks. It may even allow employees time to be in part redirected to improving the

efficiency of other areas of the operations. Although technology can be a good way to reduce the operating

budget, it is costly to establish and requires time to implement new technologies. The initial investment may

discourage some companies from implementing the technology. Technology is a capital investment and

consideration of the depreciation, ongoing maintenance and replacement has to be assessed adjacent to the

realistic savings. New technology introduces risks and is not applicable for new projects unless piloted first.

Department Cuts Department cuts are always a participatory process where managers of each operation unit will identify the

prospects to reduce their budgets without affecting the operation of an organisation. Managers can start by

checking the surplus funding in the past that may have been given but is not entirely essential to maintain

the level of services. When there is a vacant position available, managers will decide whether it can be held

without affecting the everyday operation of the organisation. Managers have to make sure that the budget

cut does not affect the organisation’s core strategies or the key clients, customers and constituencies’

interest.

Restructuring Cuts When an organisation undergoes a major restructure, this generally entails a shift from budgeting to strategic

planning. This change occurs as the demand from the market has changed. Restructuring an organization is

not an easy process and it involves the re-examination of the services it offers, re- evaluates the

departments and managers it should keep, the determination of the employees it needs depending on their

role in delivering new services and the application of technology as an alternative to additional employees in

its operations. The end result of the restructuring cut is usually a significant cost reduction. The

organizational changes in the 1990’s where whole layers of middle management were retrenched including

specialists with long term business knowledge is a specific example.



Continuous Improvement Continuous Improvement (CI) can be described in a number of ways. However, in the simplest terms CI is

the development of an ongoing improvement in quality and efficiency within a company. The objectives of CI

are to:

Provide a more disciplined approach to CI projects and initiatives;

Allow for greater consistency across global sites with a standard methodology and tools;

Allow for a quicker response to CI opportunities by facilitating the organization of teams and the

development of both teams and individuals;

Improve safety, cost, production, and productivity performance;

Help all employees understand the importance of, and the ways to, improve processes and

organizational relationships; and

Help capitalise on global best practices and share knowledge across sites.

The steps of CI are:

1. Identify the opportunity that will lead to continuous improvement

2. Study the opportunity with respect to the key business needs for the organisation

3. Define the current state by researching and understanding the current process, system or

organisation

4. Develop the future state then develop and test an implementation plan

5. Implement the solution and maintain high levels of communication to monitor the status of

scheduled activities

6. Follow up and document the new procedures that have been implements and assess the

effectiveness

All of these points together seek to improve existing operating procedures and improve quality and

efficiency.

Management Information System Management Information System (MIS) is the practice of managing data so that information can be delivered

with insight, understanding and value for the employees in an organisation. It is the process of organising an

information database which is easily accessible, well defined and flexible. Although techniques and

technologies will change, these principles will still remain in the core information management model.

Through the MIS, information is primarily delivered to the right person in the right structure at the right time



with a cost that adds net value to the organisation. For those that apply their knowledge gained through

experience it will produce a result with more effective outcomes.

MIS is an important tool to streamline communication throughout an organisation and serves to reduce time

generally used to search for information. This time saving ultimately makes an organisation more efficient

and hence more profitable.

DESIGN STAGE COST CUTTING Optimisation during the design stage of a project is an important measure in reducing the CAPEX of a new

project. However, reducing the CAPEX of the project does not mean reducing the cost spent on design.

Mining companies may be reluctant to spend large amounts of money on the design stage of a project due to

the fact that until production commences they will see no revenue for all their expenditure. Cutting corners at

this early stage, although saving in the short term could prove costly in the future. A high quality plant design

can ultimately save on plant CAPEX and reduce unexpected costs in the future.

Typically in the early stages of design for a mineral processing operation a number of options will be

investigated to determine the best possible design for the plant. These options will be compared on the basis

of process performance, CAPEX and OPEX.

Unit Design It can be tempting for mining companies to only design a plant to the nominal operating conditions in order

save on CAPEX. Ultimately this could be detrimental to plant performance and result in less availability or

throughput. Increased flow rates due to surges need to be accounted for around the plant and it is important

that the equipment is able to handle this. Pumps are a typical example of this and the design should

incorporate a 10-15% allowance for surges.

Another trap is designing equipment to handle the average ore characteristics rather than the maximum. Ore

hardness is one characteristic where the maximum value should be taken when designing equipment.

Although sizing to these specifications may result in larger, more expensive pieces of equipment being

required, designing the plant to handle average ore hardness can result in lower crusher throughputs, higher

recirculating loads and less overall plant utilisation.

Similarly, individual equipment availability needs to be taken into account. In a processing system with little

surge capacity the overall availability is equal to the lowest availability, not the average. If a critical piece of

equipment is off line the entire process stops. A high maintenance piece of equipment like belt filters may

have an availability of 65-70% due to cloth replacement and cleaning. If an overall plant availability of 90% is

desired, standby filters would be needed to achieve this. Removing these standby units in order to save on

CAPEX will ultimately result in lower plant throughputs and smaller revenue.

Alternatively if there is equipment which is known to have a low availability, but the capital required makes

the use of multiple units unviable, instead of expecting unreasonable availabilities these units could be

decoupled from the process. That is, placing adequate surge capacity before or after the units to allow for

the increased down time. These units will require a higher throughput than the rest of the processing plant to



maintain the overall plant throughput with less up time. A common example of this in a minerals processing

operation will be a crushing circuit which has been decoupled from the processing plant through the use of a

crushed ore stockpile. A crushing circuit will generally have an availability of around 80%.

Plant Layout Plant layout can affect the CAPEX for a project. Optimizing the plant location and plant layout can produce

significant saving when considering access and plant roads, material handling distances, electrical and

communications infrastructure, and plumbing. The design must be optimised with these factors in mind

without having the equipment so close that access for maintenance is compromised. Computer aided design

packages (i.e. SolidWorks) can be used to virtually construct the plant and be used to produce quantities.

Optimising the design in the virtual world, can enable direct savings in the real world. A more compact plant

but with room for later expansion flows through to savings on civil, piping, electrical and construction costs in

both the short and long terms.

Contingency Mineral processing projects in development will always have a contingency added to the price. Although this

may seem like an ambiguous addition to the price it is in fact an important item in the project cost. This cost

represents the project unknowns which experience tells us will eventuate as the project progresses.

Contingency is directly related to the risk of the project whether that is political, environmental or otherwise.

The more risk associated with a project the higher the contingency will need to be. For example, if a project

is to be constructed in an area with political instability or a tropical area that is prone to cyclones, an

allowance will need to be made in the event that the project becomes delayed or requires extra capital to

complete. As the project develops and gets closer to production, the chance of one of these events occurring

suddenly decrease thus the contingency does also.

The contingency for a project in the pre-feasibility stage may be between 20-30% of the direct costs.

However, as the project gets to the detailed design stage it would be expected that this number will drop to

10-20%. Of course this contingency must be determined on a case by case basis and these numbers can

vary. Estimating contingencies can be a black art. In order to reduce the CAPEX of the project it may be

tempting for some companies to reduce or remove the contingency. Although there is a possibility that the

contingency will not be needed, it is there for a reason and cannot simply be dismissed. Reducing or

removing this could prove costly in the event that the contingency is required. (See Fig 2)



FIG 2 – Value Adding Through Project Life

ENGINEERING SHORTCUTS Engineers have a duty to provide their services in a manner consistent with the "standard of care" of their

profession. A good working definition of the standard of care of a professional is that level or quality of

service ordinarily provided by other normally competent practitioners of good standing in that field,

contemporaneously providing similar services in the same locality and under the same circumstances. An

engineer's service need not be perfect. Since the engineer, when providing professional services, is using

judgment gained from experience and learning, and is usually providing those services in situations where a

certain amount of unknown or uncontrollable factors are common, some level of error in those services is

allowed.

When you hire an engineer you "purchase service, not insurance," so you are not justified in expecting

perfection or infallibility, only "reasonable care and competence". The fact that an engineer makes a mistake

that causes injury or damage is not sufficient to lead to professional liability on the part of the engineer. In

order for there to be professional liability, it must be proven that the services were professionally negligent,

that is, they fell beneath the standard of care of the profession. When one hires an engineer, one accepts the

risk, and the liability, of that professional making a mistake similar to mistakes other normally competent

engineers make, using reasonable diligence and their best judgment.

The standard of care is not what an engineer should have done in a particular instance, it is not what others

say an engineer would do, or what others say they themselves would have done, and it is just what

competent engineers actually did in similar circumstances.



Operability Operability is the ability to keep equipment, a system or a whole industrial installation in a safe and reliable

functioning condition, according to pre-defined operational requirements.

Has the designer considered the operator;

Distributed Control;

Slopes on sump floors- too flat is a problem;

Access to power, air and water points;

Walkways and access;

Flat launders- sanding;

Skirting- spillage;

Dust control;

Clean up in crushing plants;

Surge between unit operations;

Low head height;

Clear signage, pipes labelled;

Attention to particular areas- reagent mixing- lime mixing;

Floor space around mills etc;

Absence of or insufficient bunding height;

Insufficient process water storage;

Ball charging; and

Telemetry etc.

Maintainability In engineering, maintainability is the ease with which a product can be maintained in order to:

Correct defects;

Meet new requirements;

Make future maintenance or expansion easier, or



Cope with a change to maintain the plant;

Can the plant be maintained;

Access to pumps for repairs or change out;

Access equipment for mobile equipment and personnel for maintenance, clean up etc;

Prevent lubrication equipment under spillage or it being placed under points rather than well outside the area of operation;

Provide spare cyclones;

Have common equipment;

Install cranes in grinding area, flotation area;

Install bypass facilities (thickeners, trash screens);

Separate acid wash column or hopper; and

Avoid the obvious e.g. vent exhaust next to a high voltage (HV) switchyard etc.

CONSTRUCTION COST CUTTING

Labour Costs

Labour represents a significant percentage of the construction costs. When looking to reducing these costs

the “per hour” labour rate should not be the determining factor, rather the efficiency of the work force. It may

be tempting for some companies to hire cheaper unskilled or foreign workers for plant construction to reduce

the labour rate. However, if one of these workers is half the price but takes three times as long to do the

work it is not a cost effective decision.

Examples exist of projects where cheap foreign electricians have been used for African projects and poor

supervision resulting in extended commissioning problems, rewiring and damage to equipment.

Plant Modularisation Plant accommodation and site buildings have used modular design for some time now to save on

construction and installation costs. The same benefits are now being seen with plant and equipment

modularisation. Companies that produce modular equipment have a standard design and modify it as per the

process requirements, reducing engineering costs. The equipment is commissioned off site before being

dismantled and packed into shipping containers. Once at site the equipment can simply be bolted together

and it is ready for use. Using equipment modularisation can significantly reduce construction times and

hence capital cost.



Innovative Construction Techniques With construction costs representing a significant portion of the CAPEX, there are companies that are

marketing innovative techniques and materials to reduce the cost of the project. There are a number of

products and construction methods that look to reduce the amount of concrete used in construction. One

method uses earth retaining walls constructed from steel mesh. Another uses steel arches overlaid with

earth to form bridges, underpasses and stockpile tunnels. Both of these save time and money which is

associated with the concrete structures that would otherwise be required.

Reducing Construction Capital Through Unit Hire It is possible to reduce the CAPEX of a processing operation through the hire of units. However the reduced

initial cost comes at the price of an increased OPEX. This could be an option investigated by junior miners

looking to enter production but are finding it difficult to raise the initial capital. The mining fleet, mobile plant

and the accommodation camp are typical examples of units that are able to be hired for this purpose.

Second Hand Plants The option of using second hand equipment in the construction of a processing plant can mean considerable

savings to the overall CAPEX of the project. However, this needs to be considered on a case by case basis

as it is possible for the costs to outweigh the savings. The use of second hand equipment can be considered

at any stage of the project. However, due to the project timeline compared to the urgency of most second

hand sales, sourcing of appropriate equipment is generally done during the definitive feasibility study (DFS),

the last of the project before construction.

There are four factors that should be considered when determining the viability of using second hand

equipment, and these are equipment costs, equipment condition, suitability for process requirements and

spares.

Equipment Costs The equipment cost does not only include the purchase of the equipment but also the dismantling,

reconditioning, transportation and reassembling at site. This should be considered carefully as if the

equipment is located in a remote mine site or is difficult to dismantle and transport, it may be cheaper to

purchase new equipment.

Equipment Condition Although second hand equipment may be cheaper than new, the condition of the equipment needs to be

considered. If substantial refurbishing is going to be required, the initial saving on CAPEX will be outweighed

by the cost required for the plant to become operational. Also used equipment may require extra

maintenance during operation, and this possibility must be taken into account.



Suitability For Process Requirements Second hand equipment is unlikely to be the exact specifications as detailed by the proposed equipment list.

Some concessions may need to be made and it must be determined whether the equipment can handle the

process requirements.

The concept is fine however there are numerous examples where the execution has been poorly handled

resulting in the savings being less than originally considered. Lack of drawings, Vendor support, unknown

historical problems with the equipment are also issues.

Spare Parts This is a requirement to avoid the cost of having to specifically manufacture one off parts.

OPERATIONAL COST CUTTING It can be difficult to achieve any one cost cut on an existing operation that is going to make a great deal of

impact on the overall operating cost. Also a large cut will inevitably impact negatively on the process or

personnel of the organization. Therefore, the most sensible way of achieving any significant saving is to

implement a number of small cuts. With regards to resources, the sector can be broken up into two distinct

industries. These are mineral consulting companies and the production companies. Consulting companies

would use the strategic budget cuts as detailed previously to achieve costs cuts in the organisation. Where

production companies can also use these strategic budget cuts, they also have scope to reduce the

operating costs of the processing plant.

Reagent and Consumables Optimising the operating costs with regard to plant reagents and consumables starts at the first contract

tendering phase. The tenders need to be reviewed to obtain the best supplier of the product, and this may

not always be the cheapest option. The lead time for the products needs to be taken into account and also

the supplier’s production rate particularly if they are supplying to an ever increasing market, it is necessary

that supply can be maintained to the plant especially if that product is critical to the process.

This has been the case with cyanide when a global shortage impacted on the smaller supply companies.

The operations that held contracts with these smaller companies were paying enormous margins at the time

because of the difficulty to get the product. The larger companies on the other hand were able to maintain

supply to their clients but would not provide product to the companies that did not have an existing contract

in place.

It is also important to watch their usage, particularly important with the more expensive consumables, for

example cyanide in a gold plant. High usages should be investigated and lowered if possible. In some cases

it may be appropriate to use automatic control to moderate reagent and consumable use. It should be

investigated whether the capital outlay will be repaid through cost savings.



There are numerous examples of reducing cyanide, flocculant where considerable savings have been made

because technical people were prepared to challenge what had been used historically and undertake

testwork to support the case for reducing reagents.

Power Saving Using power efficiently and saving energy where possible can result in decent operating cost savings,

especially if this power supply is diesel generated. In some cases an external contractor can be brought in to

determine the savings achievable through power saving. However, it needs to be assessed where the cost to

bring an expert in is worth the savings.

One way of saving power is to ensure that the operating curves of process pumps are within the most

efficient range. Depending on the size of the pumps in the system this could equate to a significant cost

saving. There are numerous examples of this for SAG mills, crushing circuits etc by using circuit surveys and

simulation studies.

External Contractors It is important that costs for external contractors are monitored. It can be the case that if contractors are left

without close contract management, they may charge more than budgeted for the job. One provision would

be to never hire a contractor for a large job on an hourly rate. Over the life of the project the cost blowout

could be substantial.

Improved Control Improving the control systems of an existing processing plant does have the ability to increase efficiency of

the operation. However the capital to implement these changes could be significant. These improvements

could be related to automation of sampling, online analysis or automatic reagent addition. These methods

would be categorized under Technology Driven Cuts as mentioned previously. Although the improved

efficiency could reduce the operating costs this needs to be weighed against the often substantial capital that

must be outlaid for their installation. Flexibility to cope with changing ore types and flowsheets is a must.

CONCLUSIONS Sensible cost cutting is an important measure in ensuring the viability of an organisation through a range of

market conditions. This is particularly important in the resources industry where the cyclical metal prices

mean that boom and busts are inevitable. Cutting costs to remain profitable during the bust times is often

necessary to avoid business or mine closures. Careful budgeting, implementing an MIS and continuous

improvement are just a few of the measures that should be implemented to keep costs down at all times. If

budget cuts are required, large cuts should be avoided due to the detrimental effect upon personnel or the

process. A number of small, thought out cuts should be implemented, as discussed in this paper. These

small savings can cumulatively produce an overall significant saving to the CAPEX and/or OPEX costs.



ACKNOWLEDGEMENTS The author would like to thank Robert Hanna from METS who assisted and various companies, all

colleagues, engineers at various sites, METS staff and other consultants for their contribution and the

management of METS for their permission to publish this paper and the constructive criticism of various

drafts.

REFERENCES Australian and New Zealand College Of Anaesthetists (ANZCA), ‘How To Carry Out A Continuous

Improvement Project’, Guidelines On Continuous Quality Improvement [online]. Available from: <

http://www.anzca.edu.au/fpm/resources/educational-documents/guidelines-on-continuous-quality-

improvement.html> [Accessed 2 March 2011]

Mackenzie, W & Cusworth, N, 2007. ‘The Use and Abuse Of Feasibility Studies’, Project Evaluation Conference, Melbourne Victoria, 19-20 June 2007.

Maddox, D. Strategic Budget Cutting,TGCI Magazine [online]. Available from: <http://www.tgci.com/magazine/Strategic%20Budget%20Cutting.pdf> [Accessed 2 March 2011] Petty, J. Budgeting & One Day Reporting: Developing And Managing A Budget, Towards One Day Monthly Management Reporting, Course Notes.



Integrating Sustainability Principles into Mineral Processing Plant Design

G Corder1 and S Green2

1. MAusIMM(CP), Principal Research Fellow, Sustainable Minerals Institute, The University of Queensland, St Lucia Qld 4072. Email: [email protected]

2. Managing Director, SUSOP Pty Ltd, 93 Colin Street, West Perth WA 6005. Email: [email protected]

ABSTRACT Even though most major operating companies in the minerals industry have Board-level

endorsed sustainability principles and objectives, design engineers typically have no

mechanism that allows them to implement those objectives in their work. In addition, the

existing project management systems do not readily deliver the innovative solutions needed

to address key sustainability issues, such as those related to significantly reduced carbon

emissions, minimal environmental impacts, and maintaining the societal ‘licence to operate’.

SUSOP® (SUStainable OPerations) is an approach for the integration of sustainable

development principles into the design of mineral processes, which has been developed

through the Co-operative Research Centre for Sustainable Resource Processing (CSRP).

Somewhat analogous to the well entrenched HAZOP (Hazard and Operability) studies,

SUSOP® is a holistic, systematic and rigorous mechanism for identifying and assessing

sustainability opportunities and risks at each stage of project development - from corporate

planning, through design and operation, to decommissioning and rehabilitation – without

compromising financial rigour.

SUSOP® was developed, tested and enhanced through its application to real case studies in

the minerals industry. This development path provided significant insights into the process of

identifying and evaluating opportunities for improving an operation’s contribution to

sustainability and its long-term business case.

This paper will discuss the need for a methodical approach for integrating sustainability

principles into mineral processing plant design, present the key elements of SUSOP®, and

outline its development path through real case studies. In addition, the reasons why such an

approach will have growing importance into the future will be raised.



INTRODUCTION Mineral processing operations, now and into the future, will need to be fully prepared to deal

with the critical aspects of sustainability to ensure their continuing licence to operate (e.g.

community encroachment, increasing government pressures, increasing scarcity and

subsequent costs of raw materials, water, and energy). Major mining companies and

industry bodies, such as the International Council on Mining and Metals (ICMM, 2003) and

associated bodies such as the Minerals Council of Australia (MCA, 2005), have developed

and adopted Board-level endorsed sustainability principles and policies to help address

these challenges. However, practicing engineers have great difficulty in practically applying

these principles and policies into new mineral processing project developments.

To provide practical solutions to core sustainability issues and maintain the societal licence

to operate, the ideals of sustainability need to become embedded into project management

systems. To achieve this, a mechanism is required that will ensure consistency across

different projects and result in a more comprehensive understanding of the overall

contribution that a new mineral processing project can make to sustainable development.

This is crucial as the environment and local community for any new mineral processing

project offers a distinctive set of opportunities and sustainability hazards which need to be

identified and addressed in a consistent manner. Without such a mechanism as part of the

standard project management system, new mineral processing projects will continue to

replicate existing mineral processing operations with only modest enhancements to meet the

requirements of corporate sustainability principles.

This paper presents the key elements and development path of an emerging, holistic

mechanism for incorporating sustainable development principles into the design and

operation of mineral processing plants, called SUSOP® (SUStainable OPerations).

BACKGROUND In 2003, the Cooperative Research Centre for Sustainable Resource Processing (CSRP)

was formed, bringing together industry, government and researcher stakeholders to

investigate ways of improving the sustainable development performance of the Australian

and International minerals industry. A key initial focus was on several areas of a technical

nature for improving the efficiency of the industry and leading to less energy and water

consumption and less waste production and emissions. In parallel with this technical work,

researchers also investigated various “sustainability frameworks” to measure the

effectiveness of sustainability initiatives and drive better behaviours, in a similar manner to

safety analysis and statistics which supported the drive towards safer workplaces starting 30

plus years ago.



Through this research the CSRP discovered that even though most major operating

companies in the minerals industry have Board-level endorsed sustainability principles and

objectives, it was apparent that design engineers typically had no mechanism that would

allow them to incorporate those objectives in their work. Without a mechanism that

operating companies and government agencies could apply (either in-house or from

engineering consultancies), it was extremely difficult to identify opportunities for improving

sustainability outcomes and reducing risks as part of the business case for any proposed

course of action (such as a new project, plant expansion or any other business decision).

Likewise, it was determined that a suitable mechanism would need to fit within existing

business processes (as do safety, environmental and risk assessments) and provide an

easy to understand summary of the outcome on the complex issue of “how sustainable is my

business?”.

Similar to safety and risk reviews, such as HAZOP (Hazard and Operability) studies,

companies increasingly began to realise that business outcomes are actually improved if

sustainability issues are considered as an integral part of the business rather than an add-on

extra at the end of the process. This recognition became a turning point and thus the need

for such a mechanism, such as SUSOP®, became much more apparent.

KEY FEATURES OF SUSOP® SUSOP® is a guiding framework for projects that enables a proper contribution to

sustainability by the industrial facilities being studied, designed, built or operated through:

Robust integration of environmental design criteria into project processes,

Proper integration of social context into project processes,

Correct interfacing with legislative processes, Environmental Impact Statements

(EIS), and Health, Safety, Environment and Community (HSEC) best practice, and

Development of design and project management ideas.

Key Elements and Outcomes There are three major elements of the SUSOP® framework:

1) Sustainability opportunities and risks identification (SUSID™). A significant

characteristic of this element is that it includes ‘new ideas’ generation. This is the

most substantial element of SUSOP® and is made up of four steps:

Step (i) Familiarisation with Sustainability Concepts and Project Context,

Step (ii) Goal Scoping and Opportunities and Risks Identification,

Step (iii) Analysis of Sustainability Opportunities and Risks, and



Step (iv) Prioritisation of Sustainability Opportunities and Risks.

2) Sustainable Development (SD) Assessment – to conduct a detailed evaluation of the

shortlisted or high-priority opportunities and risks.

3) Decision Support - to provide assistance with decision-making at project toll gates.

All elements are supported by the SUSOP® Knowledge Base which includes information on

sustainability frameworks and principles, details on case studies, relevant SD tools and

databases to assist in the evaluation and assessment stages, resources for workshops and

relevant public domain information and data.

The main outputs from a SUSOP® study are presented in the Sustainability Register™,

which works in a similar manner to the conventional risk register and includes:

Opportunities for improving the contribution to societal sustainability and business

performance of the project,

Supporting SD Balance SheetsTM for top ranking opportunities to schematically show

the positive and negative impacts across the chosen sustainability framework (the

default is the Five Capitals Sustainability Framework which comprises natural,

human, social, manufactured, and financial capitals (FF, 2005),

Sustainability risks that could potentially impact on the project’s viability, and

Action plans for each item (opportunities and risks) in the register before proceeding

through the next project toll gates.

Description of the key elements and outcomes of the SUSOP® Framework are summarised

in Fig. 1.

Participants The mix and nature of the participants will vary depending on the stage that the project is at

and the type of project being considered. Typically core project staff (e.g. process engineers

and plant designers) plus environmental and community experts are part of the SUSOP®

study team. A trained SUSOP® leader directs the overall study and a trained SUSOP®

facilitator runs workshops. The team should include people of differing backgrounds,

experience and attitudes, and expansive, unconstrained, lateral thinkers should be included

and encouraged. This greatly enhances the potential for generating innovative and

alternative opportunities as well as identifying possible sustainability risks.



SU

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Key

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SU

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Step (i) - Familiarisation with Sustainability Concepts and Project Context

Aim: To create a common understanding of the project context and core sustainability knowledge amongst the SUSOP® Team

Participants: Full team

Linkages between sustainable development framework/principles and the project context and project stages are established.

Step (ii) - Goal Scoping and Opportunities & Risks Identification

Aim: To identify suitable SD goals for the project and opportunities to potentially achieve these goals and sustainability risks that might be counter to the SD goals, or sustainability principles, and therefore affect the project's viability.

Participants: Full team

Statement of SD Goals for distribution to the SUSOP® Participants which become the desired standard.Opportunities with their associated positive impact and achievability ratings and risks with their associated negative impact and likelihood ratings.

Step (iii) - Analysis of Sustainability Opportunities & Risks

Aim: To conduct further analysis to confirm identified Sustainability Opportunities and Risks and produce a prioritised list of linked opportunities (concepts) and risks.

Participants: SUSOP® Practitioners

A prioritised list of concepts (linked opportunities) and their associated Preliminary SD Balanced Sheets™.An updated list of sustainability risks and where appropriate indicating the associated concepts that could mitigate risks.

Step (iv) - Prioritisation of Sustainability Opportunities & Risks

Aim: To validate and confirm a shortlist of concepts (linked opportunities) and risks for future evaluation and assessment.

Participants: Full Team

A shortlist of top concepts (linked opportunities) and their associated SD Balanced Sheets™. An updated list of sustainability risks and where appropriate indicating any associated mitigating Concepts.

SUSOP® Key Element 2. SUSTAINABILITY ASSESSMENT

Aim: To conduct a detailed assessment of the identified shortlist of concepts (linked opportunities) and risks.

Participants: SUSOP® Lead and Practitioners and project personnel for review.

More accurate SD Balanced Sheets™Associated action plansDraft SUSOP® Report

SUSOP® Key Element 3. DECISION SUPPORT

Aim: To produce a final stage of the Sustainability Register™ inducing SD Balance Sheets™ to assist with decision-making at project toll gates.

Participants: SUSOP® Lead and Practitioners and project personnel for review.

Initiatives exhibiting enduring sustainability benefits and satisfying key business performance criteria.Final version of the Project Stage Sustainability Register™ including final project stage SD Balance Sheets™.Project Stage SUSOP® Report.

Key OutcomesKey Elements

Sustainability Register™

Sustainability Opportunities contributing to sustainability and business performanceSupporting SD Balance SheetsTM for top ranking opportunitiesSustainability risks that could impact project's viability.Action plans for each item (opportunities and risks) before proceeding through next project toll gates.

Fig. 1: Key Elements and Key Outcomes of SUSOP®



APPLYING THE SUSOP® FRAMEWORK In this section an example of the outputs from applying SUSOP® at the Concept stage of a

project are presented. In the early stages of a project, for example Concept, it may only be

feasible to conduct the SUSID™ element of a SUSOP® as limited data are available, making

it difficult or impossible to conduct detailed analysis which is typically done in the SUSOP®

Key Element 2: Sustainable Development Assessment. The example below is indicative of

applying SUSOP® at Concept stage.

SUSID™ Step (i) Familiarisation with Sustainability Concepts and Project An important component of the SUSID™ Step (i) is establishing linkages between the project

context and the sustainability framework as it provides the basis for the identification of

potential sustainability opportunities and risks. This is achieved using a workshop style

process, typically including the full complement of participants, so that a common

understanding of the project context and core sustainability knowledge amongst the

participants is created. Examples of linkages between the Sustainability Framework (Five

Capitals) and a project are:

Natural Capital – local eco-system impacts, generation of CO2, closure planning,

Human Capital – accessing or enhancing regional skills,

Social Capital – regional enterprise development,

Manufactured Capital – local or shared facilities/infrastructure (e.g. roads,

technology), and

Financial Capital – returns on investment for project.

SUSID™ Step (ii) Goal Scoping and Opportunities and Risks Identification Similar to step (i), this step utilises a workshop style process typically including the full

complement of participants. Sustainability goals are identified through a structured

prompting process. These goals should be compatible with the company’s corporate

sustainability principles and policies and provide the platform for the identification of

opportunities and risks. Examples of sustainability goals are:

Neutral carbon footprint,

Significantly enhanced local skills, and

Partnerships to create local enterprises.



Sustainability opportunities and risks are identified through structured techniques, including

specific prompting elements based on sustainability concepts, to identify opportunities that

could potentially achieve one or more of the sustainability goals and mitigate various risks

that may affect the project’s viability. Examples of sustainability opportunities associated

with the above goals are:

Alternative fuel or energy sources with a lower carbon footprint,

Training schemes to enhance the skills set of the local workforce, and

Potential local enterprises to support operation (e.g. tradespeople) or re-use/recycle

by-products (e.g. metal recycling).

Potential risks to a project relating to the broader sustainability issues should be identified

and, if possible, potential solutions to mitigate these risks should be proposed. For example,

initiatives might be identified under normal project management processes that appear

feasible from a techno-economic perspective but have the distinct likelihood of creating

community outrage, thereby putting the project at significant risk.

SUSID™ Step (iii) Analysis of Sustainability Opportunities and Risks SUSOP® leader and practitioners conduct further analysis to verify the outcomes from the

previous step and produce an initial shortlist of opportunities and risks. Opportunities that

could make the biggest contributions to sustainability, and risks that could potentially have

significant negative impacts on sustainability, are ranked highly in the short listing process.

Examples of short listing of the abovementioned opportunities with simplified supporting

analysis are:

Training schemes to enhance the skills set of the local workforce because it provides

critical benefits to the region,

Alternative fuel or energy sources with a lower carbon footprint because it could

reduce operating costs, and


by-products because it creates wealth in the local community.

SUSID™ Step (iv) - Prioritisation of Sustainability Opportunities and Risks Similar to steps (i) and (ii), a workshop with the same participants is undertaken to validate

the shortlist of opportunities and risks for future evaluation and assessment. The

participants review the prioritised list of opportunities and risks taking into consideration the



contextual setting (regional environmental and social consideration) for the project.

Examples of prioritising of abovementioned opportunities are:

Training schemes to enhance the skills set of the local workforce,

Alternative fuel or energy sources with a lower carbon footprint, and


by-products.

All identified opportunities and risks will be carried forward through the Sustainability

Register™, with prioritised opportunities and risks being identified. Eventually all

opportunities and risks on the Sustainability Register™ must be addressed (either

implemented or resolved).

SD Balance Sheet™ The purpose of an SD Balance Sheet™ is to schematically present the positive and negative

impacts of the opportunities and risks compared with the business-as-usual approach. An

example of an SD Balance Sheet™ is presented in Fig. 2. Using a consequence-based

approach to rate the magnitude of the negative and positive impacts, the SD Balance

Sheet™ concisely illustrates the comparative benefits of the outcomes from the SUSOP ®

process versus the standard outcomes from the business-as- usual approaches. In the

example SD Balance Sheet™ in Fig. 2, outcomes from applying SD principles compared

with outcomes from the business-as-usual approaches show:

There is less of a negative impact on natural capital as measures have been adopted

to reduce the environmental impacts e,g. re-using or reducing of waste or tailings,

Initiatives would be implemented to improve social capital e.g. local enterprise

developments,

Improved financial benefits due to lower costs from initiatives e.g. utilising waste heat

sources, and

Manufactured capital (that is, infrastructure) has lower positive impact which may be

due to directing capital towards initiatives identified through the SUSOP® process.



Fig. 2: SD Balance Sheet™ Example (Note: change in impact of capital for application of SD principles

is with reference to the impact of respective capital in business-as-usual e.g. impact of natural capital for

business-as-usual is -3.5 compared with impact of natural capital for application of SD principles, which

is -1, and therefore the bar is shaded light grey.)

Distinguishing Feature Unlike sustainability assessment tools which essentially evaluate the sustainability

credentials of several pre-determined options, the SUSID™ element of the SUSOP®

generates new options or opportunities and identifies potential risks. From the outset, the

development team stressed that this feature be included in the eventual framework. By

having the SUSID™ element, SUSOP® generates outcomes and benefits to business

performance that would not normally result from standard project management systems.

This is one of the key features of SUSOP® which distinguishes it from other sustainability

assessment tools and techniques currently available.

THE DEVELOPMENT OF A SUSTAINABILITY FRAMEWORK The section gives an overview of some of the development decisions from the genesis of

SUSOP® to its current maturity which is detailed in SUSOP® Foundation Manual (SUSOP

Pty Ltd., 2010). Towards the end of this section the learnings and enhancements to the

framework that resulted from two major ‘live’ case studies are discussed. This was a critical



component of SUSOP®’s development and helped build the necessary credibility with

industry.

Collaborative Team SUSOP® was developed through a collaboration of researchers and practising engineering

consultants, with the express intention of creating a mechanism that incorporated the

principles of sustainability into process design and operations. Being developed through a

Co-operative Research Centre, CSRP, it was possible to draw on the expertise of both the

industry participants and research providers which ensured a balance between real-world

practicalities and academic rigour.

In the early stages of development, a key research question for developing SUSOP® was

formulated from the outcomes of a major review of the design for sustainability techniques

and tools (McLellan et al., 2009):

“What is the most rigorous and defensible mechanism that enables sustainable development

principles to be incorporated into the design and operation of resource processing?”

Throughout the development process, the collaborative development team frequently

referred to this question to ensure that the resulting SUSOP® mechanism included these

qualities while still being a practical and workable approach that is accepted by industry.

Project–Production Cycle In developing the mechanism, the collaborative research team specified that it must work in

conjunction with the current practices and constraints in the industry and be integrated into

the standard project development cycle. In achieving this, the collaborative research team

considered it critically important that SUSOP® could be applied to all stages of the Project-

Production Cycle - from corporate planning, through design and operation, to

decommissioning and rehabilitation (Green et al., 2010). This would enable it to have the

flexibility to provide the appropriate guidance, tools and deliverables to fit the level of detail

and opportunity available at each phase of a project (Corder et al., 2010). As a result, the

application of SUSOP® is designed to work across the full Project-Production Cycle with

specific tailoring at each project phase.

Potential for Impact The potential for a mineral processing project to make a proper contribution to sustainability

is closely related to the project stage. As the ability to impact overall project outcomes is

greatest in the initial stages of the Project Cycle (Havranek, 1999 and Termini, 1999), the

earlier SUSOP® is applied the greater the chance that the project will make a significant



contribution to sustainability as illustrated in Fig. 3. The opportunity to create consequential

change diminishes as the project progresses, but can still be achieved to an extent, even if

the costs are higher and the resulting benefits less. In designing SUSOP®, the collaborative

team have taken this into consideration so that it will allow for the greatest potential for

developing sustainability opportunities or identifying risks within the known constraints and

limitations of a particular project stage.

Fig. 3: Ability to Impact on Outcomes versus Project Stage (based on the level of Engineering definition

(AACE, 2005))



Case Study Approach From the beginning the collaborative team made a decision that SUSOP® should be

developed using a case study approach. The reasoning was that the application of SUSOP®

to ‘live’ case studies would deliver considerable learnings which would greatly help in

enhancing and refining the mechanism and would help build credibility with industry. Two

‘live’ case studies were undertaken at the Concept and Pre-feasibility stages. From an

industry perspective, applying SUSOP® to the two case studies produced sustainability

opportunities and identified constraints that assisted the decision making in the subsequent

project development phases. These case studies played a major part in significantly

improving the SUSOP® mechanism and learnings from these case studies are presented in

the next section. Details on these case studies have been reported elsewhere (Green et al.,

2010) and summaries of the context and outcomes are provided in Table 1.

Table 1 Summaries of Case Studies (Green et al., 2010)

Case Study #1 – (Project in Concept phase) Case Study #2 – (Project in Prefeasibility

phase)

Context

Deadline for site selection due to pending expiration of retention leases.

Standard business practices could not provide definitive guidance on which sites to retain and which to forgo.

SUSOP® was applied to make the distinction on a sustainability basis.

Sustainability performance of a range of proposed effluent treatment options at a mineral processing operation was examined.

The aim was to apply the SUSOP® mechanism to identify sustainability opportunities and evaluate benefits and impacts of the proposed options.

Outcomes

SUSOP® produced a staged integrated plan for implementing renewable energy and local skills to avoid high energy and high expatriate labour costs.

Plan supports the aim of developing a low operating cost operation and several sustainable development benefits.

This, along with the other key identified opportunities, was the basis for making a recommendation on the most favourable site location.

SUSOP® generated sustainability opportunities to support effluent management.

Examples included creating local enterprises to provide consumables, producing ‘green’ by-products and implementing a small-scale wetland to replace a more conventional effluent treatment facility.

The indentified opportunities would substantially satisfy key elements of the company’s Sustainability Policy.

Case Study Learnings Each case study produced learnings which were then incorporated into SUSOP® to make it

considerably more robust and effective in generating opportunities for delivering both

sustainable development benefits and improved business performance. It is highly unlikely



that these learnings and subsequent enhancements would have emerged without the setting

of real projects at the Concept and Pre-feasibility stages. In short the learnings focussed on:

Importance of understanding the project context,

Need to include outputs into project management system through Sustainability

Register™,

Need to categorise opportunities in terms of impact and timing, and

Value of clustering related opportunities into ‘concepts’.

More explanation of the key learnings is presented in Table 2.

Table 2 Summary of Case Study Learnings

Project Context

Participants need to understand and appreciate the true context of the project under consideration.

As much information as practically possible should be made available to the SUSOP® study team, including the project history, the options already covered and the relationship of this project to other projects in the company’s portfolio.

The team can then consider and draw out the potential implications of the project on the surrounding community and environment and the wider range of stakeholders and identify potential sustainability/business risks.

Recording

Recording and reporting of risks and opportunities is critical for integrating into the ongoing project plan.

The need to flag sustainability “highlights” and potential “show stoppers” guides project development.

This resulted in the formation of the Sustainability Register™ which is like the conventional risk register, and is treated in the same manner, and ensures a path into the project management system.

Categorisation

The need to sort opportunities into different categories rather than produce a single (long) prioritised list.

This approach assists the client in understanding the timing and scope for implementing these opportunities.

For example, converting a boiler from coal to gas fired is relatively straightforward technically but has limited sustainability benefits while utilising waste heat from a pyro-metallurgical option might have a much larger sustainability benefit but is not easily achievable with current technology.

Clustering

• Opportunities should be clustered into ‘concepts’.

• This provides a more robust approach for understanding and appreciating the dependencies between the individual opportunities and the overall sustainable development benefits.

• For example, Fig. 4 illustrates how three opportunities can be linked together into a concept, ‘solutions to reduce greenhouse gas emissions’, which would support several aspects of the sustainability frameworks.



Fig. 4 Simple Example of Linking Opportunities into Concepts

GROWING FUTURE SIGNIFICANCE As mentioned earlier in this paper, major mining companies have adopted Board endorsed

sustainability principles and/or policies articulating their commitment to sustainable

development. To show real support to the ideals of sustainability, developers of new mineral

processing projects need to adopt approaches within their project management systems that

generate initiatives to enhance the contribution that the resulting operation will make to

sustainable development. This goes beyond being compliant with minimum requirements, or

in colloquial terms, just ‘ticking off’ against each principle once the project is complete. The

sustainability principles or policies should be the basis for driving innovative solutions into

new projects.

As SUSOP® requires a commitment from the project developer, it is important to recognise

the benefits and value that SUSOP® will generate over and above the standard project

management systems. If the benefits strongly outweigh the effort then adoption of such a

mechanism is warranted. For instance, consider the HAZOP process; although it can be

time-consuming, HAZOP is accepted and ingrained into the current project management

systems because practitioners appreciate that it makes an effective and efficient contribution

to a safe operation. Similarly, SUSOP® is building recognition with the industry as an

approach that delivers both sustainability benefits plus improved business performance over

the lifetime of the operation and beyond. As a result of research and development of

SUSOP®, some key factors have been indentified that are considered essential

characteristics for an SD mechanism to be widely adopted as a matter of practice when

developing a new mineral processing project. These are presented in Table 3.



Table 3: Essential Features of SD Mechanism for New Mining Projects

Financial Rigour

No compromise is made on defined financial objectives (e.g. net present value, rates of return on investment, etc.).

It should help reduce inherent or hidden business risks and thereby secure financial outcomes.

Licence to Operate

A more robust licence to operate is secured by identifying early on in the project life, social, human and environmental considerations or issues.

These can then be “engineered out” or “engineered in” during the project development rather than “managed” once the project is complete.

Holistic approach

Employ a cross-disciplinary team to assess a project with respect to the wider ecological and social context in which the project will operate.

Allow project decision makers to identify and review innovative and integrated solutions that incorporate environmental and social contexts.

Sustainability Showstoppers

Elicit critical sustainability issues that could be technically feasible and financially attractive but might, for example, generate significant stakeholder concern or even outrage and affect the operation’s social licence to operate.

In a conventional project management process these issues are likely to progress unresolved, potentially leading to a major business risk once the project is operational.

Ideas Generation

A structured approach to generating new ideas is far more likely to result in better outcomes than relying on chance to come up with value-adding innovations.

Consider and investigate alternative, price competitive resources that could improve an operation’s viability. Examples are alternative energy sources and industrial by-products, the value of which is likely to be enhanced by upfront decision making.

CONCLUSIONS Through the Cooperative Research Centre for Sustainable Resource Processing (CSRP), a

collaborative research effort between industry and research organisations identified the key

characteristics of a mechanism for incorporating the ideals of sustainability into project

management systems. These include such characteristics as a holistic approach employing

a cross-disciplinary team; no compromise on financial rigour; innovative solutions to

“engineer out or in” environmental and social issues; identification of sustainability risks that

could translate to business risks; and generation of new ideas to support innovation. In

addition, the concept of a sustainability register, similar to a conventional risk register,

ensures ongoing review in subsequent project stages until all identified risks and

opportunities are either implemented or adequately addressed.

These characteristics have been incorporated into the development of SUSOP®, a guiding

framework for projects that enables a proper contribution to sustainability by the industrial

facilities being studied, designed, built or operated. SUSOP® can be applied to any stage of

the project cycle - from corporate planning, through design and operation, to

decommissioning and rehabilitation stages – and its development has been enhanced

through the application to real case studies at the Concept and Pre-feasibility project stages.



Implementation of SUSOP® aims to give greater certainty in developing resource projects

that can significantly enhance the contribution that industry can make to society and

sustainable development.

The first edition of the SUSOP® Foundation Manual, which focuses on the Concept and Pre-

feasibility project stages, was published in December 2010.

ACKNOWLEDGEMENTS The development and trial application of SUSOP® was carried out under the auspice and

with the financial support of the Co-operative Research Centre for Sustainable Resource

Processing (www.csrp.com.au), which was established and supported under the Australian

Government’s Cooperative Research Centres Program. The SUSOP® research and

development team comprises industry participants, Hatch and GHD, and research providers,

the University of Queensland, University of Technology Sydney and CSIRO.

REFERENCES AACE 2005, Cost estimate classification system - as applied in engineering, procurement and construction for the process industries, Association for the Advancement of Cost Engineering.

Corder, G. D., McLellan, B. C. and Green, S. 2010, 'Incorporating sustainable development principles into minerals processing design and operation: SUSOP®', Minerals Engineering, vol. 23, no. 3, pp. 175-81.

FF 2005, Forum for the Future Website, Forum for the Future, viewed December 15 2006, <http://www.forumforthefuture.org.uk/>.

Green, S. R., Corder, G. D., McLellan, B. C., van Beers, D. and Bangerter, P. J. 2010, SUSOP®: Embedding Sustainable Development Principles into the Design and Operation of Resource Extraction and Processing Operations, The AusIMM, Kalgoorlie, WA, 17-19 August 2010.

Havranek, T. J. 1999, Modern project management techniques for the environmental remediation industry, St Lucie Press, Boca Raton, Florida, USA.

ICMM 2003, ICMM Sustainable Development Framework, International Council on Mining and Metals, viewed March 7 2011, <http://www.icmm.com/our-work/sustainable-development-framework>.

MCA 2005, Enduring Value: An Australian Minerals Industry Framework for Sustainable Development, Minerals Council of Australia, viewed 24 February 2011, <http://www.minerals.org.au/__data/assets/pdf_file/0006/19833/EV_SummaryBooklet_June2005.pdf>.

McLellan, B. C., Corder, G. D., Giurco, D. and Green, S. 2009, 'Incorporating sustainable development in the design of mineral processing operations - Review and analysis of current approaches', Journal of Cleaner Production, vol. 17, no. 16, pp. 1414-25.

SUSOP Pty Ltd. 2010, SUSOP® Foundation Manual - Part A Introductory Manual.

Termini, M. J. 1999, Strategic project management: tools and techniques for planning, decision making and implementation, Society of Manufacturing Engineers, Dearborn, Mich.



Chirano Gold Mine Expansion – A Case Study

S Ellis1 and I Dunlop2

1. MAusIMM, Lead Process Engineer, Lycopodium Minerals QLD Pty Ltd, 163 Leichhardt Street, Spring Hill Qld 4000. Email: [email protected]

2. Process Plant Manager, Chirano Gold Mine, A Kinross Company, PO Box 57, Bibiani, WR, Ghana. Email: [email protected]

ABSTRACT In January 2008 Chirano Gold Mines Limited committed to expansion of the Chirano Gold Project from its

existing design capacity of 2 Mtpa ROM ore to a new design value of 3.5 Mtpa. In addition to increased

throughput, the design had to contend with increasing ore hardness and competency, while at the same time

make cost effective use of existing equipment with minimum disruption to existing production. The solution

adopted involved some novel approaches including tertiary crushing and operation of the SABC circuit in

overflow ball mill mode and parallel elution with a common acid wash column. This paper describes the

approach taken to the design of these elements to achieve a cost effective and safe capacity expansion.

INTRODUCTION The Chirano Gold Mines Limited, CGML, treatment plant was originally designed to treat ore from nine open

pits (Obra, Paboase, Tano, Sariehu, Mamnao, Akoti North, Suraw, Akwa and Akoti South) at a rate of 2 Mtpa

and commenced operation in September 2005. Ongoing exploration work on the lease has identified the

potential for Chirano to develop into a long life high grade underground mine, with its first underground mine,

Akwaaba Deeps, on track to ramp up and further potential identified for the development of a second

underground mine at Paboase South.

In January 2008 CGML awarded an EPCM contract to Lycopodium Minerals QLD Pty Ltd to expand the

existing operation to a nominal feed rate of 3.5 Mtpa. Site construction work commenced in September 2008

and commissioning of the expanded plant was completed in September 2009.

EXPANSION COMMINUTION CIRCUIT DESIGN AND SELECTION The flowsheet for the original 2 Mtpa Chirano plant (Fig. 1) was based upon a conventional single stage

primary crusher and a two-stage SABC milling circuit. The basis of design assumed that the majority of ore

feed to the plant would be primary ore from the Obra open cut and design criteria for the comminution circuit

was based on Obra primary ore parameters. However, as the mine developed it was found that some

sources of ore, especially from the Tano pit, were significantly harder and more competent than the original

Obra ore design basis. Consequently for the expansion, in addition to the increased capacity, the design of

the comminution section of the plant had to accommodate an increase in ore hardness/competence. The

challenge that this presented for design was the selection of a flowsheet that provided both the increased

capacity with minimum disruption to the existing operation during construction, and at the same time made

cost effective use of the existing equipment.



Fig. 1 – Original Comminution Circuit Schematic

Design criteria and ore parameters for the original design and the expansion cases are summarised in Table

1. The 2 Mtpa comminution flowsheet incorporated primary crushing in a single Nordberg C140 jaw crusher

followed by primary grinding in an open circuit 6 x 6.44 m EGL SAG mill, and then secondary grinding in a

4.9 x 6.8 m EGL ball mill in closed circuit with twelve cyclones. The SAG mill originally operated with 25 mm

slotted grates and 80 mm pebble ports, with pebbles recycled via a Symons 4”1/4’ SH pebble crushing

circuit. The original ball mill circuit included gravity recovery from cyclone underflow via a Knelson

concentrator but due to insufficient coarse gold in the feed this circuit was subsequently decommissioned.



Table 1 – Design Criteria Parameters

Parameter Unit Original Design Expansion

Operating time (per year) hours 8000 8000

Ore Throughput (dry tonnes) Mtpa 2 3.5

tph 250 437.5

Grinding Circuit Feed F80 mm 110 12.5

Grinding Circuit Product P80 um 106 106

UCS design MPa 139 139

UCS range MPa 37 - 139 37 - 139

Rod Wi design kWh/t 21.9 21.9

Rod Wi range kWh/t 19.5 – 21.9 19.5 – 21.9

Ball Wi design kWh/t 15.9 17.6

Ball Wi range kWh/t 11.5 – 15.9 15.7 - 18

Ore specific gravity t/m3 2.77 2.77

JK SimMet Parameters, A x b 39.2 33.5

Abrasion Index, Ai 0.96 0.96

Pre-crushing ahead of SAG milling circuits is a demonstrated method of increasing the capacity of existing

circuits with numerous examples cited over the past decade, e.g. MacNevin, 1997, Atasoy et al., 2001,

McGhee et al., 2001, Putland et al., 2004, Mwehonge, 2006 and Thong et al., 2006. According to Siddall

and Putland, 2007, operations where secondary crushing was retro-fitted often show capacity increases in

excess of 40%, however the finer feed can result in a coarser SAG product shifting grinding duty to the

secondary ball mill.

Previous expansions cited have used either partial feed crushing or full feed crushing. There is a case for

retaining some coarse material as media to a SAG circuit by partial crushing, in that higher efficiency and

lower operating costs should theoretically be achievable. However this saving has to be balanced against the

increased capital cost required to manage and distribute the coarse media and the lost production resulting

from the integration of these facilities with existing operations during construction. In the case of the Chirano

expansion, the latter were seen as a significant risk and full feed crushing was adopted as the preferred

option.

Studies by OMC (Davies, 2007) on optimal crushed product size to feed the existing milling circuit,

suggested that a fine product size, and as a result, tertiary crushing would be required, as shown in Fig. 2.



250

255

260

265

270

275

280

285

290

295

300

305

0 10 20 30 40 50

Crushing Plant F80 (mm)

Mill

Th

rou

gh

pu

t (

ton

ne

s p

er

ho

ur)

Fig. 2 – Throughput vs Crushing Plant Product Size

Additional modelling studies of tertiary crushing and an expanded milling circuit for a capacity of 3.2 Mtpa

were undertaken by SMCC, 2007. The results indicated that the crushing plant would be required to

generate a product size P80 of 12.5 mm. A duplicate secondary ball mill circuit was also required to achieve a

product P80 of 106 um. For this purpose CGML procured a secondary and tertiary crushing plant, one CS660

secondary and three CH660 tertiary units and associated conveyors and screens, and a third mill. The new

mill was an unused low aspect SAG mill.

Circuit power modelling was then undertaken by OMC (Koch and Putland, 2008) which indicated that

installation of the secondary and tertiary crushing plant, and integration of the new mill, operating in “ball mill

mode” could achieve the design throughput of 3.5 Mtpa , (438 tph), although some relaxation of product P80

was considered possible.

With the finer feed material, the existing primary SAG mill required a high ball load (20 - 24% v/v) and

conversion to overflow configuration was also recommended. The new mill was installed in parallel with the

existing closed circuit secondary ball mill and operated in “ball mill mode” with a 20% v/v ball charge.

Conversion to overflow configuration by removal of the grates and lifters was also recommended.

The two secondary mills were designed to operate in closed circuit with a single common cyclone cluster, the

different mill power draws necessitating a 40 / 60 split of cyclone underflow return. The option to return

cyclone underflow to the primary mill was also retained to provide flexibility in balancing the power between

the three mills if required. This circuit design was expected to result in the pebble crusher becoming

redundant in the expansion flowsheet. The final mill design configuration selected is shown in Table 2.



Table 2 – Expansion Mill Configuration

Specification Units Primary SAG

mill (existing)

Secondary Ball Mill

(existing)

Secondary Ball Mill

(new)

Diameter- inside shell m 6.0 4.88 6.1

Effective Grinding Length m 6.44 6.83 7.3

L:D ratio 1.07 1.4 1.2

Lining Type Rubber Rubber Rubber

Feed Trunnion Inside Diameter mm 1850 1330

Discharge Trunnion Inside Diameter

mm 2080

Liner Thickness mm 140 80 80

Discharge Type Overflow Overflow Overflow

Ball Load (operating)* % v/v 19 30 20

Max Load (worn liners) % v/v 20 32 21

Mill Speed % Nc 67.5 – 75

DC Drive 75 75

Pinion Power

Operating 67.5% Nc

72% Nc

75% Nc

kW

2226

2370

2466

2145

3055

Pinion Power Max kW 2680 2370 3312

Installed Power kW 3200 2600 3500

Maximum Design Power kW 2900 2500 3500

* Note: 5% contingency allowed between operating and maximum ball charge.

The expansion circuit flowsheet is shown in Fig. 3.



Fig. 3 – Expansion Comminution Circuit Schematic

COMMINUTION CIRCUIT PERFORMANCE The expanded comminution circuit was commissioned in August 2009. Initially, the two modified SAG mills

were commissioned and operated as grate discharge mills whilst the client’s crushing circuit was also being

commissioned and operated. However, issues with reaching the crushing circuit design throughput and

product have impacted on the performance of the grinding circuit. Whilst the crushing circuit issues are

being resolved, the crushing circuit product has been relaxed from P100 of 15mm to 25mm. Consequently,

the primary SAG mill remains fitted with steel liners and a grate discharge configuration. The pebble crusher

has since been replaced with a HP200 unit fitted with extra fine liners to crush a closed side setting of 11mm.

Additionally, the primary SAG mill discharge grates aperture has been changed from 35mm to a combination

of 45mm (3 off) and 25mm (17 off). The larger 45mm aperture ports allow for an accumulation of oversize to

discharge from the mill to the pebble crusher.

Whilst the operation of the crushing circuit and primary SAG mill have differed from design configuration, the

performance of the ball mills has been in-line with the design. The operation of the existing secondary mill

has continued unchanged as per design and the new secondary ball mill has been brought on-line as

anticipated. Initially the new secondary ball mill was operated with a power draw of 3000 – 3200kW and the

resulting cyclone overflow stream sizing was over 90% passing 106 um. This particle size distribution

exceeded design and a lower power draw of 2800 – 3000kW with the new secondary mill is producing a

cyclone overflow stream P80 of 106 um.



THICKENER DESIGN AND PERFORMANCE The original thickener was installed such that it could be utilised as a leach feed thickener or a tailings

thickener. Its primary duty was a leach feed thickener, however in the event of high oxide clay ore, the

pipework was arranged such that the thickener’s duty could be changed to tailings. The original thickener

installed was a high rate type with gravity discharge. However, operational difficulties prevented this piece of

equipment from being fully utilised. The expansion scope included recommissioning this thickener along

with the installation of thickener underflow pumps and replacement of instrumentation. The flocculant mixing

plant was identified to have sufficient capacity for the expanded throughput.

Table 3 – Expansion Thickener Parameters

Parameter Unit Original Design Expansion

Ore Throughput (dry tonnes) tph 250 437.5

Solids Loading t/m2/h 1.25 1.25

Required Minimum Area m2 200 350

Required Diameter m 16 21

Actual Diameter m 18 18

Bypass Required % 0 30

Table 3 shows that the existing thickener size was insufficient to handle 100% of the expanded flowrate. A

system was retrofitted to direct 70% of the feed stream to the thickener and constantly bypass the remainder

of the feed stream. This was achieved through the use of a manually adjustable pinch valve to enable the

operator to manipulate the proportion bypassing the thickener as required. This solution was designed to

balance capital expenditure with effective utilisation of existing equipment.

Since the thickener was recommissioned, on-site flocculant trials have resulted in a more appropriate

reagent selection along with elevated pH dosing using lime. This has improved the thickener performance to

the extent that 100% of the feed is now directed to the thickener and the bypass valve has become

redundant. The design target for the thickener underflow slurry density was 50% solids (w/w) and with

improved thickener performance a thickener underflow slurry density of 52 – 54% solids (w/w) has been able

to be consistently achieved with a clear supernatant overflow and absolute minimal rake torque.

DESORPTION DESIGN AND PERFORMANCE The original flowsheet for the desorption circuit was based on a single pressure Zadra elution circuit being

operated six times per week. The original single pressure Zadra elution circuit consisted of one loaded

carbon recovery screen, one acid wash column, one elution column, one eluate heater and three

electrowinning cells. This circuit processes carbon in 5 tonne batches. The loaded carbon gravitates from

the loaded carbon recovery screen to the acid wash column where acid washing and water rinsing is



performed. Once rinsed, the carbon is transferred to the elution column for gold elution. During the elution

process, the elution column and eluate heater are in closed circuit with the electrowinning cell to strip the

gold from the carbon into solution for recovery by electrowinning onto the stainless steel cathodes. The

eluted carbon is reactivated prior to return to the CIL tanks. The loaded stainless steel cathodes are high

pressure cleaned and the sludge is recovered by filtering. The dried sludge is subsequently smelted in a

diesel fired furnace to doré bars.

The Zadra elution circuit consists of numerous batch steps, with the original design of the longest step of

elution being for 14 – 16 hours when the elution column and eluate heater are in closed circuit with the

electrowinning cell. The expansion required an increased mass of loaded carbon to be processed by the

desorption circuit. For this batch process, this equated to an increased frequency of operation of the

desorption circuit from six to twelve times per week. Therefore duplication of the desorption circuit was

required for the expansion, namely a second elution column, another eluate heater and one electrowinning

cell. A review of the sequencing of the batch steps presented an opportunity for the loaded carbon recovery

screen and acid wash column to be utilised for both elution columns as long as satisfactory safeguards were

incorporated into the design. The expansion circuit flowsheet is shown in Fig. 4.

Fig. 4 – Expansion Desorption Circuit Schematic

The upgraded components of the desorption circuit include:

Loaded Carbon Recovery Screen – The original was sized to recover a 5 tonne batch of carbon over

an 8.5 hour period. This duration required shortening, so the pump and screen were upgraded to

facilitate carbon batch recovery over 3 hours.

Acid Wash Column – Originally sized for 5 tonnes of carbon. The new design adopted operation of

the column more frequently in lieu of duplication. The discharge piping was modified for carbon

transfer to either of the parallel elution columns.



Elution Column – Originally sized for 5 tonnes of carbon. This column was duplicated to allow for the

increased operating frequency of the Zadra elution circuit. This also required duplication of

associated equipment such as eluate heater, electrowinning cell, heat exchangers, strip solution tank

and strip solution pump. The existing eluate heater, heat exchanger and strip solution pump have

also been upgraded to match the new equipment, whilst the existing electrowinning cells remain

unchanged.

The arrangement of one acid wash column connected to two parallel elution columns presented a potential

safety risk if the elution column was mistakenly assumed to be empty, but was actually in use and valves

were opened. This scenario could result in a serious incident if a person was in close proximity to the elution

column and it instantaneously depressurised and hot solution containing sodium cyanide and sodium

hydroxide vented to atmosphere. To reduce the likelihood of this type of incident, a number of safety

features have been incorporated into the design:

Two valves, one manual and one automatic, installed in the carbon transfer lines from the acid wash

column to the two elution columns. The automatic valve has been interlocked with the position of the

manual valve.

The manual valve supplied with a physical key, which has been interlocked with the strip solution

pump discharge valve of the corresponding column. Therefore, if the strip solution pump was

running, then it would not be possible for the elution column carbon feed valve to be opened

because if the elution step was in progress, then opening the elution column carbon feed valve

would instantaneously depressurise the vessel resulting in an uncontrolled release event.

The actuated valves in the elution circuit have been automated and sequenced.

Each actuated and manual valve in the elution circuit has been provided with position indication and

feedback to the PLC for position confirmation. If a fault occurs, the system will fail to a safe valve

position configuration and hold until the issue is rectified.

This upgraded desorption circuit has been fully commissioned. The two elution circuits differ in terms of the

electrowinning cell size and configuration. Table 4 lists the parameters of the two electrowinning circuits.



Table 4 – Electrowinning Circuits Parameters

Parameter Unit Original – EC1 Expansion – EC2

No. of Electrowinning Cells 3 in parallel 1

Type of Cathode Stainless Steel Stainless Steel

No. of Cathodes 3 x 12 22

Submerged Cathode Area Each m2 0.525 1.0

Submerged Cathode Area Total m2 18.9 22.0

Plating Area per Cathode Each m2 6.8 13.3

Plating Area Total m2 245 293

Rectifier Size Amps 3 x 1500 6000

The performance of the two elution circuits has been monitored and the cell pass efficiency during an elution

cycle is displayed below in Fig. 5. The single cell in the EC2 circuit has a greater pass efficiency than three

smaller cells fitted with fewer cathodes and operated in parallel. The new elution circuit (EC2) commonly

achieves lower barren carbon assays than the existing elution circuit (EC1). The barren carbon assays for

EC1 struggle to achieve 50-100 g/t consistently, whilst EC2 typically achieves <40 g/t with a shorter cycle

time. Typically, EC1 elution cycle time is 18 hrs and EC2 is 14 hrs. The improved performance of EC2 has

been related to a higher cell pass efficiency along with a higher current efficiency in comparison to EC1.

Limited space in the gold room has prevented replacement of the original existing electrowinning cells.

0

10

20

30

40

50

60

70

80

90

100

0 2 4 6 8 10 12 14 16 18

ELUTION CYCLE TIME (Hrs)

CE

LL

PA

SS

EF

FIC

IEN

CY

(%

)

EC 2 STRIP PROFILE EC 1 STRIP PROFILE

Poly. (EC 2 STRIP PROFILE ) Poly. (EC 1 STRIP PROFILE )

Fig. 5 – Expansion Desorption Circuit Flowsheet



CONCLUSIONS In 2009, Chirano Gold Mines Limited expanded the Chirano Gold Project from its original design capacity of

2 Mtpa ROM ore to a new design value of 3.5 Mtpa. This paper has outlined the design approach to three

main areas for cost effective solutions and safety considerations. The commissioning period included a

rapid ramp up to design and subsequent operation has tested the design selections and monitored

performance.

Overall the comminution circuit has been able to achieve design despite some issues regarding the client’s

crushing circuit. Whilst these are being resolved, modifications to the expanded design configuration of the

SAG mill and pebble crushing are being utilised.

For the thickening circuit, the expanded flowrate required an adjustable bypass system to be installed. This

solution sought to find a balance between capital expenditure and effective utilisation of existing equipment.

With further optimisation, the thickener performance has now exceeded design expectations due to an

improved flocculant reagent selection along with optimum operating conditions of elevated pH.

The expanded desorption circuit now consists of two parallel elution columns with a common acid wash

column. This equipment duplication warranted additional safeguards to be incorporated into the design to

reduce the likelihood of a potential safety risk when operating two elution columns with one common acid

wash column. The performance of the two elution circuits have been monitored and show the benefit of

higher cell pass efficiency along with higher current efficiency to reduce the elution cycle time for the newer

equipment.

This paper has outlined some novel design approaches applied at Chirano Gold Mine to achieve a cost

effective and safe design for a plant expansion.

ACKNOWLEDGEMENTS Permission from Kinross Gold Corporation and Lycopodium Minerals QLD Pty Ltd to publish this paper is

gratefully acknowledged.

REFERENCES Atasoy, Y, Valery, W and Sklalski, A, 2001. Primary Versus Secondary Crushing At St. Ives (WMC) Sag Mill Circuit, in

Proceedings International Autogenous and Semiautogenous Grinding Technology Conference 2001, (Minerals

Engineering International: Cornwall, U.K.).

Davies, E, 2007. Chirano Gold Mine Expansion Study Three Stage Crushing, in OMC Report No. 6541 to Chirano Gold

Mines Ltd. March 2007.

Koch, F and Putland, B, 2008. Chirano Gold Mines Limited Plant Expansion, in OMC Report No. 8194 to Redback

Mining Ltd. May 2008.

MacNevin, W, 1997. Kidston Gold Mines Case Study: Evolution of the Comminution Circuit, in Proceedings Crushing

and Grinding in the Mining Industry 1997, (IIR Conferences: Sydney).



McGhee, S, Mosher, J, Richardson, M, David, D and Morrison, R, 2001. SAG Feed Pre-Crushing at ASARCO’s Ray

Concentrator: Development, Implementation and Evaluation, in Proceedings International Autogenous and

Semiautogenous Grinding Technology Conference 2001, (Minerals Engineering International: Cornwall, U.K.).

Mwehonge, G L K, 2006. Crushing Practice Impact on SAG Milling: Addition of Secondary Crushing Circuit at Geita

Gold Mine, in Proceedings International Autogenous and Semiautogenous Grinding Technology Conference

2001, (Minerals Engineering International: Cornwall, U.K.).

Putland, B, Siddall, B and Gunstone, A, 2004. Taking Control of The Mill Feed: Case Study - Partial Secondary Crushing

Mt Rawdon, in Proceedings Met Plant 2004, (The Australian Institute of Mining and Metallurgy: Melbourne).

Siddall, B and Putland, B, 2007. Process Design and Implementation Techniques for Secondary Crushing to Increase

Milling Capacity, in SME Annual Meeting 2007, (The Society for Mining, Metallurgy and Exploration: Colorado).

SMCC, 2007. Simulations of the Chirano Grinding Circuit When Treating Secondary Crushed Feed, in SMCC Pty Ltd

Report to Chirano Gold Mines Ltd. July 2007.

Thong, S, Pass, D and Lam, M, 2006. Secondary Crushed Feed Before SAG Milling – An Operators Perspective Of

Operating Practices at Porgera and Granny Smith Gold Mines, in Proceedings International Autogenous and

Semiautogenous Grinding Technology Conference 2001, (Minerals Engineering International: Cornwall, U.K.).



Engineering and Science in Flotation Cell Design

J Euston1

1. Product Manager, Flotation – Australia SEA, FLSmidth, 5 Comserv Close, West Gosford NSW 2250. Email: [email protected]

ABSTRACT In the 1990s round flotation tanks were introduced as a result of the lower fabrication costs for larger

cells. It was soon observed that round tanks held a number of further advantages over the traditional

rectangular cells, such as better hydrodynamics, lower costs and reduced footprint per unit volume.

In the last twenty years flotation cell sizes have increased almost exponentially and designs of 250 to

300m3 are a well established standard with plans for full scale pilot installations of 500m3. The option

to install these large cells gives flexibility in circuit arrangements and allows efficient and economic

matching of flotation lines with modern mill designs.

The benefits of scale up are well known with reduced numbers of cells required, lower power costs

per unit volume, etc. These benefits will be briefly reviewed with specific reference to the installation

of 300m3 to 350m3 flotation cells at the Copperton Concentrator of Kennecott Copper in Utah, USA.

The sophisticated techniques required to model these ‘supercells’ have led to a critical review of the

fundamentals of flotation science and engineering. In addition to outlining the advantages of large

flotation cells, this paper will review the science which supports, and occasionally challenges, the

ongoing trend to larger flotation cells.

INTRODUCTION The first modern industrial scale use of froth flotation for separation in mineral processing is

commonly attributed to Broken Hill in Australia in the opening years of the 20th century. The History

of Flotation (Lynch et al., 2010) provides an excellent review of flotation history with an informative

second chapter on the BHP developments of Potter and Delprat at Broken Hill.

By the 1930s the mechanical cell market was dominated by three manufacturers, Fagergren Wemco,

Denver and Galigher-Agitair.

Self aspirated Wemco machines continue to be manufactured by FLSmidth along with the Dorr Oliver

forced air designs. Galigher was eventually absorbed into Baker Hughes (now FLSmidth) and Denver

is now owned by Metso. Chapter 4 of Lynch’s book provides an interesting history of flotation

machine manufacturing.



The original Fagergren machine had a horizontal rotor which was changed to a vertical rotor in 1934,

recognizable as an early Wemco machine. Fig. 1 compares an early Fagergren design alongside the

contemporary Wemco cell and mechanism.

Fig. 1. The Fagergren Wemco Predecessor in 1920 and the Current Wemco Design

Outokumpu are a relative newcomer to the manufacture of flotation machines and as such were able

to approach the concept with fresh eyes whilst still adopting current best practice. Initially producing

2.5m3 machines in 1959 for their operation in Finland, the company began an extensive development

programme in 1970 to build larger machines. By 1975 the new 16m3 machine was set to establish

itself as the new standard.

Until the 1980s the development of flotation was based on tanks which were rectangular and of the U

tank or hog trough design. Tank sizes increased to close to 100m3 with the Wemco 1+1® model 3000

at 85m3 being possibly the largest U tank design ever produced.

The introduction of round tanks was a major step change in mechanical cell design in the 1990s. The

first prototype round tank is reported to be the Outotec TankCell® built in 1982 in Finland with the

first operating units installed in the early 1990s. Round tank designs were developed initially to

reduce fabrication costs especially above the 30m3 size. It was quickly realised that round tanks also

possessed superior hydrodynamics and the size range was extended downwards below 30m3. Round

tanks brought a number of advantages namely :

No dead spots in the corners of tanks,

Smaller mechanisms and impellers per unit volume,

Smaller overall footprint per volume unit,

Improved residence time distribution and fewer cells per bank required,

Radial launders simpler and more efficient than paddles,



Lower cost per unit volume,

Lower power consumption per tonne of ore treated, and

Reduced maintenance.

Despite these clear advantages it appears to have been close to 10 years before round tank designs

became universally accepted. For tank sizes less than 20m3 rectangular designs may still be

considered today for low volume applications and to maintain minimum cell quantity requirements.

The development of lower grade ore bodies and the established economies of scale have led to larger

processing plants in recent years. The development of larger cells has been facilitated by advances in

both computer based and fundamental modeling techniques but also by improved process control

using froth cameras etc. and more sophisticated reagent regimes. Flotation cell installation lists have

orders for 300 flotation cells in the 1960s for a single roughing duty. Today, a bank of 10 cells could

satisfy the same duty.

The chart below summarises the growth in flotation cell sizes.

Fig. 2. Flotation Cell Sizes 1960 to 2005

Batterham and Moodie (2005) have made an interesting observation based on the actual productivity

of a flotation cell measured in tonnes per day per cell. As expected, this has increased from about 500

tonnes per day per cell to a current level of 2000 to 2500 tonnes per day per cell. They suggest

however, that this is leveling off. Comparing the author’s data with the chart above suggests that

tonnes per day per m3 of volume has decreased considerably from values around 25 in the 1970s and

1980s to a present value of closer to 10 tonnes per day per m3.

Currently, the largest standard sizes are 200m3 to 250m3 in volume although the new generation of

300m3 machines are quickly establishing themselves as the new standard for large volume

installations. The first of these were installed by Outotec at OceanaGold’s Macraes mine in New



Zealand (Coleman and Dixon, 2010). The author’s report that more than four years of development

were required prior to installation and provide a good summary of the subsequent on site evaluation of

the flotation cells using cell hydrodynamics, kinetic evaluations and metallurgical testing.

In a parallel development process FLSmidth (Dunn et al, 2009) installed the first of its SupercellsTM at

Kennecott’s Copperton concentrator in Utah, USA in 2009. This was an ambitious and extensive

project which included the installation of three pilot plants and a tank design which could

accommodate the three distinct mechanisms supplied by FLSmidth; the Dorr Oliver, Wemco and

Xcell mechanisms. FLSmidth is currently completing a major installation of 300m3 Wemco cells at

Esperanza in Chile as shown in Fig. 3, below.

Fig. 3. The Installation of 26 x 300m3 Flotation Cells at Esperanza in Chile

This paper will review the progress of this ambitious project in the context of evolving scale up and

design techniques. The complexity of the flotation process and the concomitant intricacy of their scale

up and design require an understanding of a range of disciplines from surface chemistry to

hydrodynamics. Prior to presenting the Kennecott work it might be appropriate to briefly review some

of the techniques employed and the developments in the discipline in recent years.

FLOTATION DESIGN GUIDELINES The design of flotation circuits has traditionally been based on some accepted design philosophies to

achieve the required metallurgical results complemented by experience of the installed base.

Increasingly complex computer aided modeling has become available in recent years and this,

coupled with fundamental research, has provided the design engineer with a large number of

sophisticated tools. It is now possible to design and build larger flotation machines with a high level

of confidence that they will meet both metallurgical and operational expectations. For those of us who

are not focused on these developments, the field is becoming increasingly complex. However, the



basic design premises remain essentially valid and these will be briefly reviewed to provide a context

for the scale up discussion.

It is useful to remind ourselves that the three fundamental functions of a flotation cell, those of

mixing, aeration and froth removal must always be considered. Nelson et al., (2002) have provided an

excellent discussion of flotation cell design. Any mechanical flotation machine must provide the

following functions:

Good air / pulp contact,

Adequate solids suspension,

Good mixing without stagnant zones or short circuiting,

A quiescent zone for froth separation,

Adequate froth removal, and

Adequate residence time to allow the desired recovery of the valuable constituent.

Capacity Required The calculation of the total volume required for flotation is simple enough being based on the

residence time required, the volumetric flow into the cells and the aeration factor.

Volumetric Flow x Residence Time x Aeration Factor

Residence Time The required residence or retention time is usually determined by test work from either bench scale or

pilot plant testwork. If laboratory scale test work has been performed the operational residence time

is determined using scale-up factors.

A flotation cell may be described as a reactor and as such it is more accurate to speak of a residence

time distribution within the cell. In a plug flow or batch reactor all particles have the same residence

time. In a flotation cell, however, short circuiting and varying turbulence within the cell can result in a

range of residence times for individual particles entering the cell. Residence time distributions can be

measured with the use of chemical or radioactive tracers. Fig. 4 shows a typical residence time

distribution for an operating flotation cell.



Fig. 4. The Concept of Residence Time Distribution

It is well known that flotation essentially occurs in and around the rotor. Using the Wemco design as

an example it is useful to consider the number of passes of the pulp through the flotation mechanism.

The (average) flotation cycle represents the number of passes – and opportunities for recovery – of a

particle through the flotation cell.

Fig. 5. Residence Time in a Flotation Cell

Volumetric Flow For design purposes the volumetric flow used is nominally the feed flow to the flotation bank. Of

course the volumetric flow into each consecutive cell is somewhat less than the previous cell so it may

be more accurate to determine total volume required on the average of feed and tailings flows? As

with all rules of thumb there is a tendency to conservatism and whilst this is admirable is it important

to know how many stages of the process are actually applying a safety factor. Of course, it is not

unheard of for a flotation plant to be faced with increased flows at some time after commissioning. It

is prudent then to consider the possibility of increased flows at the design stage, particularly for valve

sizing and inter tank connections.

Aeration Factor

Residence Time Vn/Qf

Average Circulation Time Vn/Qp

Average Flotation Cycle (Vn/Qf)/(Vn/Qp) = Qp/Qf

Where:

Qf = the flow into the cell

Vn = cell volume

Qp = flow through the rotor or impeller.



Froth flotation is a 3 phase system and the presence of air will reduce the effective slurry volume in

the flotation cell. The aeration factor is used to compensate for this air. Typically a value of 0.85 is

used but for Wemco cells which only aerate the top half of the cell a factor higher than 0.9 has been

suggested.

Secondary Factors In addition to mixing and aeration flotation requires that the froth be removed from the flotation cell.

The geometric efficiency which drove the move to round tanks and now even larger round tanks has a

detrimental effect on the surface area available for froth removal. Effective froth removal requires

both sufficient surface area at the froth / pulp interface and sufficient overflow length to maintain

froth movement. These factors are quantified as the froth carrying capacity and the lip loading. For

applications with low mass pull to concentrate, residence time may be the only design consideration.

However, in high mass pull applications where the concentrate flow is a high proportion of the feed,

such as coal or phosphate or in cleaning circuits, the ability to remove the froth from the cell becomes

critical. As cell sizes become larger the ratio of surface area to volume and lip length to volume

decrease. Carrying capacity can be increased by the use of external launders to increase the surface

area of the cell. In this way a 25% to 30% increases in froth surface area and hence froth carrying

capacity can be achieved. As a guide a froth carrying capacity of 2tph/m2 is often regarded as an

upper limit.

In addition the installation of radial launders effectively increases the length available for froth

overflow. Clearly the radial launders also decrease the surface area but it must be remembered that

although this affects the froth / air interface the pulp / froth interface surface area does not change. It

is usual to use 1.4tph per linear metre of overflow lip as an upper limit.

A Word on Probability Essentially froth flotation depends on the combination of a number of probabilities. The probability of a successful and sustained attachment is given by the product of these probabilities. Pc = probability of collision Pa = probability of adhesion Pd = probability of detachment Pl = probability of levitation Collision The probability of a particle colliding with a bubble depends on the relative sizes of particles and bubbles, hydrodynamics of the flotation environment but not on hydrophobicity. Probability increases as particle size increases. Adhesion The probability of a particle attaching to one or more bubbles is largely a function on particle hydrophobicity. The probability decreases as particle size increases. Detachment The probability of the particle/bubble aggregate persisting depends on hydrophobicity and the hydrodynamics of the system. Probability increases as particle size increases. Levitation The probability of the particle reaching the concentrate overflow is the final stage. If a stable particle/bubble aggregate forms the aggregate must pass through the pulp / froth interface to the overflow launder. This part of the process requires a quiescent zone but also sufficient energy for froth movement towards the discharge.



Cells in a Bank To prevent short circuiting, the total volume is divided into a number of individual cells. The

optimum number of cells is generally considered to be between 3 and 8, regardless of cell size. With

less than 3 cells short circuiting is likely to be an issue. More than 8 cells in a bank may only be

necessary to achieve a certain total volume as described above. In addition to the number of cells in a

bank the number of cells which can be accommodated on the same height must also be considered.

Table 1 below gives a guide to the number of cells in a control row along with the accepted drop

heights between rows to ensure adequate flow along the bank.

Table 1. Cells on a Level and Level Control Height

Cell Volume m3 5 10 20 50 100 200 250 300

Maximum Number of cells in a control stage

6 4 3 2 2 1 1 1

Drop height between banks mm 300 350 450 650 900 1000 1100 1200

Power Consumption Environmental and economic demands to reduce energy consumption, carbon footprints, etc. are

requiring all industries to review and demonstrate awareness of energy use across all aspects of their

operations. Mining is certainly an energy intensive industry. Govender et al., (2011) have reviewed

the energy usage in mining and have cited a US Department of Energy report which indicates that

separation in mineral processing (essentially flotation) accounts for around 4% of total mining costs.

Larger cells have been shown to reduce the overall power requirement with a drop in the kW per m3

of flotation volume. The chart below is an approximation of this.

Fig. 6. Power Consumption as a Function of Cell Size



Obviously, flotation cells should be operated as efficiently as possible. Hydrodynamic and

metallurgical testing of flotation cell prototypes includes the measurement of recovery as a function of

rotor speed to determine the optimum speed for recovery and energy conservation. The application of

variable speed drives on flotation cells is certainly worthy of consideration. Ultimately, though the

recovery of valuable material recovered is proportional to the energy input. Power draw in a flotation

cell is often measured based on the Power Number, a dimensionless number applied commonly to

stirred systems (see inset). The variables in the Power Number equation are fluid density, rotor speed

and rotor diameter. The fluid density factor has obvious process implications as regards feed solids

and must be corrected for aeration of the pulp. Rotor speed and rotor diameter are related to the power

of 3 and 5 respectively.

Govender et al., (2011) has estimated the effect of energy input on recovery as part of the Kennecott

copper trial with the Wemco 300m3 Supercell and have indicated that over the projected 20 year

lifespan of the flotation cell, imparting extra energy into the flotation cell at a cost of US$500,000 will

yield extra recovery of Copper and Molybdenum to the value of US$160,000,000. The authors

conclude that whilst an appreciation of energy reduction is important for both environmental and

economic reasons, it should be appreciated that the primary objective of mining, which is to recover

valuable mineral, is not compromised.

A Quick Word on Dimensionless Numbers Dimensionless numbers are widely used in science and engineering to compare a property independent of the size. Pi is a simple dimensionless number which is the ratio of a circle's circumference to its diameter. Obviously, Pi has the same value regardless of the size of the circle. Dimensionless numbers are used in the technique of dimensional analysis, or hydrodynamic analysis when the technique is applied to fluid systems. The use of dimensionless numbers allow for measurements in a smaller device, such as a pilot plant or prototype, to be applied to a larger model under development. Of course, care must be taken in the level of scale up. A number of dimensionless numbers are used in fluid dynamics and some of these can be effectively applied to the design of flotation cells, two common ones being the Froude number and Weber number which relate the inertia of a fluid to the force of gravity and the surface tension at the air pulp interface. The Froude number relates the size and speed of the impeller. These dimensionless numbers relate the forces that act on a fluid such as pressure, fluid viscosity, gravity, etc. to the inertia of the fluid or the resistance to motion. Froude Number (N2D)/g Weber Number (N2D3)p/y Power Number p.N3D5 Where: N = impeller speed D = impeller diameter G = acceleration due to gravity (force) p = specific gravity y = surface tension at the air pulp interface



Coarse and Fine Particles The classic plot of particle size versus recovery has been presented many times and this paper would

not be complete without its inclusion. Improving the recovery of coarse and fine particles are the twin

holy grails of flotation. Essentially for ultrafine particles, below around 10 microns, the particles have

insufficient inertia to produce effective collisions with the viscous forces dominating. Coarser

particles are too heavy and too large to maintain effective collisions in the turbulent environment of

the flotation cell. High energy input is required to effect particle bubble attachments, especially for the

finer particles whereas low shear is necessary to successfully recover coarser particles into the

concentrate.

Fig. 7. Recovery and Particle Size

Scale up is primarily concerned with increasing the size of individual flotation cells to achieve

residence time for a nominated volumetric flow. It is important that these new designs not only

maintain the grade and recovery performance of their predecessors but also achieve recovery over at

least the same size range. The increased recovery of slow floating particles, including the extremes of

the size distribution, should not be forgotten in designing larger flotation cells.

The section above has shown how energy input is directly proportional to overall recovery and

Govender et al., (2011) have suggested that any analysis of recovery should include an understanding

of recovery as a function of particle size. The authors conclude that energy input is critical to

achieving good recovery of the ultrafine particles.

Another approach to the use of high energy for fine particle recovery is to increase the shear in the

froth contact zone. This has been proposed with the Ratemax system developed by Pyramid and

Eimco which uses a flotation column sparger with an exit velocity of Mach 1.7. Jameson (2009)

describes the concept as part of his Concorde development.



The need to provide process conditions suited to coarser particles has been addressed by Eriez with

the Hydrofloat Separator described as an air enhanced gravity separator. It could alternatively be

described as a flotation cell with a fluidized bed replacing the water of the traditional flotation cell.

Another interesting advancement is the Proflote concept (Bott and Lumsden, 2009). This uses

permanent magnets to effect selective paramagnetic flocculation of ultrafine particles, producing

agglomerates of sufficient size to readily attach to bubbles. Although not suitable for all minerals,

where it is feasible, impressive increases in ultrafine recovery can be achieved. This patented

technology is based on conditioning the flotation slurry in a strong magnetic field where the mineral is

magnetised and this magnetised mineral can subsequently aggregate. These aggregated fine minerals

then float more readily and their recovery increases. Reported results have shown increases in metal

recoveries of up to 4%.

The testwork reported by Engelhard et al., (2005) showed that the increases in the <38µm mineral can

be substantial.

In a most recent trial, 20 Proflote units have been installed at a major copper / zinc concentrator. The

suppliers of the patented Proflote technology carried out a detailed statistical approach and are

indicating at 99% significance an increase in Zinc recovery of 1.3% with a decrease in Zinc in the

tails of 8%. This represents an increased income of $40,000 per day for a rental cost of $1,200 per

day.

Fundamental Research The following section will describe the tools available to flotation cell designers, essentially the

engineering tools of hydrodynamics and, more recently, computational fluid dynamics. Metallurgical

testing to confirm the efficacy of these design tools is also described. A further and increasingly

important contribution to design and scale up is being provided through fundamental research.

The ubiquitous AMIRA P9 project has been ongoing for over 45 years and continues to make a

valuable contribution to flotation knowledge, due both to its practical hands on approach and its

ability to adapt to the changing focus of its industry sponsors. The generic title of “Optimisation of

Mineral Processing Through Modeling and Simulation” summarises the overall goal of the project

which is to essentially simulate mineral processing unit operations, comminution through to flotation

based on ore properties. In its current incarnation, P9O is also investigating some more fundamental

aspects of froth flotation related to the froth phase such as bubble size and size distribution, movement

in the froth and across the pulp / froth and froth / air interfaces.



SCALING UP SCALE UP In 2005 Weber et al., (2005) reported on the development of the world’s largest flotation machine of

257m3. The author notes that in 1971 a similarly styled paper extolled the virtues of a new generation

of large 8.5m3 and 14.2m3 flotation cells. A mere five years later the major flotation cell

manufacturers (Coleman et al., 2010; Dunn et al., 2009) are once again promoting the world’s largest

flotation machines with capacities from 300m3 to 350m3. Almost before these developments can be

presented to the industry, word is spreading of 500m3 machines being developed. Fig. 8, below is an

extension of Fig. 1 and shows the planned development of cell sizes towards 500m3.

2002 – Scale Up from 160m3 to 257m3. The development of the 257m3 flotation cell represented a scale up of around 60% over the then

‘standard maximum’ Wemco machine of 160m3.

The Wemco flotation machine is characterized by having an elevated rotor which performs both

mixing and air ingestion functions and as such it is the key design consideration for scale up purposes.

Fig. 8. Flotation Cell Size Projections

The scale up process involved:

Stage 1 The use of proven scale up methodologies.

Stage 2 The application of hydrodynamic criteria.

Stage 3 The use of hydrodynamic testing to verify and fine tune the design parameters.

Stage 4 Metallurgical testwork to prove the design.

Stages 1 and 2 are carried out during detailed design whilst stages 3 and 4 are performed post

installation.



The basic premise of Stage 1 is to maintain dimensional consistency and the use of dimensionless

numbers is a key part of the process. Using these techniques and the known requirements to maintain

flow through the draft tube and the required air entrainment, the rotor size and speed can be

determined.

The hydrodynamics of the design are defined by the relationship between mixing and air entrainment

and the critical rotor parameters of submergence, rotor speed and rotor engagement.

The table below summarises the hydrodynamic scale up criteria used (Weber et al 2005).

Table 2. Hydrodynamic Scale Up Criteria Example

Criteria (water only) Definition Machine Size m3

130 160 257

Superficial Gas Velocity Qa/Ac cm/sec 2.10 1.65 2.10

Specific Cell Power P/Vc kW/m3 0.92 0.89 1.07

Liquid Circulation Vl.Ad m3/min 95.4 108 197

Air Capacity Qa/N.D3 0.15 0.14 0.10

Power Number Pg/p.N3.D5 5.7 6.0 6.9

Note 130m3 and 160m3 are measured values, 257m3 are predicted.

Where: Ac Cell surface area Ad Draft tube cross sectional area D Rotor diameter N Rotor rotational speed P Absorbed power p Liquid specific gravity Qa Air ingestion rate Vn Cell volume Vl Liquid viscosity

Based on this analysis a 1.40m diameter rotor was determined to be required operating at 90 to 110

rpm. In 2002 Dorr Oliver Eimco (now FLSmidth) completed the detailed design with the inclusion of

vertical baffles and radial launders to decrease froth travel distance and provide increased quiescence

in the upper portion of the cell. The first cell was installed at the Minera Los Pelambras mine in Chile.

Following installation the second part of the process was to evaluate and if necessary fine tune the

design. Engagement, rotor speed and submergence were tested above and below the expected

operating range and a series of models developed. These models related:

Superifical gas velocity to submergence and rotor speed,

Cell power to engagement, rotor speed and submergence (and multiple factors), and



Liquid circulation to engagement, submergence and rotor speed.

Further optimisation of the hydrodynamic characteristics was then used to maximise performance and

minimize power consumption. In general predicted and actual design criteria matched well, the main

discrepancy was an 18% lower power consumption with the actual value being 0.88 kW/m3 (water

only) compared with a predicted 1.07kW/m3, a result which the authors seemed happy to accept.

Finally, residence time distribution tests were carried to confirm the mixing regime in the cell. Also,

in order to confirm adequate solids suspension, slurry samples were collected at 1 metre increments

down the cell and analysed for solids size distribution.

2007 – Scale Up to 300m3. Compared with the previous discussion the scale up from 257m3 to 300m3 is relatively modest at

17%. Prior to engineering development FLSmidth worked closely with the Center for Advanced

Separation Technologies (CAST) in the USA to fundamentally review the mechanics of flotation and

to develop Computational Fluid Dynamics models specific to the company’s flotation equipment.

CFD is used in many fields of engineering and its application to the development of mechanical

models for flotation is now a widely accepted tool. Initially used as a tool for scale-up, CFD. has

become highly sophisticated and is now applied to many aspects of flotation cell research and

development such as the provision of hydrodynamic data (velocity, and turbulence dissipation)

needed for mathematical modelling of flotation dynamics, the investigation of flow patterns and their

impact on flotation processes and the modelling of forces on impellers and stators for stress and

vibration analysis.

At the beginning of this project FLSmidth was able to call on hydrodynamic data and CFD modeling

of over 100 existing 257m3 SmartCellsTM.

In addition to the use of CFD modeling during many stages of the design process, the design of the

final 300m3 cell design used the previously tried and tested techniques described above;

hydrodynamic evaluations of and comparisons with existing cells, experience with operating the same

design and metallurgical testing post installation



Fig. 9. CFD Models of the Wemco Rotor

FLSmidth installed the first of their 300m3 units at the Copperton concentrator of Kennecott Copper.

The proximity of the concentrator to their Technology Center in Salt Lake City, Utah, allowed ready

access for the design team. As part of the project, FLSmidth installed three pilot plants adjacent to the

full scale cell so that each of the three mechanism types supplied by the company (Wemco, Dorr

Oliver and Xcell) could be tested simultaneously using the same feed as the full scale unit. Also and

in line with their current philosophy the tank was capable of accepting all three of the mechanism

types. The ability to use all three mechanisms with only slight internal tank changes resulted in the

rated capacity of the cell ranging from 300m3 to 350m3.

Fig. 10. 300m3 Installation at Kennecott Copper Showing the Three Pilot Plants

Tables 3 and 4 summarise the parameters and measured responses for the hydrodynamic and

metallurgical evaluations of the Kennecott installation (Lelinski et al., 2009).



Table 3. Hydrodynamic Evaluation of the 300m3 Cell

Rotor Parameter (Wemco)

Measured Response

Submergence Machine Power Rotor Speed Aeration Rate

Rotor Engagement Pulp Circulation

Table 4. Metallurgical Evaluation of the 300m3 Cell

Rotor Parameter (Wemco)

Measured Response

Froth Depth Feed Assay

Concentrate Assay

Rotor Speed Tailings Assay Aeration Rate

Air Rate Absorbed Power Solids Content

Typically, the installation of a new design of flotation cell would be followed by many months of

hydrodynamic and metallurgical testing with some fine tuning prior to handing the equipment over to

the site personnel. Lelinski et al report that in the case of the Kennecott installation recoveries and

grades were higher than predicted from the start.

Fig. 11. 300m3 Examples of the Kennecott Hydrodynamic Evaluation (Lelinski et al., 2009)



Fig. 12. The First Wemco Metallurgical Campaign. The Vertical And Horizontal Lines indicate

the Performance Guarantee (Lelinski et al., 2009)

2011 – Scale Up from 300m3 to 500m3. The increased confidence with modern techniques has resulted in the plans for the next size of

flotation cell almost before the 300m3 cells have had their first full scale installations commissioned.

The techniques described above are of course still relevant but just as CFD techniques enhanced the

design of the 300m3 the technique of Population Balance Modeling (PBM) is being increasingly

applied to the future of flotation cell design and scale up.

The use of the technique of Population Balance Modeling brings together these techniques and is

shown diagrammatically in Fig. 13, below. A Population Balance Model is a mathematical description

of a process developed from a number balance in a discrete volume. In this regard it has parallels with

CFD techniques. In its simplest form a population balance relates the input into a fixed volume to the

output of a particular property with consideration of the quantity of that particular property being

created, accumulated or destroyed.

As with CFD the basic premise of Population Balance Modeling is relatively simple, although the

mathematics soon becomes a little too complicated for the layman.

The classical approach to scale up as described above uses a largely kinetic approach using first order

rate constants, residence time distribution studies and dispersed mixing for single cells and banks of

cells as reactors in series. This approach does not account for possible different residence times of

particles and bubbles in the cell and how the residence time distribution relates to different parts of the

cell. As mentioned earlier the importance of the froth and the froth / pulp interface is receiving

renewed attention and the kinetics approach does not account for bubble loading, drop back and

entrainment / true flotation. PBM provides a framework for integrating the various design and scale

up tools.

Cu Concentrate Grade

Cu

Rec

ove

ryColumn Cells

WEMCO SuperCell



Fig. 13. The Particle Balance Model Concept

CONCLUSIONS Increasingly lower grade ore bodies and the favourable economics of larger plant throughputs have

demanded that flotation cell manufacturers develop larger and larger flotation cells to satisfy the

market demand. Increasingly sophisticated techniques employed such as Computational Fluid

Dynamics and Population Balance Modeling have complemented the established scale-up and design

criteria such as hydrodynamic modeling and classical design guidelines. The success of recent large

cell installations such as the 300m3 cells at Macraes and Kennecott are a testament to the confidence

in these new techniques and the major manufacturers are already making plans for even larger

flotation cells. The benefits of these larger cells such as lower overall power consumption and reduced

footprint with fewer larger cells are clear.

What is also clear is that flotation cells sizes will continue to increase. Plans are already underway for

the installation of 500m3 units and it is likely that with the tools available to the designers such as

CFD and PBM there are unlikely to be too many surprises in the installed performance.

What the limit is to flotation cell size is difficult to predict. The authors of some of the references in

the literature over the year’s extolling the ‘world’s largest flotation cell’ are testament to the ongoing

push to larger volume conventional cells. Will the contemporary presenters of the 300m3 and shortly

the 500m3 flotation cells be regarded in the future as being quaint or will a maximum eventually be

reached as a result of hydrodynamic or engineering limitations? Current scientific research and

analysis of the flotation process is as sophisticated as the techniques used for engineering design and

is moving at an equally rapid pace. It may be that a new technology will see the traditional tank cells

be replaced with a totally new method for flotation.



ACKNOWLEDGEMENTS The author gratefully acknowledges FLSmidth for permission to publish this paper. In addition,

members of the FLSmidth Global Flotation community for their private contributions. Obviously, any

errors, opinions or omissions are solely the responsibility of the author.

REFERENCES Batterham, R and Moodie, J.P, Flotation in the Minerals Industry – Some Observations, 2005 paper

presented to the Centenary of Flotation Symposium, Brisbane QLD, 6-9 June 2005.

Bott, A and Lumsden, B. Magnetic Conditioning at the Hellyer Tailings Retreatment Plant, 2009. In

Proceedings of theTenth Mill Operator’s Conference, Adelaide, South Australia, 12-14 October 2009,

pp 247-253.

Coleman, R and Dixon, A, 2010. Tried, Tested and Proven – 300m3 Flotation Cells in Operation, in

Proceedings of the XXV International Mineral Processing Conference (IMPC) 2010, pp 3429-3440

Dunn, M, Newman, R, Gordon, I, Lelinski, D, Weber, A, Dabrowski, B and Traczyk, F, 2009.

Commissioning of the SupercellsTM – World’s Largest Flotation Machines, paper presented to

Flotatation ’09, Capetown, 2009.

Engelhardt, D, Ellis, K and Lumsden, B, 2005. Improving fine sulphide mineral recovery – Plant

evaluation of a new technology, in Proceedings Centenary of Flotation Symposium, pp 829-834 (The

Australasian Institute of Mining and Metallurgy: Melbourne)

Govender, D, Dabrowski, B, Lelinski, D, Schreiber, H, Traczyk, F, 2011. Improved Fine Particle

Recovery, paper presented at Flotation ’09, Cape Town, South Africa, Nov 11-12 2009.

Jameson, G. 2009. New directions in Flotation Machine Design, paper presented at Flotation ’09,

Cape Town, South Africa, Nov 11-12 2009.

Lane, G, Brindley, S, Green, S and McLeod, D. 2005. Design and Engineering of Flotation Circuits

(in Australia) paper presented to the Centenary of Flotation Symposium, Brisbane QLD, 6-9 June

2005.

Lelinski, D, Redden, L, D, and Nelson, M,G. 2005. Important Considerations in the Design of

Mechanical Flotation Machines, paper presented to the Centenary of Flotation Symposium, Brisbane

QLD, 6-9 June 2005.

Lynch A J Harbort G J and Nelson M G History of Flotation AusIMM 2010. Published by the

Australian Institute of Mining and Metallurgy Spectrum Series 18.



Nelson, M, G, Traczyk, F.P, Lelinski, D, 2002. Design of Mechanical Flotation Machines, in Mineral

Processing Plant Design, Practice and Control (eds: Mular, A, L, Halbe, D, N, and Barrett, D, J), pp

1179-1203 (SME: Littleton).

Weber, A, Meadows, D, Villanueva, R, Palomo, R and Prado, S, 2005. Development of the World’s

Largest Flotation Machine, paper presented to the Centenary of Flotation Symposium, Brisbane QLD,

6-9 June 2005.



The Wonawinta Silver Project – Value Engineering at Work

P G Greenhill1

1. MAusIMM, Managing Director, RMDSTEM Limited, 4/448 St Kilda Road, Melbourne Vic 3004. Email: [email protected]

ABSTRACT The Wonawinta silver project near Cobar in NSW is under development by Cobar Consolidated Resources Limited (CCR). RMDSTEM Limited is part of the owner’s team for feasibility, design, engineering, construction and commissioning. With a 51Moz resource CCR is one of the largest pure silver players in Australia. A probable reserve of 14.3Moz has been defined for a mine life of approximately five years. The silver mineralization is contained in host rock composed of clays and oxidized limestone. The ore also contains on average 1.3% lead as cerussite, and other minerals. A simple circuit consisting of slurrying of the ore, gravity treatment of the coarse fraction to recover lead, silver extraction using CIL, elution and the Merrill Crowe silver recovery process, has been developed. The “value engineering” approach to this development pioneered over the past decade by RMDSTEM has meant that RMDSTEM has delivered the bankable feasibility study at a cost of $0.07 per resource ounce. It is highly cost effective making efficient use of shareholders funds as well as being very time efficient. This paper describes the application of the value engineering approach to the development of the Wonawinta project.

INTRODUCTION The development of mining projects is an expensive and complex activity. Justifying the investment of stakeholders by proving the viability of a project is critical. A traditional feasibility study takes a route from concept, through scoping, prefeasibility and finally bankable feasibility and usually involves the study of each step and component in a consistent manner. While detailed and well rehearsed, it is not a guarantee of a successful outcome. There is an alternative approach. This paper describes the application of this alternative approach to the Wonawinta Project for Cobar Consolidated Resources. It uses a combination of the Owners Representative Team, Quantitative Risk Analysis, Value Driver identification and techno-financial modelling. It is called Value Engineering to identify the application of company technology but it broadly falls into a methodology called Value Based Project Management.



Complementary to current project management practices, Value Engineering (Fig. 1) aims to maximise value and increase the chances of success by:

a. Identifying and quantifying project value drivers, their delivery chain and impact on project value,

b. Identifying major risks and uncertainties and quantifying their impact on project value,

c. Identifying and quantifying major management options available at each stage of development, their linkages and their value added, and

d. Identifying the optimal implementation pathways through proper decision analysis.

Fig. 1: Value, Risk, Maturity Profile

Used astutely this approach can deliver a feasibility study in a highly cost efficient manner. Importantly it can also reduce the time to deliver as the number of factors that need to be examined is usually less than in the conventional approach. In fact this is one of the key objectives of the method. It has only recently been applied to the modelling of cost reduction opportunities in mineral operations (Carter et al., 2009).

THE HISTORY The Wonawinta Silver Project is approximately 85km south of Cobar and 55km west of Nymagee in central New South Wales. Mineralisation in the area was first identified by Geopeko/Norgold Ltd (Peko) regional stream sediment sampling in 1989. Anomalous lead and zinc values were identified over 20km of strike length of Devonian Winduck Group sediments along the Wonawinta Anticline. Follow-up rock sampling and rotary airblast drilling by Peko outlined a 4.5km x 1km area with anomalous geochemistry up to 14% Zn, 6% Pb and 230 ppm Ag in weathered dolomitic



Booth Limestone overlying Thule Granite basement. Follow-up drilling by Peko, CRA Exploration and several other companies led to the conclusion that the primary Pb-Zn mineralization was not economic. Pasminco first recognised the potential for a shallow, continuous, high grade Ag “weathering blanket” along the Wonawinta anticline. In 2007 CCR was attracted to the Wonawinta area by its shallow, high-grade Pb-Zn-Ag intersections, and its proximity to CCR’s Gundaroo prospect where significant Pb-Zn mineralisation and a shallow Ag resource, the De Nardi deposit, had recently been discovered in a similar geological environment. Whereas previous explorers had been focussed on sulphide deposits, CCR considered the economic possibility of a standalone oxide silver deposit. CCR’s work program following JV formation in December 2007 included three phases of drilling, a mid-stream study that indicated an NPV of $60M for a 1Mtpa open pit operation, and feasibility study (RMDSTEM, 2010) resource estimation that resulted in the definition of 51Moz of inferred and indicated resources.

THE PROPOSED PLANT The Wonawinta resource consists of two main ore types: oxidised clays and oxidised limestone ores. Oxidised Clays represent approximately 60% of the resource, whereas the oxidised Limestone represents approximately 30% of the resource. The Wonawinta processing plant will treat blends containing up to 60% oxide limestone on average. The mining department will schedule ore deliveries and blend ore types to prevent plant throughput impacts from items such as over size rock at the log washer. The processing plant will produce silver bullion and a Lead concentrate. The Silver bullion will be produced utilising the Carbon-In-Leach, Elution and Merrill Crowe processes. The Lead concentrate will be produced by a gravity circuit. The process plant flow sheet has been designed and based on a throughput of 1 million tonnes per annum. The ability to change the mining schedule so as to mine the softer material first, made it possible to adopt a “staged capital expenditure” approach. As the softer ore would not require any comminution, capital expenditure associated with the comminution circuit could thus be deferred until Year 3 of operations. The processing plant throughput will be 800,000 tonnes per year in the first three years, utilising a log washer to ‘slurry’ the ore. Any material not broken down in this wet-scrubbing process will be stockpiled and re-processed when the comminution circuit is installed. The processing plant will be expanded to treat 1 Mtpa by the fourth year of operation with the installation of the comminution circuit consisting of a vertical shaft impact crusher, a rod mill, and associated ancillary equipment.



A schematic of the Wonawinta process flow sheet is shown in Fig. 2.

addition of water to form a slurry

ClassifyingCyclone

slurry passes through a pre‐leach feed tank

leach tanks where sodiumcyanide solution is added

carbon adsorption tanks

tailings storage facility

Carbon elution circuit

Merrill Crowe circuitPrecipitate dried,

retorted and smelted

ROM

molten material poured into a cascade of moulds

Silver exported

Gravity concentrator

Pb / AgConcentrate

Coarse Screen

Log washerStockpile for

future processing

Fine Screen

To Smelter

Stockpile for future processing

Fig. 2: The Wonawinta Process Flowsheet

HOW DID WE GET TO THIS POINT? The starting point was a quantitative risk analysis, or QRA, of all the elements contributing to the project’s capital and operating cost. The QRA showed the range of uncertainty and the level of impact of each factor on the project economics, and clearly identified some factors that had a greater impact than others. The focus was first on those factors with the greatest level of impact. The principle was to determine sooner rather than later if there were any showstoppers, likely to fast-fail the project.

How Does It Work? The basic premise is that understanding the factors that impact the value of any project will allow good, logical decisions to be made. More importantly, if an estimate of the quantitative impact can be made then there can be a prioritisation of the factors allowing the most important (highest impact) to be focussed on.

Tools The tools used are numerous but key to it are:



The Value Driver Analysis, Qualitative Risk Assessment, and Quantitative Risk Assessment

The Value Driver Analysis This tool, originally developed in the DuPont de Nemours company in 1914, is a set of financial ratios and key figures relating to the Return on Investment (ROI). This decision tree is used to find the drivers of a specific outcome, usually some sort of measure. For instance, the top level drivers of profits are revenues and costs. The purpose of the tree is to help determine which drivers are controllable and which drivers have the most impact on the outcome you want. The study team applies its own experience and also engages other experts if required to capture every factor that influences the value of the project. If that value is the Net Present Value then the contributing factors are

Revenue, CAPEX, OPEX, and Timing

Each of these is then broken down into its contributing factors which are in turn broken down into their contributing factors. Experience and judgement is used to determine how far the factors in the tree need to be broken down. The result is a diagram such as Fig. 3 which would be typical for a mining operation. The Recovery factor has been expanded to illustrate the lower level factors. Each other factor is similarly treated in the full process. Note that the diagram progresses from financial factors on the left to technical factors on the right. In this way a project model can be constructed based on technical performance criteria and costs.



NPV

CAPEX

OPEX

REVENUE

TONNES

HEAD GRADE

RECOVERY

IMPURITIES

TC/RC

GRAVITY REC

Metal Price

TIMING

LEACH

RECOVERY

FEED

DENSITY

FEED PSD

TEMPERATU

RE

RESIDENCE

TIME

CN CONC.

Fig. 3: Typical Value Driver Tree

Qualitative Risk Assessment The AS/NZS ISO 31000:2009 standard for Risk Management is employed. This was used as required to document and review risks identified for the project. Critically, it enables identification of risks that if not dealt with have either a potential catastrophic impact or a high likelihood of occurrence. Each of these risks must be mitigated via action to reduce or manage their impact. Typically anchored scales are used to assess the risks and place them in the risk assessment matrix. Quantitative Risk Assessment Sets of risks recorded and analysed in qualitative risk management, can be a foundation of quantitative risk analysis. Where risks can be quantified by the use of suitable measures and ranges, information can be used to model the process. Ranges can be set represented by the uncertainty in the risk. Quantitative risk analysis

may determine the likelihood that a project will be completed on time and within a budget,

identifies the critical project parameters which affect project schedule, possibly determines project success rate, and enables decisions to be made about viable project alternatives.

These three tools interact as shown schematically in Fig. 4. Both the Value Drivers and the Qualitative Risks contribute to the Quantitative Risk Assessment by identifying factors that



are incorporated in the QRA techno-financial model. The model generates a suitable indicator of the project value and identifies the key drivers of that value as well as the factors to which the project are most sensitive.

Fig. 4: The Tools

The Owner’s Team The Owner’s Representative model provides the environment in which the other tools are utilised. A team consisting of senior client representatives, project manager and senior study representatives coordinate and review all elements of the study. The project manager and their team engage and oversee other specialists as required and for essential purposes of the study. In this way, front end engineering is done at the time it adds most value and only when it is confirmed to be the best approach.

IMPLEMENTATION During 2007, in parallel with resource definition drilling, there was a preliminary assessment of mining and processing methods and costs. These were very much in the style of a scoping study. At the same time, non-technical issues were also reviewed and assessed. Issues such as permitting requirements, licensing (water, access), environmental impacts, infrastructure availability and requirements, and native title were examined in relation to factors such as the location, land use, and population. It was concluded at the time that none of these issues presented a high risk to the project and so were most appropriately dealt with as required as



the project was developed. All of these issues were re-examined in subsequent risk assessments and actions initiated as required. There were a number of key findings that informed subsequent work.

The mining study found that, being relatively shallow and clay hosted, ore could be mined by open pit and would not require drilling and blasting.

The metallurgical test work indicated the ore could be processed with a simple scrubbing circuit followed by cyanide leaching and elution – very much in the style of oxide gold production, which came to prominence in Western Australia in the 1980s.

At around 97 grams or 3 ounces per tonne, the ore is not considered high grade. This indicated a need for low cost mining and processing routes to keep the cost of silver production low. The metallurgical test work pointed to a route without crushing and grinding that would eliminate these steps and their associated capital and operating costs from the process. This encouraged the view that a relatively low cost process route could be designed. Progressively further work was done gathering and estimating mining and metallurgical data and costs culminating in a revision of the earlier NPV models. In late 2008 (Cobar, 2008a) an updated techno-financial model based on a set of revised assumptions calculated an NPV of $48M. The key assumptions are in Table 1.

Assumption Value

Total Ore Resource 6.5M tonnes

Head Grade: Ag - 97 g/t

Pb - 1.3%

Annual Processed Ore: 1M tonnes

Silver Production 3.3 Moz pa

Estimated Capital Cost $37 M

Silver Recovery 94%

Lead Recovery 55%

LOM cash cost of production (net of Pb credits)

$8.30/oz

Silver Price AUD$16.9oz

Pb Price AUD$1800

Table 1: Key Assumptions mid 2008

The key drivers of NPV and uncertainty are shown in Fig. 5 and Fig. 6 respectively. This sensitivity analysis set the direction for the next stage of work. Clearly increasing the ore reserve and developing a suitable mining method were worth focussing on, particularly as the ore reserve contributes most to the uncertainty of the NPV. Metallurgical factors such as silver and lead recovery are important to the project value but also important are the form of the products. Selling silver in concentrate rather than as bullion has major impact on the NPV



and also on its uncertainty. This directs an effort in developing a cost effective flow sheet that allows production of bullion. Similarly, improving the net smelter return for lead (by improving the grade) adds substantially to NPV while reducing its uncertainty.

Top Drivers of the NPV

12%

12%

12%

-25%

16%

16%

16%

16%

9%

21%

-12%

-12%

-12%

5%

-16%

-16%

-16%

-16%

-28%

-28%

-40% -30% -20% -10% 0% 10% 20% 30% 40%

Pb Recovery

Net Smelter Return Pb

Pb Price

Evaluation Period

Ag Recovery

HG - Ag

Ag Value Discount

Ag Price

Annual Ore Mined

Total Ore Reserve

Dri

vers

of

NP

V

Impact on NPV

10%

-10%

Fig. 5: NPV Drivers

Drivers of Uncertainty of NPV

-76%

-49%

-49%

-37%

-37%

-32%

20%

19%

-19%

-19% 19%

19%

-19%

-20%

23%

37%

37%

49%

49%

69%

-100% -50% 0% 50% 100%

HG - Zn

Net Smelter Return Zn

OPEX Mining Cost

OPEX -Processing

Pb Recovery


HG - Pb

HG - Ag

Ag Value Discount

Total Ore Reserve

Infl

uen

cin

g f

acto

rs

Impact on NPV

Min

Max

Fig. 6: Uncertainty Drivers

Following this a review of mining costs based on a change from owner operator to contract mining generated reduced estimates of capital and operating (cash) costs of $27M and $6.25/oz. Revised silver prices of A$14.65/oz and lead price of A$1869/t were used in the revised model generating an NPV of $60M. (Cobar, 2008b) Through 2008 and 2009, drilling continued to contribute to the estimation of the resource. During 2008, additional metallurgical test work resulted in the following conclusions:

• heavy liquid separation and characterisation indicated the potential of producing a gravity concentrate for lead and residual silver giving

o recoveries of 58-77% silver, and 50-73% lead, and o grades of 1200-2400g/t silver, and 24-53% lead.



• Cyanide leaching for silver gave 79-97% silver recoveries, little grinding was required and crucially the process showed low cyanide and lime consumptions.

• Highest leaching recoveries were demonstrated for the clay ore which also did not require grinding.

The deposit was demonstrating a simple metallurgy that met the requirements of a low cost process route whereby

• Crushing or grinding may not be required. – Washing and scrubbing may be enough for ore processing,

• Silver is amenable to cyanide extraction, • Good silver recovery through to bullion, and • Gravity concentrate provides saleable lead product with silver credits.

A conceptual flow sheet (Fig. 7) was prepared at this stage.

TrommelScreen

CIP

Desliming

Gravity Separation

Tailings Storage

Ag MC

Pb / Ag Conc

fines

100 tph

33 - 65 tph2-15 tph

Fig. 7: Conceptual Plant

The concepts in this flow sheet have been refined but are still the basis of the flow sheet of the current plant design. On the basis of this information, the selection of more specific processes and equipment proceeded. This allowed initial design criteria and improved cost estimates to be prepared. Much of the conceptual plant design was conventional in nature. Gravity circuit was assumed initially to handle a relatively high proportion of the feed stream through jigs, and CIP/CIL was highly conventional. The recovery of silver was also conventional with carbon adsorption and elution followed by either electrowinning or Merrill Crowe. Capital and operating cost estimates of an electrowinning circuit were very high. The alternate elution/MC process was reasonable and in accord with the need for a relatively low cost capital base for the operation. The risk of such a plant was further reduced by seeking out vendor sourced packages with performance guarantees. By July 2009 additional work had been completed that added resource ounces, developed mining schedules, and following additional metallurgical test work to refine the flow sheet, estimates of capital and operating costs were obtained and the process commenced to obtain project approval. This information (Table 2) was updated in the techno-financial model and rerun.



The model indicated a number of key changes. The project was no longer sensitive to the ore reserve and the mining rate had a lessened impact. Importantly now the silver head grade and recovery dominated and give a focus for ongoing work.

Wonawinta Silver Project

Ore throughput Yrs 1 – 3: 0.8Mtpa Yrs 4 – 6: 1Mtpa

Silver production 2.4Moz pa

Initial mine life 5.2 years

Silver recovery (CIL & gravity) 93.2%

Lead recovery 55%

Estimated capital cost (including Mining Pre-strip) $33M (Initial) $5M (Expansion)

LOM cash cost of production (net of Pb credit) $7.15/oz

Table 2: July 2009 Project Assumptions

Static Sensitivity of NPV

-22.2%

-21.3%

-14.3%

7.5%

6.5%

6.3%

5.9%

-5.1%

-5.1%

-5.1%

2.2%

-5.9%

-5.9%

5.1%

5.1%

5.1%

-2.2%

-7.5%

-6.5%

22.2%

21.3%

14.3%

-30% -20% -10% 0% 10% 20% 30%

CAPEX-CIL

Pb Recovery-Gravity seperator

HG - Pb


OPEX -Processing

Discount Rate

Average Striping Ratio

OPEX - Mining

Annual Ore Mined

Ag Recovery -CIL whole ore

HG - Ag

Infl

ue

nc

ing

fa

cto

rs

Change in NPV

-10%

10%

Fig. 8: Sensitivity of Project – July 2009

At this time a risk assessment was done. This was driven by the project entering a key phase of commitment. The objective of the future work was to deliver a feasibility study that was bankable. This requires attention to all elements of risk, uncertainty and project value. The risk assessment process allows all factors to be considered and all knowledge developed to date to be used to document critical risks to the project. The outcome is shown in Table 3.



Risks ImpactRating5 = Cat

LikelihoodA = high

Total Score

Geology and Mining3) Difficulty in controlling grade 4* B 4(E)4) Difficulty in handling the over-burden (Clays) 3 B 9(H)5) Difficulty in handling the ore 3 B 9(H)

Processing3) Clays increase difficulty of processing 3 A 6(H)

Infrastructure1) Insufficient water for processing from local aquifers 5 C 5(H)

Permitting1) Inability to achieve mining licence in 12 months 4 C 8(H)

Marketing1) Difficulty in finding a buyer for the concentrate 5* C 5(H)

Other risks1) Stakeholder engagement + expectations 5 D 7(H)

Table 3: Highest Risks for the Wonawinta Project – July 2009

The risk assessment process highlighted key risks. Impacts were articulated and mitigation strategies were developed for each of them. Several examples are given below. Difficulty in Controlling Grade

• Ore cannot be visually distinguished from the waste • This could reduce grade from 105 to 90 g/t, reducing the project value by 35%. • Mitigation strategies were developed including:

– Definitive grade control study during the feasibility study, and – Continuous grade drilling

Clay Content in Ore and Waste

• Clay content of the ore and waste is likely to cause significant handling difficulties both in the mine and in the process plant.

• Mine: moisture & clay = very sticky & slippery = poor productivity, increased costs, increased safety risks.

• Plant: materials handling, dispersing the clay in the slurry, and increased water demand.

• Mitigation strategies included: – Mine feasibility: rigorous geomechanical assessment of ore and overburden, – Mine operation: appropriate selection and specification of equipment, and – Plant: slurry ore ASAP, and careful design of the process plant

Inability to achieve mining licence in 12 months

• Permitting is on the critical path and timeframe is constrained to achieve production. • Moderate likelihood that the project could be delayed by 6 to 12 months should any

difficulties be encountered. During this period a maiden reserve estimate of 4.6M tonnes at 97g/t Ag and 1.4% Pb, for 14.3Moz of silver was released. (Cobar, 2010). The project economics were updated with silver prices of the time and the after tax NPV was calculated as $46M. The techno-financial



model and risk assessment activities directed the focus of the work until September 2010 when the feasibility study report was delivered to the independent technical expert for review. This is a critical path element in obtaining bank finance. During that period based on the outcomes of the risk assessment, the mitigation strategies were further developed and actions taken. Examples of these actions are given below. Difficulty in controlling grade.

– Applied a global 3g/t dilution factor to the reserve. – Grade control drilling is planned at 20 x 40 m spacing. – GPS location will be used to improve accuracy. – Field XRF with sensitivity to ~10ppm will be employed. – There is a clear mining contract that defines the scope of responsibility for

grade control and the processes to be employed. Difficulty in handling overburden/ore.

– More geotechnical work has found that the materials have better handling characteristics.

– The review of equipment and contractors has identified both equipment and contractors with capability in this type of ore.

– In the plant, the slurry can be processed with pH control in a range that is operationally manageable. Early slurrying reduces handling of ore

Clays increase the difficulty of processing.

– Testing of ore slurrying and slurry behaviour has demonstrated reasonable handling characteristics at the desired solids content.

– A log washer was considered as an alternative to a drum scrubber. The log washer will handle the hardness of any of the ore material tested so far.

– The geotechnical reports indicate that the ore is of lower plasticity and so unlikely to agglomerate.

– Sample preparation for HLS analysis indicates that the ore is easily comminuted by water on a 3mm sieve.

Marketing – inability to find buyers.

– Potential buyers have been identified. There were no adverse indications that Pb concentrate cannot be sold.

On completion of this work, the model was updated with improved information and the risk assessment was done again. The comparison of the risk assessment is shown in Table 4. Only one risk (water availability) remains high (H). A drilling program was commissioned to test the flow from several additional water sources. The sensitivity analysis has evolved to that shown in Fig. 9 and Fig. 10. The sensitivity to mining factors (strip ratio, OPEX) has further declined as the mining plan has been refined. Silver recovery continues to dominate reflecting its role in revenue generation for the project. Aspects such as the capital cost of the TSF, or the capital cost of the water pipeline, on the other hand, were well down the list in terms of impact.



Geology, Exploration and Mining 8 July 2009 6 Sept 2010

Difficulty in controlling grade 4 B 4(E) 2 C 18(M)

Difficulty in handling overburden 3 B 9(H) 1 C 22(L)

Difficulty in handling ores 3 B 9(H) 1 C 22(L)

Engineering Construction Commissioning Processing

Clays increase the difficulty of processing 3 A 6(H) 3 D 17(M)

Infrastructure

Insufficient water for processing from local aquifers 5 C 5(H) 5 D/E 7(H)

Permitting

Inability to achieve mining license in 12 (now 5) months

4 C 8(H) 4 D 12(S)

Other Risks

Marketing – Difficulty in finding buyer 5 C 5(H) 2 C 18(M)

Table 4: Comparison of Risk Assessments

Sensitivity Analysis

-0.4%

-0.4%

-0.5%

-0.5%

-0.7%

-0.8%

0.7%

-1.3%

-2.0%

2.0%

3.3%

3.3%

-4.3%

3.9%

4.30%

21.5%

21.5%

22.3%

0.4%

0.4%

0.5%

0.5%

0.7%

0.8%

-0.9%

1.3%

2.0%

-2.4%

-3.3%

-3.3%

4.3%

-4.7%

-4.3%

-21.5%

-21.5%

-22.3%

-0.3%

-0.4%

0.3%

0.4%

-30% -20% -10% 0% 10% 20% 30%

Rosetta Engineering & Project Margin

Power Consumption

Water Pipeline

Leach & Adsorption CAPEX

Elution & Carbon Regen CAPEX

OPEX - Labour - Mining / Geology

Consumption of Zinc

Opex Mining - Ore

Ore SG

Mill Labour

NaCN Addition to plant

Concentrate Grade - Pb

Pb HeadGrade (%)

Grav Sep - Lead Recovery

Opex Mining - Waste

Waste SG

Strip Ratio

Silver Adsorption Efficiency

CIL - Silver Recovery

Ag HeadGrade (g/t)

Infl

ue

nc

ing

Fa

cto

rs

Change in NPV

-10%

10%

Fig. 9: Sensitivity Analysis – Sept 2010



Range Sensitivity Analysis

0.4%

0.5%

0.5%

-0.7%

0.7%

0.8%

2.6%

1.3%

2.0%

2.1%

-3.3%

-2.7%

-3.30%

4.30%

-6.9%

-21.2%

-21.7%

-22.30%

-0.4%

-0.5%

-0.5%

0.6%

-0.7%

-0.8%

-1.3%

-2.0%

-2.1%

1.1%

2.4%

3.3%

-4.3%

4.6%

3.0%

4.3%

22.3%

0.4%

0.3%

0%

-0.4%

-0.3%

-30% -20% -10% 0% 10% 20% 30%

Rosetta Engineering & Project Margin

Power Consumption

Leach & Adsorption CAPEX

Elution & Carbon Regen CAPEX

OPEX - Labour - Mining / Geology

Ore SG

Consumption of Zinc

Opex Mining - Ore

Water Pipeline

Mill Labour

NaCN Addition to plant

Opex Mining - Waste

Silver Adsorption Efficiency

Waste SG

Pb HeadGrade (%)

Strip Ratio

CIL - Silver Recovery

Grav Sep - Lead Recovery

Concentrate Grade - Pb

Ag HeadGrade (g/t)

Infl

ue

nc

ing

Fa

cto

rs

Change in NPV

Max

Min

Fig. 10: Range Sensitivity Analysis – September 2010

The project NPV has increased again based on previously used silver prices. Since then further work has refined the operation of the gravity circuit and selection of appropriate equipment, carbon elution and Merrill Crowe packages have been specified and ordered and detailed design has advanced to near final stages. On 1st March 2011 Cobar announced the securing of the debt finance portion of the funding package for the development of the project (Cobar, 2011) and a parallel equity raising for the remainder of the finance.

CONCLUSIONS By following this approach the feasibility study was completed to bankable level for approximately $3.5M or around 7 cents per resource ounce. At every key milestone of the project, the quantitative risk assessment model indicated positive project values and lessening of risk through the progressive attack on key risk factors or those demonstrating high influence on the project. The QRA is the key tool to identify whether, at an early stage, the project should be failed. It directs attention to the issues that need to be addressed and enables staged expenditure that is the most efficient use of shareholder funds. In this way a project that is not feasible can be identified early and either abandoned or the issues contributing to the poor economics addressed to see if there are alternatives that may make the project viable.



ACKNOWLEDGEMENTS

The author acknowledges Cobar Consolidated Resources Limited for permission to publish this paper. My colleagues at RMDSTEM Limited who have developed and pioneered this approach are also recognized for their efforts and support.

REFERENCES Carter, A, Gillespie, B and Gilbert C, 2009, Finding cost efficiencies in mining operations through effective value driver

modeling, Price Waterhouse Coopers, Performance Improvement Group, Brisbane, February 2009

Cobar Consolidated Resources Limited, 2008a, ASX Announcement 31/10/08

Cobar Consolidated Resources Limited, 2008b, ASX Announcement 27/11/08

Cobar Consolidated Resources Limited, 2010, ASX Announcement 18/06/10

Cobar Consolidated Resources Limited, 2011, ASX Announcement 1/03/11

RMDSTEM Limited, 2010, Cobar Consolidated Resources Wonawinta Project, Feasibility Study

Standards Australia, AS/NZS ISO 31000:2009 Standard for Risk Management



An Approach to High Solids Slurry Pipeline Design

M Griffiths1 and N Steward2

1. MAusIMM, Project Engineer, Golder Associates Pty Ltd, Building 7, Botanicca

Corporate Park, 570 - 588 Swan Street, Richmond Vic 3121. Email: [email protected]

2. Principal Engineer, Golder Associates Pty Ltd, Building 7, Botanicca Corporate Park, 570 - 588 Swan Street, Richmond Vic 3121. Email: [email protected]

ABSTRACT

This paper outlines an approach to high solids slurry pipeline design with data generated from a closed pipe-loop test system. The pressure and flow rate data generated is used to develop a pseudo-shear diagram for the purposes of system design. Following this approach enables the design of a full pipe flow system maximising system life and reliability. The alternative free fall system is explained and the disadvantages highlighted. Environmental and social considerations are discussed; mainly water and land use, and how designing for high solids slurry pipeline systems can be used improve these aspects of a mining project. Also discussed is how the use of high solids slurry pipeline systems can help to mitigate the risk associated with surface tailings storage facility failure, with the Bulyanhulu project used as a positive example.

INTRODUCTION

The design of high solids concentration tailings slurry disposal and mine paste backfill pipeline systems require the determination of the pipeline friction losses generated by the slurry. This friction loss data is the key to designing and operating any high solids concentration slurry transport system. Without accurate friction loss data the design is based on engineering assumptions, which result in a high risk for the client and the designer.



Fig. 1 Closed loop pipeline test rig.

The friction losses of the thickened slurry can be measured over a range of flow rates through various pipe sizes in a closed loop pipeline test facility. Fig. 1 shows a typical closed loop pipeline test facility during an in situ mine site tailings transport investigation. A closed loop pipeline test facility allows the slurry condition, both solids concentration and additives, to be varied easily. This test methodology has geometric similarity, and in many instances is similar in size to the final piping system. Rheological evaluation of tailings slurries from Angas Zinc, Mt Isa, Morningstar Gold and Cadia East, as well as a tuff based backfill product, has been carried out in Australia. Reports have been produced using the rheological results generated in the closed loop pipeline test facility proposing or specifying the required systems pipeline, pump and motor requirements. The closed loop pipeline test facility is the first step in the rational design of a slurry reticulation system. The risk in pipeline design lies in describing the flow behaviour, or rheology of the thickened slurry incorrectly, i.e. how the thickened slurry reacts in a pipeline. The flow behaviour, or rheology, dictates the interaction, final selection and specification of piping, pumps and electrical requirements. Undertaking pipeline test work reduces the project risk as regards paste backfill and thickened tailings transport system design.



The objectives of this paper are to: 1) Place thickened tailings disposal and paste backfill into industry context, 2) Discuss free fall and full flow underground backfill system design, 3) Define the typical reticulation system design process, 4) Outline pipeloop testing, and 5) Present rheological data from the pipeloop testing and how it is used in

system design.

INDUSTRY CONTEXT The main drivers in the mining industry behind the use of a thickened tailings disposal or paste backfill are its environmental benefits, better product delivery and its improved underground support properties. Thickened tailings use less water to transport high concentration tailings slurries for disposal, which is one obvious benefit to the mining industry, (Cooke, R). Less obvious is the significant reduction in the risk associated with the possibility of failure of the typical tailings storage facility. Using less water to transport solids means that the process water recovery can be controlled and maximised for reuse within the processing circuit. Less water is retained within the tailings storage facility, and smaller process and return water infrastructure, including pumps and piping, are required. Pumping a high solids concentration slurry will also result in less air space being required per tonne of tailings to be stored, hence a smaller storage facility as well as facilitating co-disposal with the coarse mine waste. Thickened tailings, as used for the Bulyanhulu tailings storage facility in Tanzania (Figs. 2 and 3), will also reduce the risk of liquefaction in the event of failure of a tailings storage facility, (Theriault et al.). In this instance each layer of deposited thickened tailings dries to form a hardened thin layer. By alternating deposition points around the tailings storage facility the drying time the tailings needs is provided, allowing the dry material to be built up over time into a stable facility.



Fig. 2: Thickened tailings disposal at Bulyanhulu.

Fig. 3: Thickened tailings disposal at Bulyanhulu.

Very high solids concentration tailings slurries, or paste, are used underground to backfill mined voids. This benefits the mining operation by:

1) Providing geotechnical support to the local country rock, and 2) Disposing of tailings underground.

Paste backfill typically involves combining the full tailings stream with a binder such as cement, or a cement blend, to provide the strength required from the paste backfill. Addition of cement also improves the rheology of the slurry through the introduction of fines that hinder the solids from settling out, allowing lower flow velocities to be used. The increased solids concentration of the paste backfill (lower water content) results in a lower binder requirement to achieve the support strengths necessary underground.



PASTE BACKFILL SYSTEM DESIGN

Typically a binder is added to paste backfill when it is to be used for underground support. The binder is needed to stabilise the paste backfill to prevent liquefaction and to meet the geotechnical strength requirements. The vertical component of the underground paste backfill system generates energy that must be dissipated. By maintaining full pipe flow conditions the velocity within the system can be controlled. If the velocity is not controlled, the system will allow the paste backfill to free fall and result in excessive wear of the pipe, borehole pipe damage at the air/backfill interface and ultimately, a system failure, (Steward, N R and Slatter, P T, 2009b). A full pipe flow backfill system dissipates the vertical component by using the system friction losses to control the flow velocity of the paste backfill. The paste backfill is conveyed (pumped or dropped) to the top of the vertical pipeline, or borehole, and the gravity force generated by the mass of the paste backfill and the elevation difference between the surface and the bottom of the pipe produces the head available to deliver the backfill to the stope (Fig 4). Balancing the system, balancing the available head and the pipeline frictional losses, requires accurate friction loss data for the paste backfill. The wear of the piping system can minimised by controlling the velocity within the system.

Fig. 4: a balanced full flow system. In a free fall backfill system the material is pumped from the production plant to the top of the vertical pipeline, or borehole, and free falls under gravity to impact with the air/backfill interface at some point within the pipe as can be seen in Fig. 5. The void left between the air/backfill interface and the top of the bore hole allows the backfill to free fall under gravity. This uncontrolled free fall increases the velocity of the backfill material and is likely to result in abrasive wear on the pipe wall and impact damage at the air/backfill interface. This is due to the back pressure of the system not being sufficient to balance the energy generated by the vertical component. The column of

Pressure

profiles

Vertical

pipeline

Horizontal pipelineStope

SurfaceBackfill



paste backfill self regulates, or self balances, the distribution system pressure losses by increasing or decreasing the height of the vertical column to overcome these pressure losses.

Fig. 5: The free fall system. The correct size of the vertical pipe in the distribution system can only be specified accurately if the frictional loss of the system for the paste backfill is correctly and accurately determined. This then results in a controlled backfill transport velocity, ensuring full flow and system integrity. For a given solids concentration and flow rate, a reduction in the pipe internal diameter increases the frictional losses and increases energy dissipation. This results in a lower flow velocity and hence a minimisation of abrasive wear and impact damage. Impact damage caused by the free fall of paste backfill leads to failure of the paste backfill system if not mitigated.

Pressure

profiles

Vertical

pipeline

Horizontal pipeline

Stope

Backfill

Free fall – backfill interface

Surface



SYSTEM DESIGN PROCESS

The development of a high solids concentration pumping and pipeline distribution system follows the following design process: 1) Determine the geotechnical requirements of the backfill material (tailings or

other).

2) Evaluate the materials to be used e.g. rock, aggregate, sand, tailings.

3) Determine the environmental impact of the backfill material to be used, i.e. effects from flushing, water run-off, surface disposal, chemical run-off etc.

4) Design the backfill product, based on the available materials, through carrying out tests to meet the Geotechnical Engineering/ Environmental/ transport requirements.

5) Design the high concentration backfill slurry production plant required to produce the required backfill material.

6) Carry out pipe loop tests on the backfill material in order to determine the pipeline pressure gradients for pump and motor selection, and for the pipeline specification. Where the backfill material cannot be transported by pipeline, alternative transport methods must be evaluated.

7) Finalise the transport system design, borehole, wheeled transport, pipeline or a combination of transport methods.

8) List the equipment requirements of the total production and reticulation system for both surface and underground operations.

9) Develop the general arrangement diagrams, process flow diagrams and piping and instrumentation diagrams for the backfill material production plant and distribution systems.

10) Present the options to the client.

11) Client selects the backfill disposal system for design.

12) The engineering drawings of the backfill material manufacturing plant and distribution system, as well as pipeline stress analyses, are carried out to the required codes of practice.

13) Construct the backfill material production plant and distribution system.

14) Document the completed total system; design, operational, safety, trouble shooting and quality control procedures.

15) Commission the system, including “as built” documentation, review and hand over.

Each point above can be detailed in terms of its requirements and deliverables depending on the specific operation.



CLOSED LOOP PIPELINE TESTING METHODOLOGY

Backfill material and system testing is carried out by pumping the test backfill material through a closed loop pipeline and measuring the pressure losses over a range of flow rates and in a range of pipe diameters at specific points within the system, (Steward, N R and Slatter, P T, 2009a). Prior to any test being undertaken, the closed loop pipeline is validated by testing the system with water. The pressure gradient vs flow rate data for water transported through the closed loop pipeline is compared to the Colebrook-White and/or Hazen-Williams equation. Correlation of the data with the established water curves confirms that the equipment and system configuration are working correctly. Hence any data generated using the system can be used confidently for design of the backfill reticulation system, as well as the specification of the associated equipment. The tailings to be used for backfill are sampled from either the mines processing plant or pilot plant operations, and loaded into either polypropylene bins or bulk bags. Any water decanted plus any additional water required for the test work is stored in drums to be used during testing. Supernatant water is removed from the settled tailings and results in a higher initial solids concentration, which is the typical maximum starting solids concentration for the test program. If a higher solids concentration is required the material will be dewatered further typically through draining or pressure filtering while for lower solids concentration tests, supernatant water recovered, and any make up water supplied, is used. The tailings sample is loaded into the agitated hopper of the closed loop pipeline system and pumped though the system including the preselected test pipe section. The test process involves taking pressure readings while increasing the tailings flow rate from a minimum to a maximum and back to the minimum flow rate achievable without blockage. This completes a test cycle for the selected test pipe size and tailings solids concentration. From this data the pressure gradient (∆P) and bulk velocity are calculated. These parameters become part of the inputs to the pseudo shear diagram used for the design of the backfill system. This test work and resultant data also give an understanding of the low velocity behaviour of the backfill material. Its settling/non-settling characteristics will dictate how the backfill slurry is to be treated and is an important design consideration. Water is added to the tailings sample in the closed loop pipeline to reduce the tailings concentration in preparation for the next test in the same test pipe size. The subsequent tailings solids concentration test is run in the same manner as the previous test, from the minimum to the maximum flow rate and then back to the minimum, monitoring the pressures generated. The test sequence for the selected pipe size is completed after the tailings solids concentration has been reduced to the minimum required in the test program.



Fig. 6: Pressure gradient vs. velocity diagram produced in the closed loop pipeline The next test pipe size is then selected and the test sequence is run again. This process is carried out for a minimum of three pipes sizes. The data collected includes:

Solids concentration, Flow rate, Differential pressures, Viscosity, Pump output pressure, and System temperature.

This data is converted initially into flow rate vs. pressure gradient graphs (Fig. 6).

THE PSEUDO SHEAR DIAGRAM

The flow rate and pressure gradient data is converted into a pseudo shear diagram. An example of this is given in Figure 7. In the pseudo shear diagram the data generated for different pipe diameters is co-incident over the range, for the specified solids concentrations, where the flow in the pipe is laminar. The intersection of the co-incident data at the ordinate is the material yield stress, i.e. that pressure required to make the material move.



Fig. 7: Pseudo Shear diagram for design, produced from data of pressure gradients versus flow rate produced in the closed loop pipeline Where the individual pipe sizes break away from the laminar region is termed the laminar to turbulent transition, and the relationship after that breakaway is indicative of turbulent flow behaviour within the pipeline. The co-incident data relationship lends itself to scaling the data to larger pipe sizes than those tested, providing that the laminar to turbulent and turbulent flow regions are understood. Transition prediction methods such as the Slatter Re3 model can be used to determine at what point the transported tailings transitions to turbulence. However, this model needs to be understood in the context of the rheology and the product being transported, as the model is not predictive for all material types.



CONCLUSIONS

For the design of any high solids concentration slurry system (i.e. paste backfill or thickened tailings), the pipeline frictional losses must be accurately determined to ensure a successful and competent reticulation system. This data is required to determine the correct pipe sizes for the reticulation system, as well as to specify the required pump and motor sizes. Closed loop pipeline testing is able to accurately measure flow rates and pressure losses from which pressure gradient vs bulk velocity diagrams can be developed, and hence the pseudo shear diagrams, which are used for the design of the system. This closed loop pipeline testing method of rheological evaluation uses pumps and pipe work that if not similar in size, are geometrically similar to the final system, as well as incorporating the instrumentation that would be used in a thickened tailings or paste backfill system. This decreases the design risk and increases confidence in the final system. Full flow system design underground paste backfill distribution systems designed using the closed loop pipeline testing process enables the velocity of flow in the vertical pipeline/borehole system to be controlled, hence minimising pipe wear and the risk of system failure.

REFERENCES

Cooke, R. High Concentration Tailings Transportation System Optimisation. Paterson & Cooke Consulting Engineers (Pty) Ltd, South Aftica. Paterson, A and Cooke, R, 1997. Patterson & Cooke Consulting Engineers 1997 Text Book. Slatter, P, 1999. The Role of Rheology in the Pipelining of Mineral Slurries. Min. Pro. Ext. Met. Rev. 1999, Vol. 20: pp281-300. Steward, N R and Slatter, P T, 2009a. The transport of fly ash pastes through pipelines. Australian Bulk Handling Review July/August 2009, pp 88-94. Steward, N R and Slatter, P T, 2009b. The ramifications of free fall in a backfill distribution system. Theriault, J A, Frostiak, J and Welch, D. Surface Disposal of Paste Tailings at the Bulyanhulu Gold Mine, Tanzania. Golder Associates Ltd and Barrick Gold Corporation.


The Integration of Geometallurgy with Plant Design

G Harbort1, G Cordingley2 and M Phillips3

1. MAusIMM, Process Manager, AMEC Minproc Limited, Level 2, 10 Browning Street, West End Qld 4101. Email: [email protected]

2. Formerly: General Manager Operations – Andash Project, Andash Gold Ltd. 3. MAusIMM(CP), Principal Process Engineer, AMEC Minproc Limited, Level 2, 10 Browning Street,

West End Qld 4101. Email: [email protected]

ABSTRACT Traditional engineering design for flotation circuits use the “Rule of Thumb” approach. Typically this uses a

nominated maximum head grade for design. A scale-up factor is applied to laboratory flotation tests for

residence time. This scale-up factor is usually based on a designers experience at other flotation sites and

may or may not be relevant to the circuit being designed, or the flotation equipment being used. A small

number of locked cycle tests are assumed to represent the ore body and the results of these are often used

for financial analysis, independent of changes in mine plan throughput and mineralogy.

The engineering “Rule of Thumb” approach is only strictly accurate where there is very little variability in

throughput, head grade and mineralogy. It may provide accurate design for mature established operations

where a brownfield expansion is being considered but is unlikely to provide accurate design for a new,

greenfield site.

The use of geometallurgical modeling with floatability component simulation provides a design methodology

with significantly less associated risk. The use of geological data for optimisation of operating plants has

become a significant part of the modern process mindset. The underlying principle is to use spatial

metallurgical information to drive production planning, mine planning, blast design, blending strategies and

plant set-up. At a design stage the process designer can use geometallurgical information to evaluate bottle

necks and potential design flaws and propose the best investment strategy for the project benefit.

This paper details the geometallurgical characterisation of the Andash deposit and methodology used to

review the project’s detail design and projected production.

INTRODUCTION The Andash Gold-Copper Project is located in the Kyrgyz Republic (formerly Kyrgyzstan) approximately

260km from the capital, Bishkek in the Talas region. The Andash ore body has an estimated reserve of

540,000 oz of gold and 63,000 tonnes of copper. The gold-copper concentrator will be developed in two

phases. The Phase 1 concentrator is designed for a throughput of 1.5 Mtpa and Phase 2 will constitute an

expansion to 3 Mtpa. The targeted annual production will be 60,000 oz of gold and 5,000 tonnes of copper in

Phase 1 followed by 80,000 oz of gold and 6,900 tonnes of copper in Phase 2. The Andash concentrator

design is based on a maximum gold head grade of 1.65 g/t and a maximum copper head grade of 0.60%.


PLANT DESIGN The concentrator will utilise the following principal process areas for the recovery of the gold bearing copper

concentrate:

Three stage crushing, ore storage and reclaim;

Single stage ball milling (expanded to primary and secondary ball milling in Phase 2);

Flash flotation (Phase 2);

Sequential rougher flotation (utilising split sulphide and oxide circuits);

Combined cleaner, cleaner scavenger and re-cleaner flotation;

Sulphide concentrate re-grind (Phase 2);

Concentrate de-watering, filtration, storage and load-out and tailings thickening (Phase 2);

Disposal and decant water return.

Further details on the plant design are given in the following sections, with the Phase 1 flowsheet shown in

Fig 1 and the layout in Fig 2.

Fig 1. Andash Phase 1 Circuit Flow Sheet

Fig 2. Three Dimensional View of the Andash Circuit


Crushing Run of mine (ROM) ore will be withdrawn from the ROM bin by a vibrating grizzly feeder, with 150 mm bar

spacing. Oversize from the grizzly feeder will discharge directly to the open circuit primary jaw crusher. The

jaw crusher product (P80 125 mm) will discharge onto a sacrificial conveyor, where it will re-combine with the

grizzly feeder undersize, and be conveyed to the secondary crusher sizing screen. A weightometer will be

installed on the primary crusher sacrificial conveyor for throughput measurement, control and tonnage

accounting. A continuous belt magnet and metal detector will be installed on the sacrificial conveyor for

tramp metal detection and removal.

The undersize from the secondary crusher sizing screen reports to the tertiary crusher sizing screen feed

conveyor. The screen oversize will feed directly to the secondary cone crusher. During Phase 2 a feed

splitter and additional secondary cone crusher will be installed. The feed splitter will result in an equivalent

feed reporting directly to each of the two secondary crushers. Secondary crusher product will recombine

with the undersize on the tertiary crusher sizing screen feed conveyor.

Oversize ore from the tertiary crusher sizing screen will be conveyed to the tertiary crushing circuit feed bin.

Ore will be withdrawn from the tertiary crusher feed bin by a variable speed, belt feeder and fed directly to

the closed circuit tertiary cone crusher. Tertiary crusher product will discharge onto the tertiary crusher sizing

screen feed conveyor with oversize thereby returning to the tertiary crusher sizing screen. A feed splitter,

second feed bin and second tertiary crusher are included in Phase 2.

Undersize material from the tertiary crusher sizing screen (P80 - 10 mm) will report to the crushed ore

stockpile prior to reclaim by four variable speed vibrating pan feeders discharging to the ball mill feed


conveyor. The mill feed conveyor will be equipped with a weightometer for feed rate control and mass

measurement of the ore reporting to the ball mill.

During Phase 1 bulk supply hydrated lime will be added directly onto the load of the mill feed conveyor via a

lime silo and screw feeder installed adjacent to the conveyor. In Phase 2 a hydrated lime mixing tank and

lime slurry storage tank complete with standby pumps, duty pumps and ring main will be provided for pH

control via lime addition to the sulphide and oxide rougher flotation conditioner tanks as well as the primary

mill feed.

Ball Milling During Phase 1 of the project, the a ball milling circuit will be installed to grind material received from the

crushing circuit from an F80 of 10 mm down to a P80 of 56 µm. Phase 2 of the project will involve doubling the

plant throughput capacity with the installation of a second ball milling circuit. The two mills will operate in

series where the primary ball mill will grind material from an F80 of 10 mm to a P80 of 215 µm and the

secondary ball mill will grind the cyclone overflow from the primary ball milling circuit from an F80 of 215 µm

to a P80 of 56 µm.

Each mill will have its own dedicated cyclone cluster. Flotation feed will be sampled via an automated cross

stream sampler with the sample reporting to the OSA and also being collected for metallurgical accounting

purposes.

Sulphide Rougher Flotation The sulphide rougher flotation bank will initially consist of three 100 m³ forced draught tank type flotation

cells. An additional two 100 m³ forced draft tank type flotation cells will be added to the circuit for the Phase

2 expansion.

Concentrate from the sulphide rougher flotation circuit will be collected in a concentrate launder and will

gravitate to the sulphide rougher concentrate hopper via an in-line launder sampler. During Phase 1 sulphide

rougher concentrate will be pumped from the concentrate hopper to the cleaner feed box. During Phase 2

sulphide rougher concentrate will be pumped to the regrind mill cyclone cluster. Tailings from the sulphide

rougher flotation circuit will be pumped, via an-inline pipe sampler, to the first of two oxide rougher flotation

conditioners. Both the concentrate and tails sample collected will report via separate sample hoppers to the

OSA.

Oxide Rougher Flotation Oxide rougher feed will be conditioned in two stages. The oxide rougher conditioners will operate in series

with overflow from the first conditioner gravitating to the second conditioner. Sodium Hydrosulphide (NaHS)

will be added to the slurry within the first oxide rougher flotation conditioner to sulphidise the oxide mineral

surface (Controlled Potential Sulphidisation) and enhance the floatability of the oxide minerals.

Primary and secondary collector will be added to the slurry within the second oxide rougher conditioning

tank. The conditioned slurry will gravitate from the second conditioner tank to the oxide rougher feed box.


The oxide rougher flotation bank will initially consist of three 100 m³ forced draught tank type flotation cells.

During the Phase 2 expansion an additional two oxide rougher flotation cells will be added to the circuit.

Concentrate from the oxide rougher flotation circuit will be collected in a concentrate launder and will

gravitate into the oxide rougher concentrate hopper. Oxide rougher concentrate will be pumped from the

oxide rougher concentrate hopper to the cleaner conditioner tank, via a pipe sampler. The sample will be

collected in a sample hopper and be pumped to the OSA unit.

Tailings from the oxide rougher flotation circuit will gravitate through a launder sampler prior to reporting to

the oxide rougher (final) tails hopper. The sample collected from the oxide rougher tails sampler will be

pumped to the OSA unit. Oxide rougher tails will initially be pumped directly to the TSF. Following completion

of the Phase 2 expansion, oxide rougher tails will be pumped to the final tails thickener feed box.

Gas monitors will be installed to monitor the H2S level within the oxide rougher circuit.

Sulphide Concentrate Regrind There will be no regrind mill installation in Phase 1. Two ultra fine grinding mills will be installed for Phase 2

The sulphide rougher concentrate will be pumped from the sulphide rougher concentrate hopper to the

regrind mill cyclone cluster. The regrind mills will operate in open circuit with the cyclone underflow

gravitating to the regrind mill feed hopper. Cyclone overflow and regrind mill discharge will recombine as the

regrind mill product and gravitate to the cleaner conditioner tank.

Cleaner Flotation Sulphide and oxide concentrate cleaning to a saleable concentrate grade will be achieved utilising a

common cleaner and re-cleaner circuit. Additions in Phase 2 include a common three cell scavenger cleaner

circuit and one additional re-cleaner cell. The cell volumes are 8 m3 per cell for the cleaner and cleaner

scavenger cells and 3 m3 per cell for the re-cleaner cells. The oxide and sulphide rougher concentrate will be

pumped from the respective rougher concentrate hoppers to the cleaner conditioner tank. In Phase 2 the

regrind cyclone overflow will gravitate to the cleaner conditioner. Following the addition of primary collector,

the conditioned rougher concentrate will gravitate from the cleaner conditioner to cleaner cell No1.

The cleaner circuit will consist of six cleaner cells operating in series. Sulphide and oxide rougher

concentrate will constitute the feed to the first cleaner cell. Concentrate from the first four cleaner cells will

gravitate, via two launders and a collection box, to the cleaner concentrate hopper, prior to being pumped to

the re-cleaner circuit. Concentrate from cell No 5 and 6 will gravitate via a launder to the cleaner scavenger

concentrate hopper prior to being pumped back to cleaner cell No 1 feed box. In Phase 2 the cleaner cell 5

and 6 concentrate is combined with cell 1 to 4 concentrate and pumped to the re-cleaner circuit. The cleaner

scavenger concentrate then gravitates to the cleaner scavenger concentrate hopper and is returned to the

cleaner circuit.

In Phase 1 the cleaner circuit tails will gravitate to the sulphide rougher tails hopper for return to the oxide

rougher conditioner No 1. In Phase 2 the cleaner circuit tails will constitute the feed to the cleaner scavenger


circuit. The cleaner scavenger circuit will consist of three cells operating in series. Cleaner scavenger tails

will gravitate to the sulphide rougher tails hopper for return to the oxide rougher conditioner No 1.

In Phase 1 the re-cleaner circuit will consist of two cells operating in series. During Phase 2 a third cell will be

installed. Cleaner concentrate will constitute the feed to the re-cleaner circuit. The re-cleaner tail gravitates

back to the cleaner feed box. The re-cleaner concentrate will be collected in the re-cleaner concentrate

hopper and pumped to the trash screen ahead of the concentrate thickener. The sample is collected in the

re-cleaner concentrate is sampled and pumped to the OSA for online moditoring and collection for

metallurgical accounting..

GEOLOGY Andash copper-gold mineralisation is structurally confined to a hydrothermally silicified tectonic breccia and

quartz alteration, extending from and terminating into granodiorite. The orebody is zoned as follows:

The upper section of the ore zone spreads from surface to a depth of 40m to 75m and is composed

mainly of cataclasised granodiorite-porphyry. Hydrothermally imposed ore minerals include

chalcopyrite, native gold and with hypergene minerals including cuprite, covelline, bornite, native

copper, chalcosine and hydrotite.

The middle section of the ore zone stretches from 40 - 105m with sharply increased gold and copper

content. The ore is formed by carbonate-chlorite-sericite-quartz alterations. Hypergene minerals are

significant and are represented by malachite, azurite, covelline, cuprite, native copper, drogoethite,

bornite and chalcosine.

The lower ore zone is at depths of 60 - 175m. This zone is mainly formed by chloritised and

sericitised cataclasised granodiorite porphyry with increasing chalcopyrite and pyrite content.

Numerous barren dykes of various compositions are present within the ore body.

METALLURGICAL TEST PROGRAM

Drill Core Inspection and Sample Selection A critical component of metallurgical testing for engineering design is supply of adequate sample that has not

deteriorated during storage. Physically, core was required to show minimal post drilling oxidation and fatigue

stress and provide a representative cross section of copper and gold mineralogy, radially and at depth,

representing the anticipated mine plan. In addition, core to be utilised for grinding tests required some

selective inclusion of granodiorite and barren dike material to evaluate the effect of mining dilution on

throughput. Continuous lengths of 10m core were required for flotation only tests and 20m continuous

lengths were required for flotation/comminution tests.

Andash drill core was stored in a core shed in an industrial area of Bishkek, Kyrgyz Republic (Fig 3). Over a

two day period all drill core that may have been suitable for test work was inspected. All core examined,

except one hole, was labelled correctly according to the date of drilling, depth and hole location. Natural


transitions within core were geological and visual inspection of copper mineralogy and litholgy showed it to

relate to the core data base. Although the incorrectly labelled core was of good quality and obviously from

the Andash deposit, uncertainty relating to its origin in the ore body precluded its use for test work.

Fig 3. Drill Core Storage in Bishkek, Kyrgyz Republic

Iron staining was noted on a significant number of core segments. The selected breakage of core indicated

that this was in-situ in nature and had not occurred due to weathering during storage. No fatigue cracking or

stress release pitting was evident so the competency of the core in relation to storage had not been

compromised. Minor oxidation of chalcopyrite was evident, resulting in surface tarnishing. This was not

evident in freshly broken core and affected a small amount of the copper mineralisation present. The minor

oxidation observed was considered unlikely to affect the planned test program.

The location of samples selected for the metallurgical program is shown in Fig 4 and Fig 5.

Test Program Overview The test program scope was to quantify the extent of recovery and concentrate grade variability for

marketing, evaluate the reagent regime, concentrate regrind requirement, impact of copper oxides, flotation

response details for circuit design including tailings and effect of water recycle and competency for

comminution circuit design. Key components included:

Generation of two composite samples representing different copper mineralisation;

Bond Ball Mill Work Index tests on 14 samples;

Bond Rod Mill Work Index and Bond Abrasion Index tests on composites;

Lime consumption and grindability tests on composites;


Flotation optimisation tests including grind size, collector, residence time, sulphidisation, regrind and cleaner appraisal on the composites;

Composite locked cycle tests;

Flotation variability tests on 28 samples.

Fig 4. Front Elevation View of Drill Core Selected for Test Work

Fig 5. Plan View of Drill Core Selected for Test Work


Comminution The design comminution characteristics are typically selected from the variability test work results to ensure

that the majority of the ore body can be treated at the nominated design throughput rate. The standard

engineering design procedure involves selecting the 80th percentile comminution characteristic as the basis

of the comminution circuit design. The approach is based upon the basis of the hardest 20% of the ore body

being blended in with softer ore during stockpile operations, allowing the plant to consistently achieve

nameplate design capacity. Should significant variation exist, a detailed geometallurgical approach is

required. As shown in Table 1, little variation existed in the Bond Ball Mill Work Indices and a

geometallurgical approach was not used for comminution design.

Table 1. Bond Ball Mill Work Index Summary Bond Ball Mill Work

Index (kWh/t) P80

(µm) Average 18.5 56

Standadard Deviation 0.8 2

80th Percentile 18.7 57

Sequential Copper Assay Procedure Prior flotation test work had indicated high variability in total copper recovery (Wardell Armstrong

International, 2007). Little attempt was made to identify the various copper species and their effect on copper

recovery.

The AMEC Minproc test work program used a sequential copper determination to determine the predominant

copper mineralogy of the core samples utilised. The sequential copper analysis was a modification of the

Tessier et al, 1979 procedure and involved:

De-ionised water dissolution to determine the quantity of water soluble copper;


Acetic acid dissolution to determine the quantity of malachite and azurite;

Weak sulphuric acid dissolution to determine the quantity of malachite, azurite, tenorite, chrysocolla and cuprite (partial);

Cyanide dissolution to determine the quantity of chalcocite, bornite and chalcopyrite (partial);

Strong acid dissolution to determine the quantity of remaining copper in chalcopyrite.

In addition MLA analysis was conducted on selected head samples and sequential assay results were

calibrated against the mineralogy using a least squares minimisation technique.

The sequential copper analysis confirmed that the copper mineralisation within the Andash ore is extremely

variable (Fig 6) with dominant copper oxides (malachite, azurite), secondary copper sulphides (chalcocite,

bornite) and chalcopyrite.

Fig 6. Andash Copper Mineral Distribution

0%

10%

20%

30%

40%

50%

60%

70%

80%

90%

100%

Co

pp

er Distrib

utio

n

Sample1

Sample3

Sample6

Sample8

Sample10

Sample12

Sample14

Sample16

Sample18

Sample20

Sample22

Sample24

Sample26

Sample28

Malachite/ Azurite Tenorite/ Copper Chalcocite/Bornite Chalcopyrite

Flotation Test Work

Optimisation Tests The optimisation of flotation conditions was conducted on two different composites. These composites were

formed based on the sequential assays of individual intersections and were a sulphide copper dominant

composite and an oxide/carbonate copper dominant composite.


Prior tests had established that a flotation feed P80 of 56µm for maximum copper recovery and this figure

was used throughout tests.

All rougher flotation tests used the AMIRA P9 Floatability Component Batch Test flotation procedure with a

bottom driven 5 L flotation cell. Cleaner and recleaner tests were done with the nominated laboratory’s

procedures, modified if necessary to ensure floatability component modelling could be conducted. Rougher

tests were conducted to visual exhaustion.

Reagent Evaluation The reagent selection was conducted in two phases:

1. Reagent addition to optimise sulphide copper recovery and

2. Reagent addition to the sulphide float tailing to optimise oxide copper recovery.

The reagents trialled are detailed Table 2. The collectors PAX, A3477 and A3926 were selected primarily for

their ability to recover copper sulphides or copper oxides following sulphidisation by NaHS. AM2, A404 and

XD902 were selected based on their reported ability to recovery both copper sulphides and copper oxides

simultaneously.

Table 2. Flotation Reagents Used in the Test Prrogram Product Description AM2 Potassium n-octyl hydroxamate PAX Potassium amyl xanthate A3477 Sodium isobutyl dithiophosphate A404 Blend of dithiophosphates and mercaptobenzothiazole A3926 Alkyl thionocarbamate XD902 Alkyl hydroxamates NaHS Sulphidising agent

Sulphide Composite Reagent Selection AM2 and XD902 produced the highest copper recovery in a sulphide only float. This appears due to the

flotation of some copper oxides in additional to copper sulphides. Sequential sulphide / oxide rougher

flotation was tested under the following conditions:

1. AM2 only to produce a bulk sulphide / oxide concentrate,

2. PAX to produce a sulphide concentrate, with AM2 to produce an oxide concentrate and

3. PAX to produce a sulphide concentrate, with NaHS to -200mV SHE and PAX to produce the oxide concentrate.


All three options achieved an equivalent ultimate total copper recovery, approaching 90%. AM2 when used

alone proved very non-selective, producing the lowest grade-recovery curve. PAX, with either AM2 or NaHS

in the oxide float produced the highest grade-recovery response.

Gold recovery in the sulphide concentrate was directly related to the recovery of copper sulphides, with

ultimate gold recoveries of 85% achieved.

Oxide Composite Reagent Selection PAX produced the highest copper recovery in a sulphide float, achieving approximately 80% recovery of the

copper sulphides present.

Sequential sulphide / oxide rougher flotation was tested under the following conditions:

1. Comparison of AM2 and XD902 to produce a bulk sulphide / oxide concentrate and

2. PAX to produce a sulphide concentrate, with various NaHS additions to produce an oxide concentrate.

Treatment with AM2 achieved the highest copper recovery (62%), albeit at a high consumption rate of 400

g/t. When the AM2 consumption rate was reduced to 200 g/t the copper recovery decreased to 40%. By

comparison XD902 achieved a copper recovery of 50%, with a consumption of 200 g/t.

The oxide float sulphidisation tests indicated the optimum Eh for the rougher oxide float was between -300

mV and -400 mV (SHE), with optimum gold recovery between -200 mV and -300mV (SHE).

A sequential float, utilising PAX and NaHS produced the best grade recovery curve.

Rougher Concentrate Regrind Requirements To determine which streams could benefit from regrinding, aliquots of rougher concentrates were analysed in

a laser particle sizer. From these results it was determined that oxide flotation concentrates were finely sized

and are unlikely to require regrinding, while sulphide rougher concentrates may require minor regrinding to a

P80 of 20 µm. This was confirmed in subsequent tests

Cleaner Circuit Evaluation Cleaner flotation tests were conducted on both oxide and sulphide composites to evaluate the effect of:

1. Cleaning the sulphide float and oxide float concentrates in separate cleaner circuits and

2. Cleaning the sulphide float and oxide float concentrates in a combined cleaner circuit

When treating either the sulphide dominant copper ore or oxide dominant ore, combining the sulphide

flotation and oxide flotation rougher concentrates into a single cleaner stream produced improved copper

recovery and grade, when compared to split cleaner treatment. This was in direct contrast to prior test work.


Re-cleaner flotation tests were conducted on both oxide and sulphide composites to evaluate the effect of:

1. Re-cleaning the sulphide float and oxide float concentrates in separate cleaner circuits and

2. Re-cleaning the sulphide float and oxide float concentrates in a combined cleaner circuit.

When treating the sulphide dominant copper ore, combining the sulphide flotation and oxide flotation rougher

concentrates into a single cleaner and re-cleaner stream produced improved re-cleaner copper recovery and

grade, when compared to split cleaner treatment. The comparison for oxide dominant ore was ambiguous.

Locked Cycle Tests The locked cycle tests were undertaken when the majority of DFS design had been completed, incorporating

a split cleaner circuit based on prior test work. As such the locked cycle tests for the two composites were

conducted using sequential rougher flotation to produce separate sulphide and oxide concentrates, which

were then cleaned in separate cleaning circuits, the sulphide rougher concentrate being reground prior to

cleaning.

A summary of locked cycle results is given in Table 3.

Table 3. Summary of Locked Cycle Test Results

Oxide Composite Sulphide Composite

Copper Gold Copper Gold

Conc (%Cu) Rec (%) Conc (g/t Au) Rec (%)

Conc (%Cu)

Rec (%)

Conc (g/t Au)

Rec (%)

Cycle 1 25.8 31.8 130.7 57.4 24.7 82.8 67.1 75.3

Cycle 2 25.9 37 119.9 67.2 23.6 81.9 62.8 70.4

Cycle 3 28.1 43.6 108.9 66.5 22.7 80.2 65.9 75.8

Cycle 4 28.1 39.3 135.4 67 24 81.8 66.0 77.3

Cycle 5 28.8 41.2 109.2 50.7 22.6 80.8 58.4 72.4

Cycle 6 30.2 40.7 128.6 66 23.7 80.5 61.9 75.4

Cycle 7 29.0 43.6 98.4 65.0 22.6 82.9 61.4 75.3

Cycle 8 25.9 42.9 97.6 66.1

Cycle 9 25.2 46.7 104.3 69.0

Variability Tests Rougher flotation variability tests were conducted on each of the individual samples. The flotation conditions

were set according to whether the samples were oxide/carbonate copper dominant or sulphide copper

dominant. The rougher flotation conditions are given in Table 4.

Table 4. Conditions for Variability Flotation Tests

Ore Designation

Grind (p80)

Sulphide Float Oxide Float

pH RT Reagents (g/t) RT Reagents (g/t) Eh

mV (min) MIBC Lime PAX (min) AM2 NaHS PAX Sulphide 56 8.5 7.5 43 615 40 20 100 10

Oxide 56 8.5 4 43 615 20 27 459 30 -


300

The variability tests confirmed that the ore body exhibits significant variability in total copper rougher

recovery and minor variability in rougher gold recovery. The wide copper variability was directly related to

copper mineral variability. The recovery of copper sulphides was relatively robust with 96% rougher recovery

when no copper oxides were present, declining to 90% at a copper oxide to copper sulphide ratio of 0.8. At

very high oxide copper content the sulphide recovery decreased significantly (Fig 7).

The rougher copper oxide recovery was heavily dependent on the ratio of acetic acid soluble copper, e.g.

malachite to weak acid soluble oxides such as tenorite. When copper carbonates were the only oxide copper

minerals present their recovery approached approach 75% and it declinied with higher weak acid soluble

copper content increased.

Due to the wide variability in these it was decided that the locked cycles tests could not effectively be used

as a measure of life of mine performance. At this stage test sample was approaching exhaustion, with limited

sample available to conduct further flotation tests.

A geometallurgical approach was adopted to evaluate performance over the mine life.

Fig 7. Effect of the Oxide Copper Content of Ore on Copper Sulphides Recovery

0

10

20

30

40

50

60

70

80

90

100

0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0

Oxide Copper to Sulphide Copper Ratio

Co

pp

er

Su

lph

ide

Ro

ug

he

r R

eco

very

(%

)

GEOMETALLURGICAL DEVELOPMENT Traditional engineering design for flotation circuits uses the “Rule of Thumb” approach. Typically this uses a

nominated maximum head grade for design. A scale-up factor is applied to laboratory flotation tests for

residence time. This scale-up factor is usually based on a designers experience at other flotation sites and

may or may not be relevant to the circuit being designed, or the flotation equipment being used. A small

number of locked cycle tests are assumed to represent the orebody and the results of these are often used

for financial analysis, independent of changes in mine plan throughput and mineralogy.


The engineering “Rule of Thumb” approach is only strictly accurate where there is very little variability in

throughput, head grade and mineralogy. It may provide accurate design for mature established operations

where a brownfield expansion is being considered but is unlikely to provide accurate design for a new,

greenfield site.

To overcome this “Rule of Thumb” approach to design Geometallurgical test programs and data review have

been developed to provide realistic information into simulation packages to provide data for process

evaluation and plant design. This approach was applied to the Andash project to aid final design and risk

minimization.

This geometallurgical development was conducted in a number of stages:

1. Incorporation of sequential copper assays into the resource model and mine plan,

2. Development of two series of floatability parameters based on the rougher performance in the locked

cycle tests,

3. Generation of a floatability component model to simulate the locked cycle tests,

4. Calibration of the model so that model results equalled actual locked cycle test results,

5. Determination of maximum mineral recovery with ore of various copper mineralogy, based on the

variability tests,

6. Development of a JKSimFloat simulator incorporating floatability data, proposed circuit design,

flotation feed rates and flotation machine characteristics and

7. Operation of the simulator on each month of operation indicated by the mine plan developed in Item

1.

Resource and Mine Plan Development Sequential assay results were input into the resource model. Development of a mine plan based on this

geometallurgical resource model allowed a mine plan and concentrator schedule to be developed based on

copper mineral species (sequential copper extractions).

The scheduled feed grade to the concentrator on a copper species basis is shown in Fig 8. Key points in the

schedule are:

At the commencement of mining heavily oxidised material was stockpiled for end of mine life

treatment. This resulted in an initial concentrator feed where the majority of copper was acetic acid

soluble, e.g malachite.

During the second year of operation the starter pit penetrates a high grade, sulphide pod, with a

significant increase in both feed grade and aqua regia soluble copper, i.e. chalcopyrite.


Year 3 represents a significant mine expansion with lower grade copper carbonate material feeding

the concentrator. This period coincides with the planned expansion of the concentrator facilities.

During year 4 increased amounts of copper oxides and copper carbonates are mined from the

expanded pit and presented to the concentrator. The deepening pit also results in significant

amounts of chalcopyrite being mined and treated.

In the ensuing three years the head grade to the concentrator remains relatively constant, although

the copper mineralogy changes significantly. Acetic acid soluble copper is progressively replaced

with chalcopyrite as the pit deepens. The amount of copper oxide material also decreases.

Year 7 represents treatment of stockpiled material with substantial amounts of copper oxides,

although acetic acid soluble copper still represents 50% of the concentrator feed.

Fig 8. Life of Mine Mineral Grades in the Concentrator Feed Schedule

Floatability Component Modelling Floatability component modelling and simulation, as developed in the AMIRA P9 project (Alexander and

Morrison, 1998) has received increasing acceptance as a plant optimisation tool. It is one of a suite of tools

that AMEC Minproc use during commissioning and plant optimisation campaigns. The Andash Project

represented its first use within the company for geometallurgical modelling and checking of detail design.

Ore floatability components were calculated from the sulphide dominant composite rougher kinetic locked

cycle test results. As no bubble size measurements were conducted , a bubble size of 1.1 mm was assumed.

It was also assumed that the batch cell froth recovery was 100% and entrainment was zero.


As the model utilised a mineral based system it was assumed that acetic acid soluble copper was malachite,

weak sulphuric acid copper was tenorite, cyanide soluble copper was chalcocite and the copper extracted in

the final aqua regia was chalcopyrite.

Floatability parameters were calculated for both sulphide rougher feed and oxide rougher feed and are

shown in Table 5. This indicated that chalcopyrite was highly floatable in the sulphide rougher feed and was

depressed in the oxide flotation stage. Both malachite and tenorite had low floatability in the sulphide

rougher. Both became floatable in the oxide rougher with the weighted floatability of malachite being

approximately double that tenorite.

Table 5. Floatability Parameters and Distributions

Mineral Location Floatability Component Parameters and Distributions

mfast mSlow m Non Float

M Total PSlow PFast

P Non Float ∑P

Tenorite Sulphide Rougher Feed 0.00 0.18 0.82 1.00 1.55E-03 1.00E-04 0.00E+00 1.77E-05

Oxide Rougher Feed 0.30 0.08 0.44 0.82 1.00E-02 1.00E-04 0.00E+00 3.65E-03

Malachite Sulphide Rougher Feed 0.00 0.11 0.89 1.00 1.55E-03 1.00E-04 0.00E+00 1.15E-05


Cu Sulphides

Sulphide Rougher Feed 0.89 0.04 0.07 1.00 1.55E-03 1.00E-04 0.00E+00 1.39E-03


Gold Sulphide Rougher Feed 0.63 0.19 0.19 1.00 1.55E-03 1.00E-04 0.00E+00 9.93E-04


A simulation was the created to match the configuration of the locked cycle test, including cleaner stages.

The model results were then compared to the measured locked cycle results to see if floatability was

conserved across the locked cycle test. The comparison indicated that mineral floatabilities in the respective

rougher concentrates were not conserved across to their respective cleaner circuit feeds. Floatabilites in the

cleaner concentrate were however conserved into recleaner feeds. The major variations in the cleaner feeds

were:

1. A decrease in the floatability of tenorite and malachite in the sulphide cleaner feed,

2. An increase in the floatability of chalcopyrite in the sulphide cleaner feed and

3. A decrease in the floatability of tenorite in the oxide cleaner feed.

A series of floatability transfer matrices were incorporated into the model to calibrate it against the locked

cycle test. This provided acceptable recoveries of copper and sulphur, with the modelled final concentrate

grade matching the locked cycle test in terms of copper and sulphur. Modelled iron grades were lower in the

final concentrate however and this resulted in over estimation of non sulphide gangue. To correct the iron

imbalance non sulphide gangue was adjusted in the model to contain iron. (Later QEM-Scan analysis of

concentrate indicated the higher iron levels were caused by the flotation of copper-goethite),


The calibrated model was then compared against the second locked cycle test and showed good agreement.

A further two locked cycle tests were conducted with different configurations and simulated. The simulated

results agreed with actual results and it was decided that the model could effectively be used to replace

further locked cycle tests and simulate full circuit operation.

Calculation of Maximum Mineral Recovery Both the floatability component model and JKSimfloat make use of the fraction of non-floating mineral in the

feed, calculated from the maximum mineral recovery. The rougher flotation variability tests provided the

maximum mineral recoveries at various mineral head grades. This indicated a multivariable linear function,

as described below with malachite as an example.

RMal = aMal . FMal + bMal . FTen + cMal . FCc+ dMal . FCp

Where RMal is the recovery of malachite; FMal , FTen , FCc and FCp are the sample head grades for malachite,

tenorite, chalcocite and chalcopyrite respectively and aMal, bMal, cMal and dMal are solved parameters.

This methodology was also used for total copper recovery and showed good agreement between modelled

copper recoveries and the variability test recoveries (Fig 9).

Fig 9. Parity Chart for Calculated and Measured Copper Recovery


Development of a JKSimfloat Plant Simulator To overcome the previously stated “Rule of Thumb” approach in plant design and evaluation the

Geometallurgically developed data if fed into Geometallurgical modelling with floatability component

simulation providing a design methodology which significantly reduces the associated project risk.

Simulator Inputs The floatability component simulations conducted for the Andash project used JKSimfloat. Key inputs to the

software included:

Throughput rates and feed percent solids from the Process Design Criteria;

Feed grades and mineralalogy from the geometallurgical mine plan;

Floatability parameters from the calibrated floatability component model. Changes in floatability were

incorporated via floatability component transfer matrices.

Nonfloating mineral fractions as calculated from the calculation of maximum recovery methodology

Flotation equipment number and sizes from the Process Design Criteria using a ”Rule of Thumb”

approach

Recommended air flow rates supplied by manufactures


Bubble size, froth recoveries and entrainment rates sourced from the AMEC Minproc optimization

database

Water additions to launders and pump hoppers from the Process Design Criteria.

Simulator Outputs The outputs from the simulator included:

Steady state residence times for all flotation cells;

Mass, water and pulp flows for all streams in the circuit;

Mineral and element distributions for all streams in the circuit;

Mineral and element grades for all streams in the circuit.

Integration with Design

Base Case Circuit Simulations Seventy-eight simulations were conducted on the Andash flotation circuit developed during the DFS. This

circuit consisted of sequential sulphide and oxide circuits and split sulphide and oxide cleaner circuits. Each

simulation was conducted on one month of the concentrator feed schedule.

These simulations were initially conducted to determine the variability of copper and gold recovery and

concentrate grade throughout the planned operating life. Copper and gold recoveries over the mine life are

given in Fig 10, with concentrate grades given in Fig 11.

The simulations provided detailed mass balances across each month of the mine plan and were used as a

cross check of the DFS design. A number of significant problems soon became apparent from using a “Rule

of Thumb” approach. Although the simulations showed only minor variation in recoveries they indicated

major variation in stream pulp flow on a month to month basis. It was expected that these variations would

become worse over shorter time frames.


Fig 10. Life of Mine Copper and Gold Recoveries

0

10

20

30

40

50

60

70

80

90

100

YEAR 1 YEAR 2 YEAR 3 YEAR 4 YEAR 5 YEAR 6 YEAR 7

Copper and Gold Recovery (%)

Copper Recovery Gold Recovery

Fig 11. Life of Mine Concentrate Grades

0

10

20

30

40

50

60

70

80

90

100

YEAR 1 YEAR 2 YEAR 3 YEAR 4 YEAR 5 YEAR 6 YEAR 7

Concentrate Copper and Gold Grade

Copper Grade (%Cu) Gold Grade (g/t Au)

The ratio of simulated residence time to the “Rule of Thumb” design residence time was used to conduct a

review of circuit stability. The residence time ratios for both oxide and sulphide split cleaner and recleaner

circuits are shown in Fig 12 and Fig 13.


In the Figures a ratio below 1.0 indicates a shortage of flotation residence time, with recovery losses. A ratio

above 3.0 indicates circuit starvation with associated feed pump cavitation, operation of tailing valves below

their recommended optimum duty points, with associated level control problems and high recirculating loads

coupled with fluctuating grade and recovery in concentrate. The preceding issues can be overcome through

increased rougher concentrate production, possibly endangering the final concentrate grade, the significant

addition of make-up water or changes to the circuit configuration.

The sulphide cleaner and recleaner circuits typically operated within acceptable limits. This was not the case

with the oxide cleaner circuit where significant variations are evident and circuit stability was expected to

cause operating difficulties and loss of metallurgical performance. The oxide recleaner circuit would never

have operated to design without major modifications.

Fig 12. Split Cleaner Circuit Residence Time Ratios

0

1

2

3

4

5

6

7

8

9

10

1 4 7 10 13 16 19 22 25 28 31 34 37 40 43 46 49 52 55 58 61 64 67 70 73 76 79

Month of Mine Plan

Rat

io o

f A

ntic

ipat

ed

Res

iden

ce T

ime

to

Nom

inat

ed D

esig

n R

esid

enc

e T

ime

Suphide Circuit Oxide Circuit Fig 13. Split Recleaner Circuit Residence Time Ratios

0

2

4

6

8

10

12

14

16

1 4 7 10 13 16 19 22 25 28 31 34 37 40 43 46 49 52 55 58 61 64 67 70 73 76 79

Month of Mine Plan

Rat

io o

f A

ntic

ipat

ed R

esid

ence

Tim

e to

N

omin

ated

Des

ign

Res

iden

ce T

ime

Sulphide Circuit Oxide Circuit


Detail Design Optimisation An initial solution evaluated was to combine sulphide rougher concentrate and oxide rougher concentrate

and treat them in a combined cleaner circuit utilising the same amount of flotation equipment specified in the

DFS design.

This provided a major improvement in the ratio of anticipated residence time to the nominated design

residence time and hence circuit stability. Simulations however failed to match those of the test work, with

any recovery benefit undermined by large decreases in concentrate grade caused by the copper oxide

flotation chemicals generating high levels of gangue entrainment.

A modified combined cleaner circuit was simulated with staged addition of rougher sulphide concentrate and

rougher oxide concentrate to the cleaner circuit. As shown in the Fig 14 this provided an increase in copper

recovery of approximately 2.5% when treating sulphide ore and little tangible benefit when treating oxide ore.

Considering the ore mineral split this should represent an overall improvement in copper recovery over the

mine life of 0.5%.

Fig 14. Simulated Grade-Recovery Curves for Various Circuit Design Options


0

5

10

15

20

25

30

35

0 10 20 30 40 50 60 70 80 90 100

Cumulative Copper Recovery (%)

Cum

ulat

ive

Con

cent

rate

Gra

de (

% C

u)

Sulphide Dominant Split Cleaners Sulphide Dominant Combined Cleaner 2 Oxide Dominant Split Cleaners

Oxide Dominant Combined Cleaners 2 SulphideDominant Combined Cleaners 1 In addition, the revised circuit required two less flotation cells and was expected to be overwhelmingly more

stable than the DFS split cleaner circuit design. The modified circuit was incorporated into the Andash detail

design. In addition the simulation mass balances were used to calculate cleaner circuit make up water

requirements. This resulted in all water addition piping to the cleaner circuit being changed in the detailed

design to further maximise circuit stability.

CONCLUSIONS A sequential flotation circuit with a standard sulphide flotation rougher, followed by an oxide flotation stage

rougher, using controlled potential sulphidisation was found to provide the best metallurgical performance for

Andash ore.

Locked cycle tests were of limited value due to the high variability in the resource. Geometallurgical

modelling using sequential copper assays and floatability component modelling provided a design

methodology with significantly less associated risk.

ACKNOWLEDGEMENTS The authors acknowledge the permission given by Kentor Gold Limited and AMEC Minproc Limited to

present this paper. Thanks is also given to Ian Walton who conducted many of the simulations critical to this

paper.

REFERENCES Alexander, D J and Morrison R D, 1998. Rapid estimation of floatability components in industrial flotation

plants, Minerals Engineering, Vol. 11, No. 2, pp. 133-143, 1998


Tessier, A, Campbell, P G C and Bisson, M, 1979. Sequential extraction procedure for the speciation of

particulate trace metals, Analytical Chemistry, Vol 51, No 7, pp. 844-851, June 1979

Wardell Armstrong International, 2007. Feasibility study for the Andash Cu-Au Project, Chapter 5 –

Metallurgical Test Work, WAI/61-0317, January 2007



Cost Effective Concentrator Design

G Lane1, P Dakin2 and D Elwin3

1. MAusIMM, General Manager Technical Solutions, Ausenco Minerals & Metals, 144 Montague Road, South Brisbane Qld 4101. Email: [email protected]

2. Principal Designer, Ausenco Minerals & Metals, 144 Montague Road, South Brisbane Qld 4101. Email: [email protected]

3. Principal Designer, Ausenco Minerals & Metals, 144 Montague Road, South Brisbane Qld 4101. Email: [email protected]

ABSTRACT This paper discusses the factors that contribute to the cost effective design of a concentrator.

Concentrator design and layout outcomes are functions of the team (engineer’s and owner’s)

participating in each particular project. The benchmark for relatively modest projects was set in the

1980s and 1990s during the ‘gold boom’ when numerous cost effective plants were designed and

constructed on a lump sum basis in a very competitive market. A number of factors contribute to lack

of transference of the lessons learnt in gold plant design to concentrator design including established

paradigms in the design and layout of concentrators, lack of experience in cost effective design,

operator’s preferences for flowsheet and layout and simple lack of appreciation of the impact of plant

footprint on materials quantities and resultant capital costs.

Experiences with recent copper concentrator projects (both small and large) are used as case studies.

INTRODUCTION Cost effective concentrator design is not an isolated paradigm. It needs to interface with the project

infrastructure constraints, owner’s needs, vendor’s capabilities, constructor’s logic and operator’s and

maintenance team’s preferences. However, cost effective design has some ‘rules of thumb’, namely:

Keep the execution strategy and plan simple, sift the ‘baggage’ from the facts early, have a

plan (and agreed scope) and stick to it,

Minimise the number of interfaces across all parties as every interface requires

‘management’,

Invest in good equipment as it saves you money, and

Reduce plant footprint as capital and operating costs increase with increasing plant footprint.

Capital costs will escalate if;



Scope is poorly defined and the execution strategy meanders (scope and design are not

frozen),

Simplicity is replaced with opportunism (hope),

Pipe rack locations are used as the basis of plant layout, and

‘Expandability’ is a necessity.

SETTING THE SCENE Nirvana for the project developer and plant designer is a ‘cost effective plant that works’ to

expectations. There are many approaches to plant design that range from the grandiose to the shoddy

and mean. ‘Cost effectiveness’ is a different paradigm. It relies on sound judgement and a balanced

assessment of what is required for the circumstance. The ‘art’ of designing cost effective CIL plants

was mastered in the late 80s by Australian engineering companies who competed on a lump sum turn

key basis for plants in the 100 kt/y to 5 Mt/y throughput range. The need for these plants arose

through the development of CIL technology, the plethora of modest grade opportunities in Western

Australia in particular, the high gold price and the flexible nature of gold metal marketing (Close,

2002). These circumstances resulted in the need for innovation, short project schedules, low project

capital costs and a ‘money making machine’ approach. On occasions the commercial and design

‘recipe’ came unstuck. Notable projects, such as Three Mile Hill, initially failed to meet throughput

targets due, in part, to a ‘one size fits all’ approach and insufficient test work to define the design.

The lessons learnt from the ‘cost effective’ design examples of the Australian gold industry in the 80s

and 90s are still the basis of sound CIL plant design although the frequency of projects in Australia is

now low and there is greater focus on Africa. The ‘art’ is practiced by some of the same individuals

and mostly in small to medium sized engineering companies. Interestingly, as the size of the

engineering company increases the ‘art’ is diluted by other commercial imperatives. Larger

engineering businesses target Tier 1 mining clients. These clients may express interest in cost

effective engineering but the demands of their business, the number of interfaces and the physical size

of the projects makes meeting the aspirations much more difficult. Tier 1 companies generally have

long life projects where the desire for durability, flexibility, life-cycle optimisation from the outset

override the urgency and design practicalities often associated with smaller projects.

‘Value engineering’ is often used to trim scope and cost but the fundamentals and paradigm of the

project are often immovable or have a large coefficient of restitution (or resistance) due to complex

standards, systems and approvals processes.



Cost effective design needs good engineers and designs and motivation from the client. For their six

large copper concentrators, Xstata has opted for a ‘standard concentrator’ approach as a method of

reducing design costs. Another project in South America opted for a ‘low cost’ design. However, the

design failed to meet many of the ‘rules of thumb’ and the plant is struggling to meet throughput.

Another copper concentrator currently in design has had an interesting history where the initial ‘high

value centre’ design failed to deliver a cost effective outcome. Subsequent external review raised a

number of pertinent issues and proposed an alternative design that went to the other extreme and was

too ‘low cost’ to meet process needs. The final outcome will be a balance between capital cost

aspirations and sound design.

KEY DRIVERS Typical concentrator layouts tend to be designed very conservatively because designers:

either don’t know how ‘close to the wind you can sail’ (fit for design) or aren’t sufficiently

experienced to understand the cost/risk/benefit relationships. This results in the additional

elevation of equipment based on false ‘standards’ or previous practices, e.g. concentrate

launder slopes, setting height of mills and flotation cells.

lack of innovation in design (as this requires experience and effective risk management). This

can include using the topography and gravity instead of pumps, not using pipe racks to

dominate layouts or building close-stacked vertical processes. Good examples are placing the

mixing tank on top of a storage tank or the MCC under the mill feed platform.

tend to ‘re-invent the wheel’ instead of re-using, improving or adapting previous proven

design. Existing design libraries aren’t well known about or published.

don’t have field installation, commissioning and site as-built experience, particularly on

projects they have designed. Designers with this experience are better equipped to understand

the basics of layout and operability and then implement these lessons in future designs.

come from other industries, e.g.petro-chem or oil & gas industries. As a result of their prior

history they have little or no relevant commissioning or operational experience relevant to

flexible but compact plant design.

Key drivers for a cost effective plant design are listed below.

Optimise the plant footprint with the aim of reducing concrete, structural steel, piping and

electrical/control cable and raceways. Push all areas close together, e.g. grinding close to

CIL/flotation, desorption and reagents close to CIL/flotation, air/water services close to plant,

big drives close to the switchroom.



Keep elevation of the equipment to a workable minimum. The elevation of run-of-mine bins,

mills, cyclones, flotation cells and thickeners are key drivers (Lane et al. 2005).

Do not use dominating piperacks or large platform areas.

Have common platforms, stairs, pipe and cable ladder supports.

Platework and lining should be kept lean, e.g. only put wear liners in the chute areas exposed

to wear and not all internals.

Design with a fit for purpose attitude as though it is your money you are spending.

Do not accept second best.

To progress design in a cost effective manner the following guidelines need to be applied (Lane and

Dickie, 2009).

The ore body and its mineralogy, geometallurgy and process responses need to be sufficiently

understood to allow process and market risks to be managed effectively.

Process flowsheets need to be ‘signed off’ in the first few weeks of design. Any changes

result in design change notices that cause rework and administrative churn within the project.

If the flowsheets cannot be ‘signed off, detailed design is not ready to commence.

Duty/standby equipment needs should be defined in the process flow sheets.

Process instrumentation and preliminary piping diagrams need to commence early and be

frozen from a scope perspective at the 40 per cent design complete stage.

Survey and geotechnical data generally hold up progress when finalising the location and

earthworks detail. Site survey and geotechnical studies need to be completed in the study and

front end engineering phases.

Client maintenance preferences and local crane availability impact on the decisions to use

overahead cranes, tower cranes, monorails or davit cranes for various duties. These decisions

need to be made early in the plant layout process to avoid rework in all disciplines.

Local weather or environmental issues may define the need for a plant under roof, inside

buildings or with other protection. Clear definition of environmental needs is required prior to

commencement of detailed layout design.



The cost of installing plant in buildings, particularly in the typical South American style, is

high. Clear definition of client preferences is needed prior to commencement of layout design.

In-country materials of construction costs need to be understood in order to make cost

effective structural decisions (e.g. concrete versus steel).

Project expansion requirements and timing need to be clearly defined in the FEED phase.

Concentrate transportation methods need to be defined in the FEED phase (e.g. truck, rail or

donkey).

Reagent delivery and on-site storage requirements need to be defined in the FEED phase

based on plant access limitations (seasonal weather and/or other social and environmental

factors).

The water balance needs to be finalized by area as the design is developed with particular

focus on the storage method (ponds versus tanks).

Environmental approvals need to be finalized and permitting requirements (traffic, run-off,

dust, noise, fumes, and materials safety) need to be clearly defined in the FEED phase.

From project management and execution perspectives the following issues need to be considered.

The ‘options’ need to be considered and evaluated prior to detailed engineering and project

execution proper. Value engineering assessments can occur during the design process but

these need to be limited to low level issues and not matters of scope or issues material to the

schedule. Value engineering exercises to contain capital cost after 30 per cent engineering

completion mean that the project was not set up initially with the correct capital and/or design

expectations,

Critical vendor equipment certified data needs to be expedited. Detailed design can continue

without vendor data if the team has sufficient experience to understand the impact that vendor

data can have on design. Vendor specifications by the engineering company may need to be

prescriptive to accelerate schedule. This compliments the use of good equipment as if this

equipment is similar to that used on other jobs the vendor information can be more easily

expedited.

Simplicity in approach is ‘king’. Packaging aspects of the engineering for completion by ‘low

cost engineering centres’ can be a recipe for disaster unless the packages are well defined and

managed. Engineering needs to be progressed to between 40 per cent and 70 per cent



complete prior to remote completion, and slightly less if key lead engineers migrate with the

packages.

The fabrication strategy needs to be developed cognisant of local (to the project) capability

and capacity, low cost off-shore alternatives and logistical issues such as the consolidation of

equipment and fabricated items. Pre-assembled modules may provide opportunities for some

locations where on site fabrication costs are high or people with the requisite skills are in

short supply. A logistics study is required at an early stage.

Steel sections standards differ between countries and need to be reconciled with fabricator’s

norms.

Project manager capability ‘to support the team to perform at maximum capability’ is a key

driver particularly in maintaining a high level of clarity from the client interface to the

drawing floor.

Engineering managers need to be able ‘to lock down the scope’, understand ‘fit for purpose

design’, assign responsibility, support the leads and motivate the team.

CASE STUDIES

Introduction

Project names are not used when discussing most of the case study examples below. Photographs

from other publications and the public domain are used as examples to illustrate particular design

features.

Large and small concentrators present different challenges. Small concentrators (less than say 10 Mt/y

and single concentrate) are simpler to arrange as there is typically one crusher, stockpile, SAG mill,

ball mill, flotation train and tailings thickener. Larger concentrators with multiple SAG and ball mills,

multiple flotation trains and large service runs demand an additional level of complexity for

maintenance access, service equipment and service runs.

Small Concentrators

In many respects, small concentrators of less than 10 Mt/y and particularly less than 5 Mt/y are easier

to layout in a cost effective manner provided that all contributors are of a like mind.

The following examples indicate what to avoid in order to optimise project value. Fig. 1 illustrates a

number of design features that increase plant footprint.



The space between the unit process operations are for mobile crane access and potential expansion

(regrind mills). Mobile crane access is most effective when there is no impact on plant arrangement.

If pipe racks need to be installed or extended to allow adequate access, the use of mobile cranes may

not be cost effective and alternatives such as tower or portal cranes should be considered.

The arrangement of the unit processes in Fig. 1 requires extensive pumping of slurry between unit

processes and the installation of large pipe rack ways. These can be avoided by thoughtful design in

most circumstances.

Fig. 1 - An example of a possible small concentrator layout

The derivation of the Fig.1 design is interesting in that the original DFS design and layout was a

typical open air design on relatively small footprint that had small pipe racks between facilities and

maintenance access by local davits and monorails. A metallurgical review with about 15 per cent of

the engineering completed resulted in the inclusion of flash flotation and this changed the cyclone

tower design substantially. In addition, it was decided to leave room for a possible regrind mill and

the milling facility and flotation circuit were separated to affect this change. Flash flotation cleaning

was added when engineering was about 25 per cent completed and installed in the location allowed

for the future regrind mill. A maintenance study was conducted at about 40 per cent engineering

complete and access for mobile cranes increased the separation between the unit processes and pipe

rack lengths increased. In addition, overhead gantry cranes were allowed over cyclones and the

primary crusher requiring significant structural strengthening due to the local seismic conditions.

Hence, this is a good example of a project where a design approach was not frozen and maintained

throughout the engineering design and where a series of relatively minor modifications led to a less

than optimum outcome due to their incremental impact on the layout.



Large Concentrators The design of large concentrators requires a large team of designers, often multiple parallel lines of

equipment and critical consideration of operating and maintenance activities due to the size of the

wear items, the weight of replacement equipment and the volume of consumables. In general, this

leads to a more complex decision making process and a greater tendency to conservatism. The

approach to design can also change to one that is driven by key engineers and designers to one that is

driven by a more over-arching approach based on ‘proven track record’ or prior designs.

At 12 Mt/y capacity with a single train SAG and ball mills in the grinding circuit it is relatively easy

to design a cost effective concentrator (Fig. 2). These projects generally require modest size teams and

can be lead effectively by a competent engineering manager using simple engineering systems. As

projects become larger, the team grows and the infrastructure and technical issues increase,

particularly when equipment selection considers large capities and/or novel design features. However,

it is still possible to design concentrators with twin train grinding circuits with up to 25 Mt/y capacity

with relatively simple and compact layouts (Fig. 3 for example) if the concept is set early, agreed by

the owner and conveyed effectively to the engineering group.

Fig. 2 - The 12 Mt/y concentrator (Lane et al., 2008)



Fig. 3 - The simple layout of a 25 Mt/y twin train SABC concentrator

There are numerous examples of different approaches to concentrator design throughout the world.

Plants with large capacities, such as at www.citicpacificmining.com, with multiple trains of the largest

grinding mills have large footprints. The grinding area layout is at the other end of the spectrum from

the concepts promoted in this paper. The mills are elevated and separated and as a consequence the

bulk materials quantities are high. This combined with location-related costs for materials leads to

high capital cost outcomes. There are reasons for the layout, particularly associated with the mill

erection process where the mills were assembled overseas, transported to site and lifted into place

using a purpose designed system, but the proportional costs of civil and structural works associated

with this style of plant design are significcantly greater than those for smaller plants such as in Fig. 2.

The South American market benchmark for large concentrators has been set by Bechtel, e.g. Los

Pelambres and La Candelaria. However, the style of these plants leads to high capital cost due to

relatively large footprints compared to Australian counterparts such as Cadia (Staples et al., 2008).

One of the more recent concentrators constructed in South America is illustrated at

www.amec.com/andacollo and is said to be an ‘innovative, low cost design due to the low grade of

the copper deposit’. However, there are aspects of the comminution circuit layout and maintenance

strategy that offer opportunities for further improvement in the context of cost effectiveness. For

example, the use of a Tower Crane may have allowed a significant reduction in structural steel in the

cyclone tower by allowing overhead cranes to be removed while the grinding floor layout is

‘relatively spacious’.

CONCLUSION

There is no panacea solution but there are some key issues to consider consider in the design and

layout of any plant as pointed out in this paper. However, the critical consideration is to give the



owner(s) what they want in meeting targets, budgets and project timing. To achieve this, it is the

engineer’s role to optimize the design within the owner’s constraints to achieve maximum value from

the project.

ACKNOWLEDGEMENTS The authors would like to acknowledge the following:

Eddie McLean for reviewing the paper and providing suggestions for improvement.

All those engineers and designers who have contributed to the authors experience over the

years, including those from Ausenco.

REFERENCES Close, S E, 2002. The Great Gold Renaissance (Surbiton and Associates: Melbourne).

Lane, G, Staples, P, Dickie, M, Fleay, J, 2008. Engineering Design of Concentrators in Australia,

Asia and Africa – What Drives the Capital Cost, Procemin 2008, Santiago, October

Lane, G, Dickie, M., 2009. What is Required for a Low Cost Project? Project Evaluation 2009

Conference, Melbourne, AusIMM, April.

Lane, G, Green, S, Brindley, S, McLeod, D, 2005. Design and Engineering of Flotation Circuits in

Australia, Centenary of Flotation Symposium, Brisbane, June.

Staples, P, Lane G, Messenger, P, 2008. Horses for Courses – Tailoring Front End Design to Project

Requirements, 40th Annual Meeting of the Canadian Mineral Processors, Ottawa, Ontario, Canada

January 22 to 24.



‘We’re Metallurgists, not Magicians!’

E McLean1

1. MAusIMM, Development Manager, Ausenco Minerals & Metals, 144 Montague Road, South Brisbane Qld 4101. Email: [email protected]

ABSTRACT

The decision making check-list and hierarchy of information for process and plant design is generally

wide-ranging and complex.

The nature of grades, mineralisation and ore types in an ore body, its mine development sequence, and

the body of results from metallurgical and physical test work present complex and competing

influences on flowsheet development, values for design, equipment selection, plant efficiencies and

operating strategy.

The challenge is how to manage all these variables, to understand their impacts on performance, and

how best to provide a cost-effective plant design and robust process that meets daily/weekly/monthly

operating objectives. Throughput, production and grade or product specifications need to be met. The

ore body is not average; its range of metallurgical and physical characteristics may be moderately to

highly variable.

This paper assesses some of the key competing influences and conflicting issues in design and

equipment selection, and their effect on production and performance expectations.

THE DESIGN SEQUENCE

Each item of process equipment, the flowsheet and process logic, and the operability of the treatment

plant is there by design because a sequence of process and engineering activities have lead to their

installation in project implementation and construction activities.

The development activities follow the staged sequences shown in Table 1. Although not exhaustive,

and without showing the complexity of linked stages and iterations for reviews and revisions, this

table simplifies a complex process and identifies the main process and engineering decision-making

activities.



Table 1 - Development sequence with principal process design and engineering activities

Stage Development activities Description of key activities in development stage I

- In

put

Geology, mineralisation, resource model

Discuss deposit(s) with exploration geologists

Review long and cross sections showing lithology and structures

Geological interpretation, reconcile metallurgical classification for zones or domains

Sample selection and sampling protocols

Representative samples and the basis for selection

Characterisation – metallurgical and physical, of the major ore types, ore zones or domains

Variability – metallurgical and physical, of parameters which influence design and production

Production composites, as required, for selected periods during the initial two to three years of operation

Test work, by laboratory and vendor

Preparation of samples, integrity of sample, prevent oxidation

Batch or semi-continuous, small or large scale weight for flowsheet development and to obtain data for process design

Reproducibility and robustness of tests; simulation of operating conditions; check/repeat and diagnostic tests

II -

Def

init

ion

Analysis of data

Collate and compare, interrogate for trends and relationships, investigate outlier data

Ore sequence in likely mine plan and weighting by ore characterisation

Processing strategy and process alternatives, flow sheet development

Value for design Recommend value/specification for unit process and for process stage

Define range of operating conditions and duty point

III

- E

ngin

eeri

ng

Calculation – size and select equipment

Text book or standard methodology, with efficiency factors

Include equipment factors – industry applicable, vendor recommended

Include operating factors – experience, reconciliation operation with previous design and performance

Include agreed or nominated engineering margin, owner margin, allowance for engineering risk

Tender and select for procurement

Vendor receives engineering specification, criteria, data sheet, drawings

Vendor uses proprietary modelling, calculations and data base to size, then vendor submits equipment which complies with throughput and duty required

Vendor selection also considers commercial aspects: performance warranties, penalties for non-performance, reputation, competition (cost)

New or second-hand equipment may be offered

Technical and commercial evaluation by Engineer: recommend to purchase

IV –

Del

iver

P

roje

ct

Start-up and ramp-up operations

Implementation plan to coordinate and schedule activities

Install and commission in accordance with standards and procedures

Maintenance, planned and in accordance with recommendations

Operator experience and training



Reviewers of the performance of project delivery through to plant start-up and operation (Agarwal et

al., 1983 and Tostengard, 2002), have identified fundamental issues which contributed to ‘financial

shock’ from attaining nameplate capacity at a slower rate that anticipated, or not achieving this at all.

They noted that, notwithstanding experienced and competent mining, design and planning personnel

on the project, the chief technical problems, start-up difficulties and poor performance for new

minerals processing operations occurred from the following items:

representativity - high heterogeneity of ore bodies - difficulties with adequate

sampling - testing of the ore body was inadequate (stage I in Table 1),

definition and process development - complexities of the process route - influence of

process water (stage II) ,

engineering - design deficiencies with mechanical equipment - equipment design

and/or installation is inappropriate (stage III), and

operation - operators lack knowledge and skills to properly operate the plant -

management is unable to cope with the problems experienced during initial operation

- maintenance is not executed properly, particularly during the initial operation

(stage IV).

From their research covering several projects, the authors stated that although problems with

mechanical equipment were expected, these risks can be reduced through efficient design and scale-

up. They attributed a number of these problems to basic design deficiencies such as incorrect

specifications, unit capacity and duty.

Although all items identified by the authors are integral to a successful project outcome, and critically

important in some cases (such as ore representativity and sampling), this paper reviews the following

key items (from Table 1 and as referred to by the authors) which have a high ability to influence the

final outcome:

analysis of the data to develop a viable processing strategy,

values for design to specify the duty for the unit process, and

calculations to size and select the process plant and equipment.

The sequence of activities in Table 1 corresponds to the schematic project flowchart (Lane et al.,

2008) in Fig. 1 which shows the relationship between inputs (green boxes = stage I) and outcomes,

engineering (blue boxes = stage III) and project execution (yellow boxes = stage IV). The project

design and definition stage (orange oval = stage II) is pivotal: all information must pass through this

activity and following interpretation and analysis, forms the basis for subsequent engineering design



and the planning for the execution phases of the project ensue. Activities and decisions in this stage

are able to impact on the project cost, schedule and plant performance.

Fig. 1 – Relationship between design inputs and project outcomes (from Lane et al., 2008)

THE DATA SEQUENCE

Practical and robust interrogation, interpretation and analysis of the data from the body of

metallurgical information available are fundamental to successful design. The quality of data at

various stages of project development is also of consequence with respect to its use and applicability

for design.

Samples and Data

Samples for metallurgical and physical testing are domain-orientated and can be categorised as

characterisation samples and variability samples.

Characterisation samples normally represent a lithology, oxidation, alteration, mineralogy, grade, or

spatial property of the ore, as appropriate, for which metallurgical or physical unit parameters are

determined by the test work for that specific ore classification. Variability samples are usually from

each domain or ore classification in which one property varies with, for example, metal grade, sulphur

grade, depth, location in the geological structure.



Data generated from tests on both these types of samples can be associated with blocks of ore in a

resource model; some of the more common applications are described as follows:

Hardness and competency values populate blocks of ore in which the dominant

lithology or alteration type is identified (characterisation).

Leach recovery values populate blocks of free-milling gold ore in which the extent of

oxidation is identified, i.e., for wholly oxidised, partially oxidised or transition, nil or

fresh ore (characterisation).

A relationship of recovery with a variable (or variables) can be descried by an

algorithm to populate blocks of ore as for example in gold ores, relationships between

leach recovery and feed grade, leach recovery and sulphur grade, and extent of

refractory gold and arsenic grade (variability).

Mill feed from a mine plan and ore schedule can then be interrogated to provide

information over a period (month, quarter, half-year) to support production and cash

flow forecasting.

Use of resource composites which are blended by ore types and grade to represent samples such as

whole-of-pit, life-of-mine and annual averages do not necessarily provide the basic information and

‘building blocks’ for flow sheet development and design that characterisation and variability

programmes do. Plant operation is unlikely, if ever, to treat such samples on a sustained basis.

Specific ore properties that may cause metallurgical and operational difficulties are diluted in a

composite sample and any impact these have on performance is dampened, or perhaps not identified

in the normal scatter of test results.

For any design based on average values plus allowances for mechanical and vendor margins, there is a

high likelihood that the expected performance and corresponding production will not meet planned

targets when treating ores over sustained operating periods with characteristics significantly above the

average. Furthermore, the opportunity to catch up production when ore with characteristic values

below average are treated may be constrained by physical and volumetric limits in the plant and

equipment; for example, pump capacity, launder size, screen area, thickener area, filter rate.

Data Analysis

In the early stages of a project development when drilling and geological interpretation are ongoing

and resource definition is in its formative stages, the body of test work data is limited. Data is

generally in small sets of information and the data selected is typically the maximum or the average of

the range. The outcome of the calculation is then tested by incremental increases to the data to assess



trends and to calculate sizes of plant and equipment. As well, benchmarking of the values used with

comparable industry data or in-house data from similar ore types and similar applications should be

done to improve confidence at this early stage.

As the project develops, not only does the volume of data available from several stages and iterations

of metallurgical test work programmes increase but so too does the complexity of information from

the amount of detail available from an improved understanding of ore characteristics and variability.

Superimposed on these are possible ore feed schedules to the plant with variables by tonnage, grade,

mineralisation and lithology, and ore exploitation sequence depending on the mining strategy, mine

development plan, pit financial modelling and resource constraints.

Analysis of the characterisation and variability data by ore type to determine values for design

includes:

A statistical analysis which identifies the average, the standard deviation, minimum

and maximum values, values at nominated percentiles (default is typically 75th or

80th) of the data population.

An assessment of outlier low/high data and include basis/comments, or discard with

justification.

A plot of values in an ordered sequence which graphically shows the distribution,

where the bulk of the values lie and where the outliers or sparsely populated values

lie.

Identify those values which characterise the major ore types, unit process

performance, process streams and water quality.

Determine a weighted value appropriate for principal blends of ore in the feed to the

plant.

Sources of Data

Input data and information for studies from concept to final feasibility, for basic engineer packages

(BEP) or front end engineering design (FEED) are from several sources. These can be grouped as

follows:

Project, operating and production parameters which are set by the client. These are

based on financial and organisation parameters. These are usually agreed and fixed,

and form part of the design basis, such as operating schedule and availabilities.



Test work (by laboratories, vendors, and consultants), resource models and

consultants’ reports. These are the main sources of data for the project as they contain

detailed test work and analyses for the resource, its geology, mine plan, all unit

processes, the water quality, and environmental controls. This body of work also

contains reports from client’s specialists, processing specialists, other consultants and

third-party audits.

Operating practice and industry standards. Sources of data are in-house experience,

operating history on similar treatment plants, published or confirmed operating

information.

Vendor data, handbook, regulatory and environmental standards, industry codes.

Sources of data are service and performance catalogues on like-for-like applications,

published engineering or discipline text books, published regulations.

Engineer’s data base, operating and commissioning experience. This would apply in

instances, typically as follows - if data from previous sources are not available - if the

number of test samples and testing is considered limited and/or incomplete - if the test

work does not replicate the unit process performance well enough - if test data is from

batch operation and needs adjustment for continuous basis - if scale-up factors need to

be allowed for based on previous experience.

As a guideline, the hierarchy of sources of information and data for process and plant design is usually

in the order as listed. The client parameters (first group) normally form the financial, operational and

production basis for the project. Test work data (second group) are ore and deposit specific and

consequently form the prime source of information for flow sheet development and unit process

design. Data from the remaining three groups of information would be used in preference where these

could be demonstrated to be more applicable and more representative than those in the first two

groups.

Design Value

The design value is the input value used in calculations to size and select the processing equipment

item (e.g. crusher, grinding mill, hydrocyclone, thickener, float cell, filter) or a unit process (e.g.

carbon adsorption circuit, counter current wash circuit). The value can be a grade or an ore/mineral

characteristic, an attainable unit parameter, a rate or capacity for the nominated duty and operating

condition for a process equipment or specific circuit.

The design value does not necessarily relate to production schedules or integrate to a mass balance.

The mass balance for the plant represents an operating condition for a continuous circuit in steady



state in which inputs equal outputs for all items (i.e., solids, water, metals, and elements) in the

balance. Plant mass balances are prepared for an average ore condition and for a series of cases which

represent specific feed parameters or operating conditions.

The design value does not include factors such as catch-up capacity, any additional design margins,

factors or other allowances. Design factors and margins should be assessed and identified in the

calculation where necessary to achieve the design objective, specific equipment duty or operating

condition. Factors include allowances for items such as wear, pegging in screens, minor surges,

emergency relief, certain start-up or shut-down conditions, short-circuiting. Operating margins are

allowed for items such as selected process equipment, conveyors, and cyclone feed pumps, other

slurry pumps, reagents pumps, water pumps and air services.

Equipment Sizing and Selection

Calculations for equipment sizing show the design value, the input flow(s) and additional factors

where these materially affect the equipment size and basis of selection. The input flow(s) for the

calculation are normally from a nominated (usually maximum) operating case or from a short-term

event which may include a step-change due to batch flow stream addition. The calculation using the

average mass balance condition is also normally carried out to provide a measure of the range of

operating performance for the unit.

Selection and recommendation of the unit process equipment or plant should meet two main criteria:

to function efficiently in accordance with the parameters and specifications nominated; and to be cost

effective from both initial capital and life-cycle assessments. Selecting the unit process equipment or

plant achieves the desired balance between competing demands of

(a) undersizing equipment to minimise costs, which may increase the risk of

underperformance, and

(b) oversizing equipment which, whilst providing a measure of confidence to meet the

operating duty, not only over-commits initial capital for the equipment item (or unit process)

but also can increase costs for concrete works and steel structure to install this item.

CASE STUDIES

Case studies for selection of design values for a number of key process criteria and for the design

basis for a typical gold plant are described in the following sections.



Selection of Design Values

Selection of a design value from their population of data is described for comminution, gravity and

feed grade cases. An example of variability data analysis for gold recovery with sulphur is given.

Each data set is from a different project, and provides an example of the basis used to select a value

for design.

Comminution – competency

The drop weight test ‘A’ and ‘b’ parameters are measures of the competency of an ore and are used in

grinding calculations for semi-autogenous and autogenous grinding mills. Fig. 2 shows the correlation

of Axb values with number of samples tested for three ore types in a supergene/hypogene type ore

body with skarn mineralisation.

0

20

40

60

80

100

120

140

160

180

200

220

240

260

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18

Axb

val

ues

Number of samples

Type1

Type2

Type3

25th percentile

Fig. 2 – Correlation Axb values with number of samples tested for various ore types

The Axb value selected for design from the distribution plots for each ore type is based on the 25th

percentile in each (lower Axb values represent more competent ore). The design value for each ore

type and the characteristic blend of ore types in the mine plan are shown in Table 2.

Table 2 - Distribution of ore types by the principal mine plan blends

Lithology Axb

(ore type) WI-ball kWh/t

Mine plan (wt%)years 1-3

Mine plan (wt%) years 4-8

Type 1 77 12.8 50 0

Type 2 76 11.5 20 20

Type 3 38 15.9 30 80

Total -- -- 100 100



The (weighted) Axb design value for each mine plan blend is 65 for the 1-3 year period and 46 for the

4-8 year period. The corresponding Bond ball mill work indices are 13.5 kWh/t and 15.0 kWh/t,

respectively.

The ore blend is less competent (higher Axb value) and less hard (lower work index) in the initial

three years, and consequently the unit power requirements for the SAG mill and ball mill are lower

during that period compared to the latter five years. Operating options are either

(a) ability to treat higher tonnages during the initial years based on the installed power and

nominated throughput for years 4-8, or

(b) to allow for expansion of facilities and increase grinding capacity to maintain the

nominated throughput for years 4-8.

Comminution – hardness

Standard Bond work index rod mill and ball mill tests are measures of the hardness of an ore and are

used in grinding calculations for rod and ball mills; they also form an integral component of SAG mill

grinding calculations. Fig. 3 shows the distribution of work indices with samples tested for a volcanic,

volcaniclastic, co-magmatic intrusive type ore body.

16

17

18

19

20

21

22

23

24

25

26

27

28

29

30

31

32

0 2 4 6 8 10 12 14 16 18 20 22 24 26 28

Wo

rk in

dex

(WI)

kW

h/t

Number of samples

Rod WI

Ball WI

75th percentile

Fig. 3 – Bond ball mill and rod mill work indices for samples from the dominant type

The work index values selected for design from the distribution plots, based on the 75th percentile, are

rod mill work index 28.9 kWh/t and ball mill work index 18.9 kWh/t.



Gravity recovery

Gravity recovery tests on ground samples (5 to 10 kg per sample) using laboratory scale centrifugal

bowls together with shaking tables or vanners measure the gravity recoverable gold from steams in a

ball mill circuit closed with cyclones. Fig. 4 shows the results from standard gravity recovery tests on

three ore types with auriferous quartz vein stockwork mineralisation.

0

10

20

30

40

50

60

70

0.2 0.4 0.6 0.8 1.0 1.2 1.4 1.6 1.8

Rec

ove

ry g

ravi

ty A

u%

Feed Au g/t

Type1

Type2

Type3

Type3A

Fig. 4 – Gravity recoverable gold from three ore types in quartz veined gold deposit

Gravity recovery is variable and no trend was observed by ore type or by feed grade. Outlier values

from the smallest volume type 3 ore (3A) were reviewed then discarded from the population for

evaluation. Recoveries between the 90th percentile and the maximum were then assessed, and as

there was very little difference in values in this range, 50 per cent gravity gold recovery (laboratory

conditions) was selected. The design value was based on this value, less a discount factor, to estimate

comparable performance in the plant. For a mass balance case for typical performance at steady state

operation, the average gravity gold recovery 38 per cent, less a discount factor for operations, was

used.

For downstream leach/CIL, the carbon circuit was designed to manage short-term stoppages in the

gravity gold circuit. This ore has low silver grades (less than 3 g/t Ag typically) so excessive silver

loading on carbon is not a consideration.

Feed grades for design

The cumulative distribution of copper feed grades from the mine plan and ore schedule for a long-life

copper-gold porphyry mine is shown in Fig. 5.



0.18

0.20

0.22

0.24

0.26

0.28

0.30

0.32

0.34

0.36

0.38

0.40

0.42

0.44

0.46

0.48

0.50

0 10 20 30 40 50 60 70 80 90 100

Cu

gra

de,

%

Cu

cumulative frequency %

LOM

years 1-3

years 4-12

75th percentile

Fig. 5 – Copper cumulative distribution from mine and ore schedule

The copper distribution in years 4-12 was selected as the basis for design for the flotation and

concentrate handling areas as the highest copper feed grades by period are encountered in the feed to

the concentrator. The 75th percentile was used for the design copper value. Gold grades are typically

higher in the initial three years than in the year 4-12 operating period. Gold has an economic

contribution, but no material influence on the sizing of the copper concentrator equipment.

Although copper feed grade was generally lower in the initial three years, it was preferred that higher

unit capacity is installed at the outset to avoid future disruptions to ongoing operations from

construction of additional flotation and concentrate handling facilities. Additional capacity was

therefore available in the initial years of operation for catch-up and ability to cope with any short-term

fluctuations.

Gold recovery variability

An apparent scatter of gold recovery results from standard grind and cyanide leach tests on gravity

tail, which were carried out on a routine basis for a large open pit gold resource, was resolved (Smith,

2005) by assessing the response by sulphur grade increments. The leach residue versus leach feed

grade trends for four sulphur increments are plotted, as shown in Fig. 6. Each sulphur grade

represents approximately equal proportions of the sulphur distribution in the mine model for the pit.



0.00

0.10

0.20

0.30

0.40

0.50

0.60

0.0 0.5 1.0 1.5 2.0 2.5 3.0

Res

idue

, A

u g

/t

Head, Au g/t

2.3% S

1.8% S

1.4% S

0.8% S

trend 2.3%S

trend 1.8%S

trend 1.4%S

trend 0.8%S

Fig. 6 – Gold residue and leach feed grade relationship by sulphur increments

A relationship of residue gold grade with feed grade for sulphur grade levels was obtained. Further

statistical analysis of this data base was required to develop a model which estimated the leach residue

gold grade for the range of sulphur grades expected in the ore schedule, and thus calculate gold

recovery from the corresponding feed gold grade.

Design Basis

A case study for the design basis for key areas in a conventional gold plant is described. The circuit

comprises the following key process stages: crushing; SABC grinding, gravity, leach/CIL, thickening,

cyanide detoxification, desorption, regeneration, gold room cathode and gravity preparation, smelting.

Plant feed characteristics

Feed ore to the mill comes from three open pits. Feed distribution by tonnage proportion and by grade

for the main ore schedule periods are shown in Table 3. The yearly intervals correspond to optimised

pit development to maximise gold and silver grade, and revenue. The majority of the ore comes from

Pit 1, with the initial ore from Pit 2 boosting gold and silver grades in the first two years. Pit 3 is the

lowest grade and is developed towards the end of the project life.



Table 3 - Mill feed from open pits, distribution by weight and average grades

Year Pit 1

(wt %) Pit 2

(wt %) Pit 3

(wt %) Total

(wt %) Grade

Au (g/t) Grade

Ag (g/t)

1 - 2 80 20 0 100 2.2 20

3 - 7 95 5 0 100 1.6 9

8 - 9 60 15 25 100 1.6 6

10 - 12 0 20 80 100 1.1 2

Characterisation – comminution and recovery

Ores from all pits were characterised by lithology and alteration types and tested for their

comminution competency, hardness, abrasion properties. Three categories were identified and

grouped as ‘hard’, ‘medium’, ‘soft’ according to UCS, Axb, rod mill work index and ball mill work

index values. Their distribution in the corresponding mill feed tonnages, by period, is shown in

Table 4. The corresponding recoveries of gold and silver, based on the characteristic gravity and

leach/CIL (on gravity tail) recovery performance by ore type for the blend of ores in each mill feed

category, are also shown in this table. The variability of feed grade in these periods was relatively

small and consequently the variability in gold and silver recoveries in these periods was, similarly,

minor. Comparative throughput and production performances are expressed as a throughput index

and a gold-equivalent index (converts silver value to equivalent gold) in the table.

Table 4 - Distribution ore competency/hardness and recoveries by year

Year Hard (%)

Medium (%)

Soft (%)

Total (%)

T’put index

Rec’y Au (%)

Rec’y Ag (%)

Au-equiv index

1 - 2 20 60 20 100 0.90 88 60 1.35

3 – 7 10 60 30 100 1.00 84 50 1.00

8 – 9 10 40 50 100 1.05 84 50 1.03

10 - 12 0 5 95 100 1.25 80 40 0.78

Grinding basis

A review of the ore hardness variability in the mine resource block model showed that a significant

number of blocks in the first two years of operation had a predicted work index above those typically

expected for the following five to seven years. The comminution design value adopted was based on

treatment of ores expected over the five year period for years 3-7 in the schedule. Consequently, the

capacity of the milling circuit was lower for the initial two year period when treating ores with

comminution indices above those in the 3-7 year period. The average throughput during this initial

period was limited to 90 per cent of the design throughput by the installed power of the selected mills.



However, as higher grades are associated with the harder ore, production for the first two years on an

equivalent recovered gold basis (Table 4) is 35 per cent higher than at the design throughput.

This strategy avoided over-capitalising the grinding circuit and underutilisation of grinding power in

later years (at the design throughput rate); it also prevented flow-on of higher capital costs if the

downstream process plant and equipment were increased in size to match the increased throughput

available with the additional grinding power.

Precious metal recovery basis

A carbon adsorption and desorption circuit could handle the duty required and was selected for this

operation. Although the throughput rate is lower in the initial two years, ore grades and recoveries are

highest and these have the main influence on solution and carbon inventories. The maximum gold and

silver inventories, and thus the carbon transfer, desorption and regeneration system to manage those

metal inventories occurs during the initial two years operation (Table 4). Consequently, this is the

basis for the carbon circuit design.

In gravity recovery tests on various ore types, silver recovery was consistently very low, typically less

than 5 per cent to a final gravity concentrate. Hence, for design purposes, it was assumed that all

silver in the feed was available for leaching and carbon adsorption. As the soluble ratio of silver to

gold for the design duty was about 6:1, silver management was the key determinant for carbon circuit

design and operating strategy.

The gravity recovery tests showed that this gold recovery was highly variable for all ore types and that

gold recovery was typically 10 per cent for ores from the (smaller) high grade pit and about 20 per

cent for the main pit. For design purposes, the impact of gold deportment to gravity ahead of the

leach/CIL circuit was therefore relatively minor in the initial years. It was assumed that the leach

circuit would treat all gold in feed. This ensured all gold was able to be recovered in leaching and

adsorption when the gravity circuit was not operating and also provided a small operating margin for

peaks in gold grades to leaching.

Design of the gold room and smelting activities was based on treating peak weekly inventories of

gravity recoverable gold and electrowinning cathode clean-up.

Volumetric capacity

The grinding circuit is able to treat 25 per cent more ore, classified as predominantly ‘soft’ and mainly

from Pit 3, during the latter stages of the mine life. Slurry viscosity and rheology characterisation

tests on moderate to high alteration ores that were in the soft category provided information on pulp

density, hydraulic gradients in the circuit and slurry transport properties. The proportion of soft ore



increased with mine life. This increasingly affected duty and specifications for equipment such as

slurry pumps, trash and safety screens, intertank screens, the thickener and for launders. Hence,

provision was made in the initial design to handle larger flows for periods when the slurry properties

would markedly decrease, or allowances made for additional, or upgrade of existing equipment.

CONCLUSIONS

The data from metallurgical and physical ore testing programmes is appropriate and applicable to the

extent that the sample selection is representative and correct preparation protocols are practiced. The

preferred methodology for sample selection as a basis for flowsheet development and process design

is by characterisation and variability.

Characterisation accounts for metallurgical and physical characteristics of the ore by domain, zone or

ore properties. Variability encompasses sub-sets of domains such as range of grades of economic

metal, range of grades of deleterious or penalty elements, ratio of grades of metals, spatial location

along strike and at depth.

Practical and robust interrogation, interpretation and analysis of the complexity of information and

detailed data available are fundamental to successful design. This is not possible without

consideration of possible ore feed schedules to the plant by tonnage, grade, mineralisation and

lithology, and of ore exploitation sequence including mining strategy, mine development plan, pit

financial modelling and resource constraints.

The selection of the appropriate design value for each unit process or process stage is critical.

Coupled with relevant design margins, factors, scale-up, benchmarking, operating experience and

input flow conditions, the unit process equipment or plant selected should meet two main criteria: to

function efficiently in accordance with the parameters and specifications nominated; and to be cost

effective from both initial capital and life-cycle assessments.

Events outside these parameters are magic, meaning the conjuring of tricks and illusions that make

apparently impossible things seem to happen whereas inexplicable things are special or mysterious or

are of inexplicable quality, talent, or skill - antonyms of science/metallurgy.

ACKNOWLEDGEMENTS

The inspiration for this paper comes from my beleaguered colleagues, many of whom have been in the

firing line from Project Managers and site Superintendents to explain “why” and “why not”.



REFERENCES

Agarwal, J C, Brown, S R, Katrak, S E, 1983. Taking the sting out of out of project start-up problems,

adapted from presentation at The American Mining Congress, September 1983, San Francisco.

Lane, G, Staples, P, Dickie, M, Fleay, J, 2008. Engineering design of concentrators in Australia, Asia

and Africa – what drives the capital cost?, in Proceedings Procemin 2008, October 2008, Santiago,

Chile.

Microsoft Word Dictionary – ‘magic’.

Smith, S J, 2005. Personal communication.

Tostengard, G R, 2002. Gold plant start-up performance – the financial impact [online], Performance Associates International, Inc. Available from: http://www.perfnet.com/GoldPlantPaper.htm [Accessed: 15 February 2002].



In-Pit Tailings – World’s Best Practice for Long-Term Management of Tailings

G M Mudd1, H D Smith2, G Kyle3 and A Thompson4

1. MAusIMM, Senior Lecturer/Course Director, Environmental Engineering, Monash University, Clayton Vic 3800. Email: [email protected]

2. Manager – Mining and Major Projects, Northern Land Council, Darwin NT 0800. Email: [email protected]

3. Mining and Environment Officer, Gundjeihmi Aboriginal Corporation, Jabiru NT 0886. Email: [email protected]

4. Mining Project Officer – West Arnhem, Northern Land Council, Darwin NT 0800. Email: [email protected]

ABSTRACT Mine tailings represent one of the two largest sources of mine waste from the minerals industries (the other

being waste rock). Given the problems of declining ore grades and growing production, tailings generation is

increasing exponentially across the global mining industry. Common practice is for mines to build large

containment dams to store tailings during operations, which are then rehabilitated following mine closure. At

some sites, however, the controversial practices of riverine or marine dumping of tailings is still practiced. As

recent tailings dam failures have shown, there are legitimate questions being raised about the long-term

viability of leaving tailings above ground due to the risks of collapse and failure. At present, it is rare for

mines to emplace tailings back into former mine voids following mine closure, most commonly into a former

open pit mine void (although a fraction of the tailings may be used for underground mine backfill, this is

minor in scale). The approach of in-pit tailings has numerous advantages, such as inherent physical stability,

low to negligible acid and metalliferous drainage (AMD) risks, as well as allowing more productive use of

formerly mined land. In tropical areas where surface repositories are prone to leakage with potential for AMD

and widespread dispersal of radionuclides (in the case of uranium mining), in-pit management provides

particularly strong advantages – thorough hydrogeological studies are required, however, to address residual

groundwater contamination risks and to establish sound, long-term environmental monitoring regimes for

groundwater quality and subsidence issues. This paper presents a study of in-pit tailings management and

compares it with surface repositories established in a tropical climate. The comparisons lead to the

conclusion that, at least under tropical conditions, in-pit tailings management should be considered as

world’s best practice – especially for radioactive uranium tailings.

INTRODUCTION The long-term management of mine tailings is a fundamental challenge for the global mining industry (IIED

and WBCSD, 2002; Spitz and Trudinger, 2008). The traditional approach has evolved from simple dumping

in surface piles or to rivers (or marine waters) to engineered dams for long-term storage. There remain a few

mines which use riverine (eg. Grasberg, Ok Tedi, Porgera, Tolukuma) or marine (Lihir, Simberi, Batu Hijau)

tailings disposal, while the vast majority of mines around the world use engineered tailings storage facilities

(TSFs). Commonly, TSFs use valley fill or ring-dyke designs to create the storage volume (Vick, 1990).



Over the past century, the combined trends of increasing production and declining ore grades are causing

tailings generation to increase exponentially across the global mining industry – leading to some TSFs now

reaching a capacity of the order of hundreds of millions (or even billions) of tonnes (eg. Bingham Canyon,

Escondida). This makes TSFs a major construction, operational and closure cost as well as giving rise to the

potential for significant environmental and public safety risks if TSFs catastrophically fail, as seen most

recently in Hungary while other major historical examples include Merriespruit and Bafokeng in South Africa

(Van Niekirk and Viljoen, 2005).

After mine closure, TSFs are commonly rehabilitated to regulatory requirements, such as removal of water,

placement of soil covers and revegetation, and this is expected to be sufficient for long-term environmental

protection. There are circumstances, however, when this approach is demonstrably inadequate to ensure a

high level of certainty of environmental and public health protection in the long-term – that is, ensuring no

impacts for time frames beyond a few decades. For example, two major issues are acid and metalliferous

drainage (AMD) risks from sulfidic mine waste (eg. Taylor and Pape, 2007) or erosion and exposure of

radioactive tailings from uranium mining. When these issues are compounded by an intense tropical climate

and a surrounding land use of high value conservation (eg. national park), it is clear that best practice tailings

management must evolve from the traditional paradigm of surface-based TSFs and rehabilitation.

An alternative approach to long-term tailings management which can demonstrably achieve low

environmental and public risks after rehabilitation is in-pit tailings. That is, after completion of mining, tailings

are emplaced in the former open pit void and rehabilitated. In-pit tailings removes the waste from above

ground and prone to long-term surface erosion of soil covers, as well as virtually eliminating seepage risks

easily reaching adjacent aquatic ecosystems. Since the tailings are most likely to be below the groundwater

table (following rebound after cessation of open pit mining), AMD risks are also effectively eliminated. The

two major issues which require careful investigation and assessment are the potential for groundwater

contamination and the extent of tailings consolidation and subsidence of the rehabilitated landform. Given

the large land footprint of TSFs and long-term risks such as erosion, a sound technical case can be made

that in-pit tailings clearly represents world’s best practice for tailings management.

This paper briefly reviews the common technical approaches to traditional and in-pit tailings, the relative

environmental advantages and disadvantages, as well as detailing a case study of the Ranger uranium

project which is required to use in-pit tailings and ensure a high level of environmental protection (since it is

on indigenous land in a wet-dry monsoonal climate and surrounded by Kakadu National Park, one of a very

small number of world heritage sites listed for both natural and cultural values). The paper therefore

demonstrates that in-pit tailings management is clearly world’s best practice, especially in circumstances

such as tropical climates, high value land uses and radioactive uranium tailings.

TAILINGS MANAGEMENT: A BRIEF REVIEW

The Traditional Approach Over the past century, the operational management and rehabilitation of mine tailings facilities has evolved

considerably. In the early days, tailings were often discharged to a nearby surface area or directly to



streams, with both approaches often leading to localized environmental impacts (severe in some cases). As

the scale of mines grew and recognition of the long-term extent of environmental impacts became better

understood, most mines moved to the use of engineered tailings storage facilities (TSFs), using either valley

fill or ring dyke-type dams to provide for long-term physical storage of the tailings.

There are three main engineering design approaches to TSFs – upstream, centerline and downstream,

named after their relative construction direction to the first starter dam, as shown in Fig. 1. In general,

upstream construction is quickest and cheapest but is more prone to various types of failures (especially

piping and seismic activity), while downstream is costly and time-consuming but is considered to be the

safest design approach under all types of hazardous events. The storage volume can be created through a

valley fill dam wall, or through a full ring dyke dam of four walls. All design scenarios require careful

consideration of operating conditions, especially discharge rates, water balance and long-term planning.

Further details of tailings management are given by Ritcey (1989), Vick (1990) and Minns (2007).

Fig. 1 – Upstream, centerline and downstream construction of tailings storage facilities (see Vick, 1990).

During operations, the water balance is one of the most crucial aspects to monitor and manage – since lack

of control of water is often a key pre-cursor to a TSF failure, through mechanisms such as overtopping,

piping or instability.

The traditional TSF is now a very familiar site across the global mining industry – with some now reaching

gargantuan proportions of billions of tonnes (eg. Bingham Canyon, USA, and tar sands operations in

Canada). Engineers are comfortable with their design, and in general, can design, construct, operate,

decommission and rehabilitate above-ground TSFs in a mostly straight forward manner (though long-term

studies on the effectiveness of TSF rehabilitation remain elusive). There are now TSFs built in high rainfall

areas (eg. Goro nickel project, New Caledonia) or high seismic risk zones (eg. numerous copper mines in

Chile)

An example of a rehabilitated valley fill TSF, the former Golden Cross gold mine in New Zealand and about a

decade old, is shown in Fig. 2. Visually it might appear to be a clean, successful example, but this does not

tell the full story – there have been issues with wall stability as well as minor ongoing AMD in TSF wall

seepage (as observed by the lead author in August 2009 during a site visit). The former mine site still

requires ongoing water treatment and active monitoring and maintenance. Despite the site being subject to

modern environmental regulation, the local community and regulators are still not comfortable with the

situation, as it increasingly appears that perpetual management will be required.

In the absence of a comprehensive audit of the environmental status of Australia’s hundreds (if not

thousands) of TSFs, only a cursory glimpse of issues at some TSFs can be described. At some sites, the

most common issue is wind-blown dust and associated hazards from heavy metals, especially in arid zones

(eg. in central Western Australia with concerns expressed over dust from the TSFs associated with the



SuperPit gold mine, see Cooke, 2004; or radioactive dust from the rehabilitated tailings dams at the former

Radium Hill uranium mine, see Lottermoser and Ashley, 2005), other TSF sites present ongoing AMD

pollution problems (eg. Mt Lyell; see Koehnken, 1997), while at others concerns relate to seepage and water

contamination problems. A less obvious issue is that of post-mining land use – TSFs require a large land

area, and depending on the project configuration and regulatory requirements, this land may be permanently

alienated from its former use, or it may be eventually amenable for use (depending on time taken to achieve

closure and rehabilitation, especially tailings consolidation).

FIG. 2 – Rehabilitated valley fill TSF at the former Golden Cross gold mine, southern Coromandel Peninsula,

New Zealand (mid-2009) (former open cut is in foreground, TSF in background left, water treatment plant centre

right).

The principal challenges are two-fold:

(i) long time frames – TSFs are in essence permanent solid and liquid waste storage facilities, and need to remain

effective in their engineered isolation role in perpetuity; and

(ii) land use changes – TSF land may never be put to its original use and becomes permanently alienated, or it

may retain certain restrictions (depending on project specifics etc).

Although the mining industry is achieving considerably improved technical and environmental standards (on

average) compared to 50 or even 25 years ago, there remains legitimate grounds for concern about the

nature of ongoing environmental risks from above-ground TSFs. A generalized representation of

environmental risks from TSFs is shown in Fig. 3.



In-Pit Tailings Management The discharge of tailings to a former open pit mine is a growing practice in the mining industry. There are a

variety of drivers, and these often vary considerably between different projects, but commonly include the

high capital and operating costs of TSFs, stringent environmental requirements, or simply lack of suitable

land for TSFs.

FIG. 3 – Common environmental risks from traditional TSFs (eg. radioactive uranium tailings).

In Australia, in-pit tailings has mainly been used in the gold and uranium sectors (for hard rock mines, soft

rock mines such as mineral sands or sometimes coal may also use variations of in-pit tailings), though most

sites are relatively modest in scale – only the Ranger uranium project is significant in magnitude. Some

examples include:

Nabarlek uranium mine (1979-1988): this small, high grade uranium deposit in Western Arnhem Land was

mined over a period of four months in 1979, with the ore stockpiled and processed from 1980 to 1988. Given

the availability of a worked out pit, tailings were directly deposited into the open pit, and the pit underwent full

closure and rehabilitation in 1995. Although the site is often claimed to be a successful example of modern

mining and environmental standards, there are important issues remaining, such as erosion of the surface cover

and exposure of underlying highly radioactive material as well as emerging concerns over possible groundwater

contamination (see Klessa, 2001; Bollhöfer and Ryan, 2007).

Kidston gold mine (1985-2001): this medium size gold project in northern Queensland initially operated a

conventional TSF but switched to deposition of thickened tailings and waste rock into the former Wises Hill pit in

1997, just before closure (Williams et al., 1999). Virtually no studies appear to have been conducted on the in-

pit tailings since closure, with most monitoring and research focusing on the above-ground TSF and waste rock

dumps (principally AMD risks).

Woodcutters lead-zinc-silver mine (1985-1999): this moderate mine south of Darwin (and just east of the former

Rum Jungle uranium-copper mine) had a small open pit and underground mine, and during rehabilitation

excavated approximately half the tailings and placed these in the former open pit below the water table to

minimise long-term AMD risks.

The most critical factors which give in-pit tailings strong advantages include minimal land disturbance

compared to an above-ground TSF, probable gravity pumping to the pit, low capital cost, probable lower

rehabilitation and ongoing management costs, and possible recovery of supernatant water (see Minns, 2007;

Spitz and Trudinger, 2008). The project sequence is obviously critical, as the timing of the exhaustion of a pit

and its preparation for receiving tailings will obviously influence the practicality of in-pit tailings deposition

(although returning tailings to a mined out pit as part of mine closure should be given serious consideration).



From an environmental perspective, in-pit tailings effectively eliminates long-term erosion concerns from

above-ground TSFs and the need for the land to be permanently alienated or restricted in access and use

after mine closure. The primary environmental challenges of in-pit tailings include slow consolidation rates

(which may even hamper rehabilitation schedules or designs), the higher volume of tailings compared to

original ore (ie. expansion due to blasting and grinding), and the remaining potential for groundwater

contamination. With respect to groundwater risks, the key issues are the relative hydraulic conductivities of

the tailings and surrounding rocks, as contrasted in Fig. 4.

FIG. 4 – Contrasting hydraulic conductivities (K) for in-pit tailings and surrounding rocks.

It is important to ensure that a thorough and full life-cycle analysis of costs, risks and benefits is completed

when comparing various tailings management scenarios. Overall, there are clear long-term environmental

advantages of in-pit tailings over conventional above-ground TSFs.

THE RANGER URANIUM PROJECT AND IN-PIT TAILINGS MANAGEMENT

The Ranger uranium project exists on indigenous freehold land and is surrounded by the world heritage-

listed Kakadu National Park (first proposed in 1964). The uranium deposits were discovered in October

1969, underwent extensive assessment and controversy in the 1970s and began production in August 1981.

The Ranger project has remained highly controversial throughout its history, with one of the key issues

remaining operational and long-term radioactive tailings management (amongst several other issues in

addition, such as uranium exports, water management, rehabilitation, and social impacts). The project was

subject to a major public inquiry, the Ranger Uranium Environmental Inquiry (‘REUI’, more commonly known

as the ‘Fox Inquiry’ after its Chief Commissioner Justice Russell Fox), which made numerous and strong

recommendations with regards to tailings management – the most critical being the need to emplace tailings

back in former pits following completion of mining and milling. This case study reviews the background to the

RUEI, why it strongly recommended in-pit tailings, the early years of tailings management for the Ranger

project, and recent progress and issues for in-pit tailings management at Ranger. The case study provides a

detailed assessment of the technical basis for in-pit tailings management as world’s best practice, especially

in tropical climates and radioactive tailings.

Brief Review of the Ranger Project and Tailings Management The Ranger uranium project is located in the Northern Territory, as shown in Fig. 5, and has a wet-dry

monsoonal climate. Following discovery in late 1969 and later detailed drilling, site investigations and mine

planning, the Ranger project was formally proposed in February 1974 through Australia’s first Environmental

Impact Statement (RUM, 1974). The large project envisaged at least two major open pit mines, developed in

sequence, a ring dyke TSF, waste rock dumps, conventional acid leach-solvent extraction mill, diesel power



station, and a nearby new mining town (to be called Jabiru). After completion of milling, the TSF was simply

planned to be directly revegetated on the tailings surface.

The climate is wet-dry monsoonal with an average of 1,552 mm of annual rainfall (varying from 1,032 to

2,623 mm) occurring almost exclusively during the summer months (December to March) and an annual

average pan evaporation of 2,588 mm with this rate being relatively constant throughout the year (Kabir,

2011). The wet season can deliver intense storms, related to tropical monsoonal troughs or cyclones,

including one event in late February 2007 which delivered some 750 mm of rainfall in less than 72 hours,

causing massive regional flooding and temporary closure of the Ranger site to address critical water

management issues.

Darwin

RumJungle Ja

bir

u

ArnhemLand

N

0 50 100 km

Alligator RiversRegion (ARR)

Ran

ger

Magela C

reek

East Alligator River

Pine Creek

So

uth

Alli

gat

or

Riv

er

N

Ranger Project Area

Kakadu National Park

Magela

Magela

Creek

Creek

‘009’

MCUS

Airport

Jabiru East

Mill

Pit 3

Pit 1

RP1

RP2

TailingsDam

MLA

A

0 1 km

Ponded Waters

Process Water

Stockpiles

WW

W

WW

W

W Wetland Filters

CoonjimbaCreek

Gu

lun

gu

l C

ree

k

CorridorCreek

FormerDjalkmarra

Creek

FIG. 5 – Location and site layout of the Ranger uranium project, Northern Territory.

The Ranger uranium deposits were on land which was being claimed by indigenous people, the Mirarr-

Gundjeihmi, as well as being proposed for a major national park in 1964 – and the inexorable conflict

between nuclear issues, land rights and conservation was set. By the time of the Ranger EIS in February

1974, the project was generating considerable controversy, which eventually led the Whitlam Government to

enact the Ranger Uranium Environmental Inquiry (RUEI) in July 1975 – with Justice Russell Fox from the

High Court, epidemiologist Professor Charles Kerr from the University of Sydney and civil engineer Graham

Kelleher from the public service appointed as inquiry commissioners. The RUEI was the first public inquiry in

Australia based on the Environment Protection (Impact of Proposals) Act 1974 (‘EPIP’), and led the

introduction and approach in this area (Mudd, 2008).



The RUEI released two major reports – the first on nuclear power in general in October 1976 (Fox et al.,

1976), and the second on the Ranger project and associated regional issues in detail in May 1977 (Fox et

al., 1977). The range of complex social, environmental, legal, political and economic issues were examined

in great detail by the RUEI, reflecting the deep concern held by local indigenous people and the broader

Australian community.

One of the most controversial issues was tailings management – especially Ranger’s proposal for direct

revegetation of the TSF after closure. The RUEI, after forensic scientific examination, fundamentally

challenged the proposed tailings management, and strongly recommended the adoption of in-pit tailings for

long-term closure and site rehabilitation – including that this outcome not be allowed to be altered during

operations. The two relevant recommendations were (Fox et al., 1977):

7.1 That the Ranger project be permitted to commence only if there is a firm, legally binding undertaking by Ranger

to replace in one or other of the pits the tailings and any stockpiles of low grade ore remaining after milling

ceases.

7.2 That the supervising authority not have the ability to relax the requirement that the tailings and unused ores be

returned to the pits.

The principal concerns which led RUEI-2 to this approach were radon gas emanation and radon progeny

buildup in the atmosphere (especially given the nearby indigenous communities), infiltration and seepage of

contaminated water through a conventional TSF reaching adjacent surface waters, and eventual erosion of

any soil covers (if they were used) (see Haylen, 1981 and Mudd, 2002). By comparison, the emplacement of

Ranger’s tailings into former open pits would virtually eliminate infiltration and seepage into surface waters,

and strongly reduce radon fluxes due to the tailings being below the groundwater table and covered with

waste rock and soil covers. In the view of RUEI-2, the environmental risks of a conventional above-ground

TSF, especially for uranium tailings in a region which was proposed as a major national park (later to be

world heritage listed), were severe when considering time frames of millennia – compared to much lower

overall risks of in-pit tailings.

The outcomes of RUEI-2 radically changed the approach to tailings management for the Ranger project.

When the project was formally approved by the Fraser Government in January 1979 (using the cold war-era

Atomic Energy Act 1953), a series of detailed Environmental Requirements (ERs) were prescribed. The ERs

mandated eventual in-pit tailings (clause 29a), but a sub-clause (29b) allowed at least a decade of research

to assess possible above-ground rehabilitation of the operational TSF to determine whether this could

achieve the same environmental outcomes. Thus the Ranger project was allowed a possible escape clause

from in-pit tailings – despite this being adamantly opposed by the RUEI. The approach being sought by the

Ranger project was to mine in pit #1, with all tailings going to a TSF, and then when mining began in pit #3,

initiate in-pit tailings for pit #1 and rehabilitate the TSF above ground after project completion.

The Ranger pit #1 was completed in December 1994 and was prepared for tailings deposition, which began

in August 1996. Full scale mining of pit #3 began in 1997 and is (currently) expected to cease in 2012, with a

possible conversion of underground mining after this. A possible heap leach expansion is also undergoing

formal environmental assessment and is about to begin public consultation.



Proposed TSF Rehabilitation and Related Tailings Studies Throughout the 1980s and into the 1990s a range of detailed research projects were undertaken to assess

the potential impacts of a rehabilitated TSF – including looking at the potential toxicity and geochemistry of

tailings in the Magela Creek floodplain downstream of Ranger, erosion field research and long-term

modelling at time scales up to 1,000 years, although virtually no research looked at seepage issues from the

TSF or potential groundwater problems from in-pit tailings in pit #1 (except the conceptual assessment of

Haylen, 1981). The extensive body of research was summarised by Wasson (1992), Waggit (1994), Waggitt

and Riley (1994) and Riley and Rippon (1997), as well as a raft of specialist reports and papers, providing

substantive weight to the arguments originally used by the RUEI Commissioners to justify in-pit tailings

management. There are many important points to note:

Perpetual maintenance would be required for an above ground dam (considered to be highly undesirable from

any perspective).

On the basis of even a small spill of tailings, say 2,000 t/year*, could lead to about 30 km2 of the Magela

floodplain being covered in 0.1 mm of radioactive tailings.

There is a paucity of data for the impact of released water and sediment constituents on relevant species, and

even less concerning toxicity assessment of the contained materials.

There can be no absolute assurance of ecological security because it is not possible to test the impact of

potential contaminants on all aspects of the environment, including the complex interactions among species.

Geomorphic modelling of one proposed above-ground rehabilitation structure at Ranger shows that there is

substantial erosion in its central area and on the margins of the steeper slopes – in some areas the predicted

erosion exceeds 7 m in depth after 1,000 years (proposed radon and infiltration limiting covers are often less

than this).

The fine-grained nature of the tailings facilitates rapid erosion and suspension in water, aiding longer travel

distances.

Assuming an erosion rate of only four times the natural rate, the proportion of tailings on the 30 km2 back water

floodplain would constitute 40% of the total sediment deposit.

A loss of 5,000 tonnes per year of tailings or eroded material from the containment structures would double the

sedimentation rate of the backwater floodplain immediately below Mudginberri Billabong.

The likely deposition sites for eroded tailings and waste rock include major stream courses, billabongs and 30

km2 of the backwater floodplain of the Magela Creek below Mudginberri.

Deposited tailings and waste rock fines may be subjected to a wide range of cyclical conditions, such as pH,

redox, thermocline development, floodplain and billabong hydrology, microbial activity, nutrient cycling -

especially if seasonal conditions are extreme.

Deposited materials can also undergo complex geochemical changes, with the potential to disrupt ecosystems

through effects on any number of species.

The volumes of particulates released from rehabilitated landforms will far exceed those that would have been

eroded from the lowlands prior to mining.

* Proposals for in situ rehabilitation of the above ground the tailings dam generally assumed about 10-17,000,000 t of contained tailings (eg. Armstrong, 1986; Wasson, 1992; Milnes, 1998), suggesting that a rate of 2,000 t/year of erosion would continue for thousands of years.



The tailings may also present subtle physical impacts at deposition sites, such as altering the infiltration

characteristics of soils, which in turn impact on water availability to plants and soil fauna, and this influences

plant growth and reproduction.

The physical impact of the eroding rehabilitated structure is likely to be high, leading to loss of habitat and

changes in the spatial distribution of ecosystems; infilling of wetlands and billabongs may be noticeable over

time.

Atmospheric and groundwater issues were not generally addressed, and

It is unlikely that a complete understanding of floodplain ecosystems will ever be achieved.

Current Status and Framework Although Ranger had hoped it could build a case for above ground rehabilitation of the TSF (and thereby

save significant rehabilitation costs) – expending considerable human and financial resources – it was

abundantly clear that this was not scientifically plausible nor realistic (ignoring other aspects such as social

issues and community expectations). In December 1997, Ranger finally decided to place all tailings in pits #1

and #3 and eventually remove and rehabilitate the TSF, thereby complying with ER 29a (Milnes, 1998; Li et

al., 2001) – as well as the spirit and intent of the RUEI recommendations. The Ranger project was

subsequently re-authorised in January 2000 for a further 21 years of operations, followed by a five year

period of rehabilitation. At this time a new set of ERs were developed and they now stipulate that “by the end

of operations all tailings must be placed in the mined out pits” (clause 11.2). Furthermore, two sub-clauses

specified that: (i) tailings are physically isolated from the environment for at least 10,000 years (11.3i); and

(ii) any contaminants arising from the tailings will not result in any detrimental environmental impacts for at

least 10,000 years (11.3ii). Thus, the statutory requirements for Ranger’s in-pit tailings are amongst the most

stringent in the world – and should clearly be considered as world’s best practice.

Deposition into pit #1 was completed in 2008, with deposition to the TSF recommencing in 2009 (following

additional lifts for greater storage volumes every year from 2006 to 2009). A recent aerial view of the Ranger

site is shown in Fig. 6, showing the current 1 km2 TSF, pit #1 full of tailings and process water, numerous

waste rock and ore stockpiles, water management features (RP1, RP2), still operating pit #3 and mill.

Ranger’s Tailings and AMD Risks During the RUEI, much of the debate about potential impacts from Ranger centred on concerns over the

heavy environmental impacts from the former Rum Jungle uranium-copper mine, just south of Darwin (see

Mudd and Patterson, 2010). The waste rock and tailings at Rum Jungle contained sulfide minerals, such as

pyrite, and these being exposed in the surface environment led to widespread pollution due to acid and

metalliferous drainage. During the RUEI, Ranger argued that its sulfide content was very much lower than

Rum Jungle and would be compensated by alkaline rocks which would negate any AMD.

Although the expectation of low AMD risks has proven reasonable, since no substantive AMD problems have

been reported at Ranger (though there are legitimate AMD concerns related to localized areas of the land

applications areas, better termed acid sulfate soils in this case), a completely unpredicted source of sulfides

emerged in the 1990s. The Ranger TSF contains the residual sulfuric acid used in processing the ore to

extract the uranium, as well as small concentrations of organic chemicals (ie. solvents) used in the chemical



purification of the uranium oxide. In these circumstances, the sulfate (from the sulfuric acid) can undergo

chemical reactions using the solvents for chemical energy to form sulfides – effectively the opposite reaction

of sulfide oxidation and AMD. The process of sulfide formation is strongly mediated by bacteria.

FIG. 6 – Recent aerial view of the Ranger uranium project (~late 2008; the TSF is ~1x1 km) (ERA, 2008).

A range of tailings research projects continued throughout the 1990s, especially with respect to exposing

recovered tailings samples from deep in the TSF to surface weathering, the results of which were reported

principally by ERA (2000), Milnes (2000) and Sinclair (2004). This collective work proved definitively that the

tailings showed sulfide formation and subsequent oxidation (ie. AMD) if exposed (pp 48, ERA, 2000). Thus, if

the tailings were left above ground in the TSF following rehabilitation, it would be highly likely that a

seasonally fluctuating water table would develop in the TSF, driven by the monsoonal wet-dry climate

(detailed groundwater-climate modelling is done by Kabir, 2011), thereby creating potentially ideal conditions

for some extent of AMD to develop.

It is perhaps no coincidence that this research work was also around the time that Ranger decided in

December 1997 to abandon their ambition for TSF rehabilitation and commit to long-term in-pit tailings

management. Given the intention to (eventually) incorporate the Ranger site into Kakadu National Park as

well as the reputation of the mining industry more generally, if AMD risks were to develop at a rehabilitated

Ranger site, this would be extremely problematic. By committing to in-pit tailings management, this will place

all tailings well below the groundwater table at Ranger and thereby virtually eliminate long-term AMD risks –

though the timing of the backfill of pit #3, the subsequent atmospheric exposure of the desaturated tailings



and the speed of groundwater rebound will be critical to assess to ensure no significant AMD issues arise

during filling and pit rehabilitation.

Overall, this outcome shows the critical importance of ongoing monitoring and assessment of tailings

management, but especially the need to explicitly factor in AMD risks when comparing above ground and in-

pit tailings management – including the potential for sulfide formation in TSFs.

CONCLUSIONS This paper has briefly reviewed the conventional approach to tailings management through traditional tailings

storage facilities. These solid and liquid waste structures are now very common across the global mining

industry (riverine and marine tailings dumping mines aside), and TSF engineers are familiar with their design,

construction, operation and rehabilitation. This approach, however, still leaves a significant long-term

environmental legacy, which could include alienated land, erosion, acid and metalliferous drainage risks, and

seepage and groundwater concerns. An increasingly popular approach is to consider in-pit tailings, which

effectively eliminates erosion concerns and significantly minimises AMD, seepage and groundwater risks in

the long-term. For radioactive tailings from uranium mining, the case for in-pit tailings management is even

more powerful since this also substantially reduces radiation risks – especially radiotoxic radon gas and its

progeny. At the Ranger uranium project, in the Northern Territory and surrounded by world heritage-listed

Kakadu National Park, the adoption of in-pit tailings was first recommended by the Ranger Uranium

Environmental Inquiry of 1977, although it was not mandated in statutory requirements until 2000. The

Ranger project is often described as the world’s most regulated mine – and, with regards to tailings

management, this is arguably true. This also helps to make the case abundantly clear as to why in-pit tailings

should be considered world’s best practice – not only for radioactive uranium tailings in the wet-dry tropics

but also mine tailings in general.

ACKNOWLEDGEMENTS This paper has evolved from the years of collective experience the authors have in examining the scientific

issues concerning in-pit or surface tailings management at various current and former mines in the tropics of

the Northern Territory. In particular, Gundjeihmi Aboriginal Corporation and the Northern Land Council are

acknowledged for their ongoing support of this field of work. Furthermore, various staff from Energy

Resources of Australia have been helpful over the years in addressing long-term tailings issues for the

Ranger uranium project (especially Alan Puhalovich for his efforts).

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pp 264-279 (Darwin, NT)

Bollhöfer, A and Ryan, B, 2007. Radiological impact assessment of the rehabilitated Nabarlek site, in eriss Research

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Fox, R W, Kelleher, G G and Kerr, C B, 1977, Ranger Uranium Environmental Inquiry – second report, 17 May 1977, 415

p (Australian Government: Canberra, ACT).

Haylen, M E, 1981. Uranium tailing disposal – Ranger project – a rationale, Masters Thesis, Macquarie University, North

Ryde, NSW.

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Kabir, M, 2011. Long-term impact study for climate change in the shallow unconfined groundwater recharge in Ranger

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Li, H, Cramb, G and Milnes, A R, 2001. Geotechnical characterisation of in-pit tailings at Ranger uranium mine, northern

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Chatswood, NSW, February 1974).

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Flotation Mechanism Design for Improved Metallurgical and Energy Performance

R Coleman1 and A Rinne2

1. MAusIMM, Manager – Process Equipment, Outotec Pty Ltd, 211 Montague Road, West End Qld 4101. Email: [email protected]

2. Technology Director – Flotation, Outotec Oy, Riihitontuntie 7, Espoo 02200, Finland. Email: [email protected]

ABSTRACT

The flotation mechanism plays a key role in the flotation process. It is used to provide the

contact between solid particles and bubbles and also the energy for these particles and

bubbles to attach. The mechanism must therefore provide sufficient mixing to keep the

solids in suspension so that they can be transported into the froth and recovered to the

concentrate. The design of the mechanism is critical and over the past decade Outotec

has developed and optimised a new mechanism design to improve metallurgical and

energy performance. The FloatForce® mechanism consists of a rotor and stator

arrangement and has been specifically designed to provide sustained mixing at higher air

flows and lower power inputs.

Over 500 FloatForce® mechanisms have been installed in existing and new operations.

This paper outlines the technological aspects of the FloatForce® mechanism and provides

metallurgical, operational and energy performance results from a number of flotation

circuits in which these have been installed.

INTRODUCTION

Flotation plant operators are continually looking at ways of improving metallurgical

performance and reducing operating costs. There are several options available to achieve

this, including improved reagent selection and additions, flowsheet optimisation through

modelling and simulation and the addition of new equipment that changes how the

particles behave (for example a regrind circuit). These are all driven by the operator. A

significant improvement to both metallurgical performance and operating costs can also be

achieved through an improvement to the design of the flotation equipment itself. So it is

also the responsibility of the flotation cell supplier to continually improve their products

through innovation.



The heart of the flotation cell is the mechanism. Over the past decade Outotec have been

developing, optimising and successfully installing a new flotation mechanism called the

FloatForce®. This new mechanism has seen improved metallurgical and energy

performance from the many sites in which it has been installed.

The technological aspects of the FloatForce® mechanism are described in this paper and

some case studies are provided from a number of flotation circuits in which these have

been installed.

ADVANCES IN FLOTATION MECHANISM DESIGN

Multimix Mechanism

After five years of rigorous design and experimental testing, the original Outotec Multimix

mechanism was developed in the 1970s by Dr. Kai Fallenius. The first model had a rotor

diameter of 750 mm with a suitably sized stator around the rotor. The air required for

flotation is forced down the lower shaft and introduced into the centre of the rotor. From

the centre of the rotor the air is directed through the six narrow rotor dispersion slots where

it is then contacted with the slurry outside the rotor. These highly turbulent collisions

provide the energy to form the bubble-particle aggregates. Once these aggregates are

formed they are then pumped through the stator which has full height stator blades.

The power draw and mixing decreases linearly with air addition as a consequence of the

location at which air is added. At high air addition rates this reduction in power in the

mixing zone can reduce solids suspension causing sanding.

Figure 1 shows the Outotec Multimix mechanism (rotor and stator).

Fig. 1 – Outotec Multimix rotor on LHS and rotor and stator on RHS



FloatForce® Mechanism

The development of a new mechanism design required significant laboratory-scale

development work and detailed CFD modelling on a laboratory and pilot-scale. Once the

CFD results were validated, the speed of the development of the new flotation mechanism

increased. As the basic design tests could be performed with CFD within a few days,

manufacturing of all new rotor designs was not required – neither was additional laboratory

or full-scale testing. Only promising rotor geometries were taken into physical tests. Finally

after several tests and design rounds, the new FloatForce® rotor was born. Slurry flow

through the new rotor simulated with Fluent CFD software is presented in Figure 2. The

new FloatForce® stator design was also developed during the development work of the

new rotor. As shown in Figure 2 the important area of stator for the flotation process in

Outotec’s mixing mechanism is the upper half. High wear parts can thereby be made

smaller and stronger and the whole stator becomes easier to handle and maintain.

Fig. 2 – Slurry flow through the FloatForce® rotor and stator as simulated by Fluent

CFD

The design specifications for the new rotor were to improve pumping and air dispersion

without affecting the many good features of the traditional OK-rotor. The new rotor also

needed to be mountable for existing installations without the need for larger motors or

additional features. The basic shape is close to the traditional and approved OK-rotor, but

the slurry and air flows inside the rotor differ significantly from the OK-rotor. Air is

distributed via six air ports that are connected through channels to the shaft. The air ports

discharge into six individual air dispersion slots on the outside of the rotor so that they lie

closer to the stator compared to the OK-rotor (Figure 3). Slurry can thereby fill the rotor



completely and pumping or power consumption does not drop significantly when air feed is

increased. Part of the slurry still flows through the new air slot and enables complete

mixing of air and slurry.

If existing Outotec flotation cell installations are retrofitted with the FloatForce®

mechanism, the power draw increases for the given aeration rate and motor size. The

benefits to this are two-fold in that firstly, the impeller speed of the retrofitted flotation cell

can be reduced, thereby saving power and secondly, a smaller motor may be selected for

any new installations, which also reduces energy consumption.

Fig. 3 – FloatForce® rotor

The principle of the new FloatForce® stator is modularity. Stator blades are manufactured

as separate parts that are mounted onto a stator stand/pedestal, as shown in Figure 4.

The pedestal is not a high wear part and should not require replacement for the life of the

project. The stator blades are supplied in halves, quarters or as single blades, depending

on the size of the flotation cell.



Fig. 4 – FloatForce® stator

The expected key features of the FloatForce® mechanism are shown in Table 1.

Table 1 – Expected key features of FloatForce® mechanism

Key feature Effect Result

Increased mixing at the same aeration rate

Increases bubble-particle collisions

Higher recovery

Increases the suspension of coarse particles

Higher coarse recovery

Coarser grind size (if liberation allows)

Enables the use of higher density slurry

Added solids residence time

Increased throughput

Pumps more slurry Less sanding

Maintained mixing at a higher dispersed aeration

rate

Increases bubble surface area flux Sb

Higher recovery

Maintained mixing and aeration rate (selection of

speed when specifying

Reduces power draw in no air / start-up situation

Smaller motor sizes



equipment) Lower capital expense

Lower spares cost

Modular stator wear parts Enables easier and safer stator maintenance

Faster maintenance

Less downtime

Over 500 FloatForce® mechanisms have been installed in existing and new operations and

all new Outotec flotation cells are now installed with the FloatForce® mechanism as

standard. These include all 80 TankCell-100 roughers at the BHP Billiton’s Escondida Los

Colorados Concentrator, the 27 TankCell-160 roughers at Anglo American / Xstrata

Copper’s Collahuasi Concentrator, the 10 TankCell-200 roughers at PanAust’s Phu Kham

Concentrator and the 6 TankCell-150 roughers at Oz Minerals Prominent Hill

Concentrator. The actual results of a number of other FloatForce® installations in terms of

metallurgical, energy or maintenance performance are presented in the following case

studies.

CASE STUDIES

Codelco Chuquicamata

The Codelco Norte Chuquicamata Mine is located in the north of Chile, 215 km north-east

of Antofagasta. Chuquicamata is one of the largest mining operations in the world and

primarily produces copper and molybdenum.

As part of the testing program of the TankCell®-300 installed at the Codelco Chuquicamata

Concentrator, three different hydrodynamic set-ups were studied in terms of metallurgical

performance and their impact on power efficiency. The TankCell®-300 was installed and

commissioned in October 2008 in Chuquicamata Concentrator A2, which consists of two

SAG mills and three rougher lines of eight TankCell®-160 flotation cells in a 2-2-2-2

arrangement installed in 2001. For comparison purposes, the TankCell®-300 was installed

as a first cell in one line, followed by the 160 m3 cells, in order to be compared with a

similar installed flotation volume of a parallel line. Detailed results of the comparative

testwork have been presented elsewhere (Elgueta et al., 2009, Grönstrand et al., 2009,

Yanez et al., 2009, Grönstrand et al., 2010 and Grönstrand et al., 2011).

The test work configuration and process data for this case study is presented in Table 2.



Table 2 – Test work configuration and process data

Parameter 1 2 3

Hydrodynamic set-up Base case:

FloatForce®

70 rpm

FloatForce® +

FlowBooster™

70 rpm

FloatForce® +

FlowBooster™

62 rpm

Feed rate (tph) 4001 4170 3720

Feed solids (%) 36.3 36.2 35.9

+212 micron 29.4 28.6 30.6

Three different hydrodynamic set-ups were tested – namely the base case and two

alternative set-ups. The base case had the standard hydrodynamic set-up of Outotec

flotation cells with a FloatForce® rotor-stator mechanism and it operates at 70 rpm, which

is the standard speed for this cell size. Second, the cell was fitted with the FloatForce®

mechanism and one auxiliary FlowBooster™ (Rinne and Peltola, 2008) impeller was

bolted to the shaft while maintaining speed. Finally, the FloatForce® and FlowBooster™

were kept, but the speed was reduced by 10%.

The average metallurgical performance and power consumption of the three hydrodynamic

set-ups tested are presented in Table 3. It is well known that the cost of electricity is the

most significant life cycle cost item in flotation operation with the energy cost up to two-

thirds of the total flotation cell costs over a lifespan of 25 years (Grönstrand et al., 2006) so

therefore this was measured during the test work.

Table 3 – Average metallurgical results and power consumption

Parameter 1 2 3

Copper grade in feed (%) 0.64 0.45 0.51

Copper grade in concentrate (%)

21.3 18.3 18.4

Copper recovery (%) 56.5 56.1 59.5

Enrichment ratio of copper 34.1 41.3 36.5



Agitator motor power draw (kW)

158 168 133

Blower motor power draw (kW)

43 42 42

Total power draw (kW) 201 210 175

Specific energy (kW/m3) 0.67 0.70 0.58

During the first test period the TankCell®-300 produced an average copper grade of

21.3%, a copper recovery of 56.5% at an enrichment ratio of 34. The specific energy was

0.67 kW/m3. The second period produced an average copper grade of 18.3%, a copper

recovery of 56.1% at an enrichment ratio of 41. The specific energy was 0.70 kW/m3,

which is higher than the first test period due to the addition of the FlowBoosterTM. The third

period produced an average copper grade of 18.4%, a copper recovery of 59.5% at an

enrichment ratio of 37. The specific energy was only 0.58 kW/m3,

These results clearly show that with the FloatForce® mechanism there is scope to reduce

the power input to the Outotec flotation cells without affecting flotation performance. In fact

in this case the average copper recovery increased.

Grupo Mexico La Caridad

La Caridad Concentrator Plant of the Mexicana de Cobre complex is located in the state of

Sonora, Mexico. The complex includes open pit mining, the concentrator plant, smelters

with flash furnace, and heap leaching with SX/EW. The mineralisation is mainly chalcocite,

covellite, chalcopyrite and molybdenite.

From 2005 onwards, the molybdenite recovery has been one of the main goals in terms of

value. La Caridad compiled a list of strategies in order to increase the copper and

molybdenite recoveries and one of the main actions was to review the general

maintenance of the flotation equipment on a regular basis. The rougher/scavenger flotation

circuit consists of 10 rows of 10 Outotec OK-38 U-shaped cells with four additional 28 m3

Wemco cells at the back end of each row (Figure 5). The Wemco cells were added after

the initial installation to increase residence time to assist with the recovery of the slow-

floating molybdenite particles.



Rougher Feed

10 rows of 10 x OK-38

Bank 1a

Bank 3a

Bank 7a

Tails

10 rows of 4 x 28 m3 WEMCO cells installed in PARALLEL configuration to

increase residence time per cell

Fig. 5 – La Caridad rougher/scavenger flotation circuit

The initial maintenance inspection of the flotation cells showed that all of the rotors and

stators in all rows were in a poor condition, with some of the rotors having rock

blockages. Outotec were then consulted about how to best improve the condition of the

mechanisms and a trial was set-up to determine the effect of the FloatForce® mechanism

on performance. FloatForce® mechanisms were then installed in Row 8 in Bank 3 and

Bank 7. Due to a limited budget, the worn Multimix mechanisms were kept in Bank 1.

New Multimix mechanisms were also installed in all cells in Row 9.

A total of six sampling audits were then performed by La Caridad´s Metallurgical

Laboratory from October 2008 to February 2009. Samples were taken around all of the

Outotec and Wemco cells in both Row 8 and Row 9. The average recovery results from

the six sampling audits are shown in Figure 6 for copper and Figure 7 for molybdenite. The

full results of the test work are presented by Gamez et al., (2011).

The copper recovery in Banks 3 and 7 of Row 8 showed an increase with the FloatForce®

mechanisms compared to the Multimix mechanisms in Row 9. The copper recovery in

Bank 1 of Row 8 (worn Multimix mechanisms) was approximately 7% lower than Bank 1

with new Multimix mechanisms, which clearly shows the importance of regular

maintenance on the flotation mechanisms. The overall copper recovery in Rows 8 and 9

were the same. It is important to note that the overall rougher copper concentrate grades

were also similar for both rows.



The molybdenite recovery in Banks 3 and 7 of Row 8 also showed an increase with the

FloatForce® mechanisms compared to the Multimix mechanisms in Row 9, with the

recovery in the Row 8 Bank 3 cells approximately 9% higher than in Row 9. Again, the

molybdenite recovery in Bank 1 of Row 8 was approximately 3% lower than Bank 1 with

new Multimix mechanisms. Even with the reduced recovery in Bank 1, the overall

molybdenite recovery in Row 8 was approximately 7% higher than in Row 9. The overall

rougher molybdenite concentrate grades were also similar for both rows.

0

10

20

30

40

50

60

70

80

90

100

Bank 1 Bank 3 Bank 7 Wemco TOTAL

Copper Recovery (%)

Row 8

Row 9

Fig. 6 – Copper recovery in La Caridad rougher circuit

0

10

20

30

40

50

60

70

80

90

Bank 1 Bank 3 Bank 7 Wemco TOTAL

Moly Recovery (%)

Row 8

Row 9

Fig. 7 – Molybdenite recovery in La Caridad rougher circuit



An analysis of the size-by-size recoveries for copper and molybdenite was performed to

further understand where the recovery differences arise from between the two

mechanisms. The size-by-size recovery results for the flotation cells with the FloatForce®

mechanisms in Row 8 and corresponding Multimix mechanisms in Row 9, for copper and

molybdenite, are presented in Figures 8 and 9, respectively.

0

10

20

30

40

50

60

70

80

90

100

0 50 100 150 200 250 300 350 400 450

Copper Recovery by Size (%)

Particle Size (micron)

Row 8

Row 9

Fig. 8 – Size-by-size copper recovery in Row 8 with FloatForce® and Row 9 with

Multimix mechanisms

0

10

20

30

40

50

60

70

80

90

100

0 50 100 150 200 250 300 350 400 450

Moly Recovery by Size (%)

Particle Size (micron)

Row 8

Row 9

Fig. 9 – Size-by-size molybdenite recovery in Row 8 with FloatForce® and Row 9

with Multimix mechanisms



The size-by-size results show that the FloatForce® mechanism increases the recovery

across all size classes.

From a hydrodynamic stand point, with the new mechanisms it was possible to increase

the superficial gas velocity (Jg) from 1.5 cm/s to 2.7 cm/s without any reduction in mixing

efficiency, thereby increasing the recovery. In addition the operators noticed a significant

improvement in froth quality in the banks with the new mechanisms and an improvement

to froth stability and control.

Based on the trial results La Caridad have decided to gradually replace the worn Multimix

mechanisms with FloatForce® mechanisms in all of the Outotec cells using a monthly

program for review and maintenance of the flotation cells.

Newcrest Telfer

Newcrest Mining’s Telfer Operation is located in the Pilbara Region of Western Australia

approximately 1300 km north-east of Perth, and produces primarily gold and copper. The

rougher circuit consists of 26 Outotec TankCell-150 flotation cells.

Due to the highly abrasive nature of the ore, the Multimix rotors and stators in these cells

were only lasting up to a year before requiring replacement. A number of FloatForce®

mechanisms were installed as a trial to determine the wear life and maintenance

requirements of the new mechanism compared to the Multimix. The trial FloatForce®

mechanisms were found to provide significantly better wear life, with the mechanisms

lasting more than two years. Based on the savings in spare parts, downtime and

maintenance costs, Telfer has begun the process of changing out the Multimix

mechanisms in all 26 flotation cells and replacing them with FloatForce® mechanisms.

Newcrest Cadia Valley Operations

Newcrest Mining’s Cadia Valley Operations are located approximately 25 km from the city

of Orange in New South Wales, and produces primarily gold and copper. The operations

consist of the Cadia Low Grade and Ridgeway High Grade Concentrators. The

rougher/scavenger circuit in the Cadia Concentrator consists of 14 Outotec TankCell-150

flotation cells in two parallel rows and in the Ridgeway Concentrator consists of 5 Outotec

TankCell-100 flotation cells.



A FloatForce® mechanism was initially trialled in the Cadia and Ridgeway scavenger

circuits, with the mechanism installed in the first scavenger in one row in Cadia and the

last scavenger in Ridgeway. A number of surveys were performed before and after the

installation to determine if the new mechanism improved cell performance. The full results

of the trials are presented by Cesnik (2009).

In the Cadia circuit, the FloatForce® mechanism produced a coarser concentrate than the

Multimix mechanism and a greater percentage of gold was present in the coarse size

fractions and less in the fine fractions. Overall, the coarse particle recovery increased after

the installation of the new mechanism. In the Ridgeway circuit, no significant difference in

performance was measured. As there is an inherent difficulty in the measurement of

scavenger tailings data, it was not possible to prove a statistically significant change.

Following these trials all of the rougher/scavengers mechanisms in both the Cadia and

Ridgeway circuits were changed to FloatForce® mechanisms.

In late 2010 a further test work program was carried out at Cadia to determine the effect

on metallurgical performance of reducing power in one of the TC-150 rougher cells with

the FloatForce® mechanism. Preliminary test work was first performed to determine the

range of operating conditions that could be used for the detailed test work. The preliminary

test work included measurements of solids suspension, gas dispersion characteristics

(superficial gas velocity, bubble size and bubble surface area flux) and metallurgical

surveys at the normal operating power input and at two lower power inputs. The power

inputs were varied using a variable-speed drive on the motor. The initial results indicated

that there was no significant difference in performance between the normal power input

and the lowest power input. The detailed test work will be completed in 2011 and results of

this test work will be presented at a later date.

CONCLUSIONS

With over 500 installations around the world, the new Outotec FloatForce® mechanism

has been proven to improve metallurgical performance and flotation circuit operation.

Operating costs can also be reduced through improved mechanism wear life and the

potential to reduce power input. As more and more data from case studies becomes

available, it is now obvious that significant improvements in flotation circuit operation can

be achieved through the use of the latest available technology.



ACKNOWLEDGEMENTS

The authors would like to thank Codelco, Mexicana de Cobre S.A. de C.V. and Newcrest

Mining for allowing permission to publish the results presented in the paper.

REFERENCES

Cesnik, F, 2009. Improvements in flotation cell operation and maintenance at Newcrest Cadia

Valley Operations, in Proceedings of the 10th Mill Operators’ Conference, pp 183-188

(Australasian Institute of Mining and Metallurgy: Adelaide).

Elgueta, H, Grau, R, Lamberg, P and Yanez, A, 2009. Statistical analysis of the operating

performance of Outotec’s TankCell-300 flotation machine, paper presented to Procemin

2009, Santiago, 2 – 4 December.

Gamez, A, Saltijeral, F, Lopez, O and Grönstrand, S, 2011. Respect equals recovery in

flotation, SME Annual Meeting 2011 (Society for Mining, Metallurgy and Exploration:

Denver).

Grönstrand, S, Niitti, T, Rinne, A and Turunen, J, 2006. Enhancement of flow dynamics of

existing flotation cells, in Proceedings of the 38th Annual Meeting of the Canadian

Mineral Processors, pp 403-422 (The Canadian Institute of Mining, Metallurgy and

Petroleum: Ottawa).

Grönstrand, S, Yañez, A, Morales, P and Elgueta, H, 2009. On hydrodynamic set-up of the

TankCell-300, paper presented to Flotation 09, Cape Town, 9 - 12 November.

Grönstrand, S, Yañez, A, Morales, P, Coddou, F, Elgueta, H and Perez, C, 2010. Operational

characteristics of the TankCell-300 at Codelco’s Chuquicamata Concentrator, in

Proceedings of the 42nd Annual Meeting of the Canadian Mineral Processors, pp 71 – 86

(The Canadian Institute of Mining, Metallurgy and Petroleum: Ottawa).

Grönstrand, S, Yañez, A, Wierink, G and Tiitinen, J, 2011. Cell power input or hydrodynamics

– which is more important in flotation, SME Annual Meeting 2011 (Society for Mining,

Metallurgy and Exploration: Denver).

Rinne, A and Peltola, A, 2008. On lifetime of flotation operations, Minerals Engineering, 21,

(12-14): 846-850.

Yañez, A, Morales, P, Coddou, F, Elgueta, H, Ortiz, JM, Perez, C, Cortes, G, Gomez, CO and

Finch, JA, 2009. Gas dispersion characterisation of TankCell-300 at Chuquicamata

Concentrator, in Proceedings of the 48th Annual Conference of Metallurgists, Sudbury.



The Great Oil Debate – Does Quenching Oil have a Deleterious Effect on Flotation?

C J Greet1 and J Kinal2

1. FAusIMM, Manager Minerals Processing Research, Magotteaux Australia Pty Ltd,

31 Cormack Road, Wingfield SA 5013. Email: [email protected] 2. Manager Metallurgy – Implementation, Magotteaux Australia Pty Ltd, 31 Cormack

Road, Wingfield SA 5013. Email: [email protected]

ABSTRACT Oil can enter a flotation circuit from a number of sources (for example, from mining equipment, oil leaks from the grinding equipment and coatings on the grinding media, but to name a few). Folklore generally states that any oil that contaminates the ore entering the flotation circuit is considered to have a negative impact on flotation. That is, a loss in concentrate grade and recovery. This represents a dilemma for some grinding media suppliers, who use a variety of quenching medium (air, polymer, oil or water) to achieve the desired physical properties of the ball. In the case of high chrome grinding media it is necessary, in some instances, to quench in oil to give the ball the right degree of toughness, which invariably leads to these balls having a thin film of oil on their surface. Once the grinding media arrives on site and site personnel recognise that the balls are “covered” in oil, they generally insist that the oil will have a deleterious impact on flotation performance. Obviously, any supplier “worth their salt” would take remedial action at the point of manufacture to minimise this problem. However, does quenching oil actually present a significant “threat” to the flotation process? This paper provides both laboratory and plant data describing the effect quenching oil has on the flotation of copper sulphide ore, with surprising results!

INTRODUCTION “History of Flotation” (Lynch, et al., 2010) provides an excellent insight to the development of the flotation process. Throughout this work there are references to the reagents used over the past 100 years or so, and it is interesting to note that from the beginning oils were used to enhance the flotation characteristics of the valuable minerals. For instance, the Potter process, de Bavary process, the Elmore process and processes developed by Minerals Separation all required the addition of oil to the ore to enhance the recovery of marmatite from the tailings dumps at Broken Hill during the early decades of the Twentieth Century. The rapid expansion of the flotation process into the recovery of copper sulphides was due in part to the use of oils as the flotation reagent. Oils remained important in flotation until xanthate collectors were developed in the mid-1920’s. With the advent of xanthate collectors the use of oil as a flotation reagent diminished rapidly. Other sulphydryl type collectors (for example, dithiocarbamates and dithiophosphates) have also gained acceptance in recent years, with their better selectivity against gangue species. Several of these compounds are oily, not soluble in water and tend to be added during grinding to so that they can be dispersed.



Perhaps of greater significance is the continued use of oil as a collector during copper/molybdenum flotation (Huiatt and Miller, 1972; Parlman and Bresson, 1980). These studies clearly show that the addition of oil during grinding has a positive impact on molybdenum recovery. It is thought that the oil increases the contact angle of the mineral, and therefore its hydrophobicity, which should improve the molybdenum recovery (Weiss, 1985). However, contemporary metallurgists do not appear to have this knowledge, and oil (in whatever form) added to a sulphide flotation circuit is considered to have a negative impact on grades and recoveries. This blind application of folk lore is not reasoned or reasonable. The literature tells us that small quantities of oil can actually enhance flotation.

LABORATORY EXPERIENCE The initial investigations into the effect quenching oil had on flotation performance commenced with laboratory studies examining dosage rate on copper rougher flotation under a standard test regime. Early estimates suggested that the quenching oil associated with the grinding media could equate to an addition of between 10 and 100 grams per tonne. The laboratory test programs tested a range, including these values. Quenching oil added during primary grinding In this work, quenching oil (0, 5, 10, 100, 1000 and 2500 grams per tonne) was added during grinding of a low grade porphyry copper ore in the Magotteaux Mill®. The ore was ground with a 15 percent chrome alloy to achieve a P80 of nominally 150 microns. The pulp chemistry of the laboratory mill discharge was recorded and appears in Table 1. With the exception of the 2500 g/t level, the addition of quenching oil resulted in a subtle decrease in the pulp chemical parameters measured. The reasons for this are not known, but may simply be the statistical noise associated with the test procedure as multiple tests for each condition were not performed. Table 1: The pulp chemistry of the Magotteaux Mill® discharge for copper rougher flotation tests completed on a low grade porphyry copper ore ground with 15 percent chrome grinding media at different concentrations of quenching oil added during grinding.

Oil addition, g/t pH Eh, mV (SHE) DO, ppm EDTA Fe, %

0

5

10

100

1000

2500

9.4

9.0

9.0

9.1

9.3

9.0

216

173

152

143

142

212

2.5

1.6

1.0

1.0

1.0

5.0

1.26

1.03

0.72

0.82

1.12

1.18

Copper rougher flotation tests were completed on the ground product. The copper grade/recovery curves for these tests are provided in Fig. 1. It is acknowledged that the copper grade/recovery curves for the tests completed at various concentrations of quenching oil are not identical but for



concentrations up to 100 grams per tonne the differences were subtle. At significantly higher additions the quenching oil appears to have a deleterious effect on copper flotation.

2

4

6

8

10

12

14

50 55 60 65 70 75 80 85 90 95 100

Cu Recovery, %

Cu

Gra

de

, %

0 g/t 5 g/t 10 g/t 100 g/t 1000 g/t 2500 g/t

Fig. 1: The copper grade/recovery curves for copper rougher flotation tests completed on a low grade porphyry copper ore ground with 15 percent chrome grinding media at different concentrations of quenching oil added during grinding. An examination of Fig. 1 suggested that at low dosages of quenching oil the copper concentrate grade, at the same copper recovery was decreased marginally (Table 2), but the flotation kinetics were increased (Table 3). However, at very high dosages both the copper concentrate grade and flotation kinetics were significantly retarded. Table 2: Copper concentrate grade and diluent recoveries at 80 percent copper recovery, for copper rougher flotation tests completed on a low grade porphyry copper ore ground with 15 percent chrome grinding media at different concentrations of quenching oil added during grinding.

Oil addition, g/t Cu grade, % Diluent recovery, %

Au IS NSG

0

5

10

100

1000

2500

8.5

8.4

7.7

8.9

5.1

3.8

78.8

81.2

81.7

82.5

77.1

88.7

69.1

68.3

68.6

62.6

52.1

56.0

0.6

0.7

0.9

0.6

1.9

2.7



These data suggest that quenching oil below 100 grams per tonne may in fact lead to an increase in the flotation rate of the various minerals within the pulp. This increased flotation rate may cause a marginal reduction in the concentrate grade as more gangue is recovered into the concentrate. However, the effect of quenching oil is considered to be subtle at these low dosages (less than 100 grams per tonne). As the quenching oil dosage increased there was a significant deterioration in the position of the copper grade/recovery curve. Not only was there a loss in concentrate grade, as the selectivity against non-sulphide gangue deteriorated, but the copper recovery also decreased. So, the effect of adding quenching oil during grinding will depend on the amount added. At low dosages, the effect is minor, and may actually have positive benefits, but at significantly higher dosages the effect is deleterious. Table 3: The flotation rate constant and maximum recovery values for copper rougher flotation tests completed on a low grade porphyry copper ore ground with 15 percent chrome grinding media at different concentrations of quenching oil added during grinding.

Oil addition, g/t k, min-1 Rmax, %

Cu Au IS NSG Cu Au IS NSG

0

5

10

100

1000

2500

3.43

3.47

3.62

3.67

2.73

2.86

3.63

3.87

3.81

4.28

3.11

2.84

2.77

2.89

2.91

2.24

0.96

1.87

0.15

0.23

0.30

0.35

0.14

0.25

86.9

88.0

88.8

87.7

79.0

76.0

85.7

87.5

89.4

87.9

76.1

85.3

78.1

78.0

80.1

76.3

56.0

50.2

2.6

2.1

2.7

2.6

3.6

4.1

Quenching oil added during regrinding grinding Again, quenching oil (0, 5, 10, 100, 1000 and 2500 grams per tonne) was added during regrinding of a copper rougher concentrate in the Magotteaux Mill®. The ore was ground with forged steel to achieve a P80 of nominally 63 microns. The pulp chemistry of the laboratory mill discharge was recorded and appears in Table 4. In general, the addition of quenching oil resulted in subtle variations in the pulp chemical parameters measured. The dissolved oxygen concentration in the slurry appeared to change independently of the quenching oil addition. These variations are suggested to be caused by a possible frothing effect of the quenching oil when air was added to the system. Again, copper rougher flotation tests were completed on the ground product. The copper grade/recovery curves for these tests are revealed in Fig. 2. The flotation behaviour in this case was a little different from that noted in the previous Section, in that the addition of quenching oil during regrinding resulted in an increase in the copper flotation kinetics (Table 5) for all dosages when compared with the standard test. However, it was again evident that increasing the dosage significantly did result in a deterioration in the concentrate grade as selectivity against the non-sulphide gangue decreased (Table 6). These data suggest that the addition of quenching oil lead to an increase in the flotation rate of the various minerals within the pulp, particularly molybdenum. For quenching oil dosages less than 100 grams per tonne and while the flotation rate increased, there was not a



significant reduction in the concentrate grade. However, at considerably higher quenching oil additions there was a noticeable deterioration in the position of the copper grade/recovery curve. That is, the selectivity against non-sulphide gangue deteriorated, resulting in a decrease in the copper concentrate grade. So, the effect of adding quenching oil during grinding will depend on the amount added. At low dosages, the effect is minor, and may actually have positive benefits, but at significantly higher dosages the effect is deleterious. Table 4: The pulp chemistry of the Magotteaux Mill® discharge for copper rougher flotation tests completed on a copper rougher concentrate ground with forged steel grinding media at different concentrations of quenching oil added during grinding.

Oil addition, g/t pH Eh, mV (SHE) DO, ppm EDTA Fe, %

0

5

10

100

1000

2500

10.2

9.9

10.7

10.1

10.5

10.5

180

190

190

190

190

180

2.1

1.0

6.0

5.2

4.8

2.5

0.83

0.82

0.88

0.76

0.79

0.63

Fig. 2: The copper grade/recovery curves for copper rougher flotation tests completed on a copper rougher concentrate ground with forged steel grinding media at different concentrations of quenching oil added during grinding. The improvement in the recovery of molybdenum is not surprising as oil is generally added during grinding to enhance its recovery in most copper/molybdenum ores.

4.00

6.00

8.00

10.00

12.00

14.00

16.00

35.00 40.00 45.00 50.00 55.00 60.00 65.00 70.00 75.00 80.00 85.00

Recovery - Cu (%)

Gra

de -

Cu

(%)

0 g/t 5 g/t 10 g/t 100 g/t 1000 g/t 2500 g/t



Table 5: The flotation rate constant and maximum recovery values for copper rougher flotation tests completed on a copper rougher concentrate ground with forged steel grinding media at different concentrations of quenching oil added during grinding.

Oil addition, g/t k, min-1 Rmax, %

Cu Mo IS NSG Cu Mo IS NSG

0

5

10

100

1000

2500

1.76

1.77

2.01

2.24

2.02

1.78

0.96

0.97

1.31

1.67

1.70

1.58

0.59

0.50

0.65

0.72

0.74

0.77

0.53

0.50

0.50

0.68

0.77

0.82

71.1

70.6

73.2

75.9

76.4

74.6

53.9

56.9

57.6

70.2

81.5

80.4

4.4

4.5

4.0

4.7

5.2

5.1

6.5

7.2

4.9

7.6

8.1

8.4

Table 6: Copper concentrate grade and diluent recoveries at 70 percent copper recovery, for copper rougher flotation tests completed on a copper rougher concentrate ground with forged steel grinding media at different concentrations of quenching oil added during grinding.

Oil addition, g/t Cu grade, % Diluent recovery, %

Mo IS NSG

0

5

10

100

1000

2500

9.9

9.8

11.5

10.3

9.5

9.0

50.5

54.1

52.3

62.3

73.9

75.2

3.6

3.6

3.0

3.2

3.6

3.8

5.2

5.8

3.3

5.1

5.7

6.5

PLANT EXPERIENCE

The laboratory studies suggest that at low dosages quenching oil may actually have a positive effect on copper flotation behaviour. But, how does one test this in a plant? The test program To obtain the desired mechanical properties the grinding media supplied to the site is quenched in oil, and as a consequence retains a residual amount of oil on the surface of the balls. The anecdotal evidence from the site was varied, with some observations suggesting there was no impact while others indicated the flotation circuit appears to get “faster” which changed the metallurgy, directly after the mill was charged with balls. However, no quantitative data existed describing the impact of the residual quenching oil on flotation performance. In order to gather some quantitative data arrangements were made to load the ball mill with five tonnes of grinding media during day shift so that the plant could be surveyed at least an hour before and up to at least two hours after charging to determine any changes in plant performance. “Spot” surveys of the concentrator were conducted at:



75 minutes prior to charging the mill, 30 minutes prior to charging the mill, While the mill was being charged, 30 minutes after charging the mill, 60 minutes after charging the mill, 110 minutes after charging the mill, and 140 minutes after charging the mill.

Samples of the following process streams were collected during each survey:

Rougher feed, Combined rougher concentrate, Rougher Cell 1 concentrate, Rougher Cells 2 and 3 concentrate, Rougher Cell 4, 5 and 6 concentrate, Rougher Cell 7, 8 and 9 concentrate, Rougher tailing, Cleaner feed, Final concentrate, and Final tailing.

The samples were weighed, filtered, dried, prepared and submitted for assay. The results were mass balanced using Algosys’ MATBAL (Version 8.0) software. In addition to collecting metallurgical surveys of the concentrator, samples of the flotation feed were collected, filtered and the solutions submitted for oil and grease analysis. The objective of this analysis was to determine if the concentration of oil and grease in the flotation feed increased after charging the mill with five tonnes of oily balls. Estimation of the quenching oil addition Work conducted in Magotteaux’s laboratory in Liege, Belgium (Bonnevie, 2009) estimated that the quenching oil retained on the surface of a ball equates to 0.4 percent of its mass. If this estimate is applied to the five tonnes of grinding media added to the ball mill it equates to the addition of 20 kilograms of quenching oil. At the time of surveying the plant, it was receiving nominally 1450 tonnes of fresh feed per hour. Therefore, the 20 kilograms of quenching oil is equivalent to 13.7 grams per tonne of fresh feed. This does not account for the circulating load, the residence time in the mill or the rate at which the grinding media was added to the ball mill (nominally five tonnes added over a 10 minute interval). Therefore, the instantaneous quenching oil addition may be higher than this, and it is expected that this would produce a spike in an oil and grease assay. The results In the first instance, the residual oil and grease assays were examined (Table 7), with the hypothesis being that the addition of media coated with oil to the ball mill should result in a spike in the oil and grease assay of the flotation feed, and with time this oil and grease assay should return to background levels. Table 7 shows that the residual oil and grease assay of the flotation feed prior to the addition of grinding media to the ball mill was nominally 15 mg/l. With the



addition of five tonnes of grinding media to the ball mill the oil and grease assay remained comparatively low (around 5 to 7 mg/l), until two hours after charging the mill when the assay jumped to 48 mg/l. Given that the residence time within the ball mill/cyclone circuit for an average particle is estimated to be less than 30 minutes the spike in the oil and grease assay two to three hours after charging the mill is most likely not related to the quenching oil, and probably comes from another source. While the data is limited, and further tests are warranted, the concentration of oil and grease in the flotation feed appear to be unaffected by the addition of five tonnes of grinding media to the ball mill. Table 7: Oil and grease assay of the flotation feed filtrate.

Time of sample collection Oil and grease concentration, mg/l

-60 to 0 minutes

0 to 60 minutes

60 to 120 minutes

120 to 180 minutes

15

9

5

48

The metallurgical surveys were mass balanced. The mass balanced head assays are provided in Table 8. While there was some variation in the head grade to flotation over the four hours that the plant was surveyed these differences were considered to be within the normal operating range. Table 8: The mass balanced head assays for metallurgical surveys conducted before and after charging the ball mill with five tonnes of grinding media.

Survey time Head assay

Au, ppm Cu, % As, ppm Mo, ppm IS, % NSG, %

-75 minutes

-30 minutes

0 minutes

+30 minutes

+60 minutes

+110 minutes

+140 minutes

0.57

0.65

0.65

0.56

0.62

0.59

0.54

0.9

1.2

1.1

1.0

1.0

1.0

0.9

320

330

317

296

281

257

254

279

294

296

248

260

243

247

7.5

8.6

8.4

7.9

7.6

7.6

7.6

90.0

87.9

88.5

89.4

89.4

89.7

89.8

The copper rougher grade/recovery curves for the seven surveys are displayed in Fig. 3. At first glance there does not appear to be any significant differences between the copper grade/recovery curves. However, on closer inspection the copper grade/recovery curves appear to improve with time. This is best illustrated in Fig. 4, where the copper rougher concentrate grade and recovery have been plotted against sampling time. Fig. 4 also includes the gold recovery. The data presented in Fig. 4 suggests that both the copper and gold recoveries were increasing prior to the addition of the five tonnes of grinding media. This may be related to the subtle increase in head grade (Table 8) during this time. Once the ball mill was charged both the copper and gold recoveries remained reasonably stable for the next two hours, averaging 91.7 and 72.3



percent, respectively. The copper rougher concentrate grade, with the exception of the first survey remained stable during the sampling campaign at around 13.4 percent copper. From these data, the addition of the five tonnes of grinding media, “laden” with quenching oil did not appear to have an adverse effect on copper or gold flotation. In fact, one could argue that the recoveries improved slightly.

8

12

16

20

24

28

20 30 40 50 60 70 80 90 100

Cu recovery, %

Cu

gra

de,

%

-75 minutes -30 minutes 0 minutes +30 minutes+60 minutes +110 minutes +140 minutes

Fig. 3: The copper grade/recovery curves for metallurgical surveys conducted before and after the addition of five tonnes of grinding media to the ball mill.

50

55

60

65

70

75

80

85

90

95

100

-150 -100 -50 0 50 100 150

Time, minutes

Cu

an

d A

u r

eco

very

, %

10.0

12.0

14.0

16.0

18.0

20.0C

u c

on

cen

trat

e g

rad

e, %

Cu recovery Au recovery Cu grade

Fig. 4: The copper and gold recoveries, and copper rougher concentrate grade versus sampling time for metallurgical surveys conducted before and after the addition of five tonnes of grinding media to the ball mill.



An examination of the flotation behaviour of the molybdenum, iron sulphide and non-sulphide gangue (Fig. 5) suggests that there were subtle variations in the recovery of both the iron sulphide and non-sulphide gangue recoveries to the copper rougher concentrate with the addition of five tonnes of grinding media to the ball mill. However, the molybdenum recovery appeared to be significantly enhanced (potentially a 5 percent increase in recovery) following the addition of grinding media to the ball mill. This is not surprising, considering the observations made above in the laboratory test work, and that it is common place to add small concentrations (nominally 5 grams per tonne) of waste oil to the grinding circuit to enhance molybdenum recovery in copper/molybdenum flotation plants.

50

55

60

65

70

75

80

-150 -100 -50 0 50 100 150

Time, minutes

Mo

rec

ove

ry,

%

0.0

1.5

3.0

4.5

6.0

7.5

9.0

IS a

nd

NS

G r

eco

very

, %

Mo recovery IS recovery NSG recovery

Fig. 5: The molybdenum, iron sulphide and non-sulphide gangue recoveries versus sampling time for metallurgical surveys conducted before and after the addition of five tonnes of grinding media to the ball mill.

CONCLUSIONS The literature indicates that oil particularly in low concentrations does not have an adverse impact on flotation, in fact oil is used to enhance the flotation of molybdenum. So, in a copper/molybdenum flotation circuit quenching oil may improve the recovery of molybdenum. The laboratory studies indicated that the addition of quenching oil up to 100 grams per tonne had little impact on flotation of copper sulphide ores. Beyond this there was a decrease in concentrate grade initially, followed by deterioration in the recovery. In the plant study it was estimated that the concentration of quenching oil added was around 13.7 grams per tonne. It was expected that the addition of five tonnes of grinding media



would produce a spike in the oil and grease assay of the flotation feed. The test results suggest otherwise. The metallurgical data indicated that the addition of five tonnes of oily balls did not have a deleterious impact on concentrate grade or recovery. In fact, the data suggests that the recoveries of both copper and molybdenum may actually improve.

ACKNOWLEDGEMENTS The authors gratefully acknowledge the support of Magotteaux in providing them with the opportunity to complete this work, and allowing it to be published.

REFERENCES Bonnevie M (2009), personal communication. Huiatt J L and Miller J D (1972), Flotation of Molybdenum and Copper Minerals with Chromatographic Fractions of Petroleum Distillate Products, presented at the AIME Annual Meeting (Preprint number 72-B-81), Society of Mining Engineers of AIME, New York. Lynch A J, Harbort G J and Nelson M G (2010), History of Flotation, The Spectrum Series Number 18, The Australasian Institute of Mining and Metallurgy: Melbourne. Parlman R M and Bresson C R (1980), New Collectors for Sulphide Minerals, Especially Copper and Molybdenum Based on Sulphur Chemicals Derived from Petroleum, in the Transactions of the SME 1980. Weiss N L (1985), SME Mineral Processing Handbook – Volume 1, Society of Mining Engineers of AIME, New York.



Carbon Prefloat Improvements at Century Mine

D G Rantucci1, T J Akroyd2 and L J Grattan3

1. MAusIMM, Project Metallurgist, MMG Century Mine, Lawn Hill via Mt Isa. Email: [email protected]

2. Metallurgy Superintendent, MMG Century Mine, Lawn Hill via Mt Isa. Email: [email protected]

3. Process Control Engineer, MMG Century Mine, Lawn Hill via Mt Isa. Email: [email protected]

ABSTRACT Organic carbon is a major ore contaminant at MMG’s Century mine. This carbonaceous matter is

associated with shale hosted sphalerite mineralisation. It occurs as elongated lamellae and as

inclusions in porous sphalerite aggregates which are fine grained and difficult to liberate from

sphalerite. The carbon prefloat circuit was designed to remove the liberated organic carbon in order to

mitigate excessive reagent consumption, contamination of the lead and zinc concentrates and poor

flotation performance in the primary zinc circuit. The consequence of organic carbon removal is that

sphalerite and galena minerals are lost to the prefloat concentrate. A Jameson Cell was installed in

late 2005 as a prefloat cleaner cell. Operational experience showed that the cell became capacity

limited during periods of high amounts of organic carbon in the feed and was subsequently bypassed.

Reactive style operating strategies were implemented to manage the effects of insufficient organic

carbon removal by the prefloat circuit but due to the relatively large circuit residence time, upstream

changes meant hours before the downstream circuits were stabilised following a disturbance from

excess carbon.

This paper describes the modifications and testwork undertaken to enable the Jameson Cell to be

operated during periods of high organic carbon in the float feed and the implementation of process

control initiatives to proactively remove the required prefloat mass from the float feed to eliminate the

reactive style operating strategy.

The process control modifications centred on dynamically targeting a prefloat mass recovery

determined by an equation relating the carbon to zinc ratio to the carbon prefloat mass as a percentage

of the float feed. The equation was developed by analysing historical data. These improvements

served to optimise the prefloat circuit and consequently stabilise the downstream lead and zinc

circuits.



INTRODUCTION MMG Century Mine is located 250 kilometers northwest of Mt Isa in Queensland Australia. It is the

largest zinc concentrate producer in Australia and second largest in the world, producing over 500 000

tonnes per annum of contained zinc.

The Century deposit is a stratiform shale-hosted zinc, lead and silver deposit. The sulphide

mineralisation occurs as fine-grained sphalerite, galena and pyrite. Sphalerite exists predominately in

two forms, with textures that may be described as normal sphalerite and porous sphalerite. Normal

sphalerite contains near stoichiometric amounts of zinc (limited substitution of cadmium and iron).

Porous sphalerite consists of variable amounts of fine intergrowths of quartz and carbonaceous matter.

The carbonaceous matter associated with sphalerite mineralisation occurs as elongated lamellae and as

inclusions in porous sphalerite aggregates (Waltho et al., 1993a). Quartz has also been identified as

inclusions in porous sphalerite (Waltho et al., 1993b). The carbonaceous inclusions within porous

sphalerite aggregates are nominally 2μm in diameter and are difficult to liberate. Ultrafine grinding to

a nominal P80 of 6.0μm typically liberates the majority of sphalerite from quartz and carbonaceous

material to enable production of a saleable product. However, as shown in Fig 1, porous sphalerite

particles with inclusions of quartz and carbonaceous material make it difficult to maximise sphalerite

liberation even at 6.0μm.

Fig 1: Optical image of porous sphalerite showing carbonaceous and quartz inclusions

The Century concentrator as it was in April 2008 is presented in Fig 2. The primary grinding circuit

consists of conventional SAG and Ball Mills which target a nominal flotation feed p80 of 55μm. The

carbon prefloat circuit represents the first flotation circuit following primary grinding. The carbon

prefloat circuit contains mechanical roughing cells and a Jameson Cell as a prefloat cleaner cell. Lead

and primary zinc circuits follow the prefloat circuit each with mechanical roughing, scavenging and



cleaning cells. Additionally, the primary zinc circuit contains a regrind circuit. The regrind circuit

size reduces the primary zinc scavenger concentrate to a nominal p80 of 16μm and feeds the first zinc

cleaner circuit. The primary zinc rougher concentrate and the first zinc cleaner concentrate both feed

the ultrafine grinding circuit. The ultrafine grinding circuit further reduces the size of the primary zinc

concentrate to a nominal p80 of 6μm followed by four stages of ultrafine flotation cleaning. The final

products are pumped 304 kilometers to the Karumba dewatering and load-out facility.

The prefloat circuit is designed to remove carbonaceous matter (here-in-after carbon) to final tail. Due

to the natural hydrophobicity of carbon and its affinity for organic reagents, it is important to remove it

prior to the lead and zinc circuits to manage excessive reagent consumption, reduce lead and zinc

concentrate contamination and poor flotation performance of the valuable mineral.

Fig 2: Century Concentrator Flowsheet, April 2008.

The effects of poor flotation performance due to excess carbon can be alleviated by increasing the

prefloat mass recovery but this results in increased zinc and lead loss. Typical prefloat mass

recoveries range between 3% and 6% of the float feed. This equates to an approximate zinc loss of

2.5% to 5.5%. Due to the proportional relationship of zinc loss and prefloat mass recovery, it is

necessary to minimise the prefloat mass recovery without detriment to the downstream circuits.



However, predicting when poor primary zinc circuit performance due to excess carbon would occur

was difficult and could only be managed by increasing the prefloat mass recovery target in response to

poor downstream flotation performance. The problem with this reactive style strategy was that the

residence time between the prefloat circuit and the primary zinc circuit is approximately one hour.

Therefore, the inherent lag extended the period of poor flotation performance in the primary zinc

circuit. The challenge was to understand the required prefloat mass recovery and how to proactively

remove sufficient carbon from the feed to prevent extended periods of poor flotation performance.

The solution came from identifying a relationship between the plant feed carbon to zinc ratio and

prefloat mass recovery.

The carbon to zinc ratio is a measurable quantity, which is an indicator of carbon throughput rate

relative to zinc throughput rate. Analysis of historical data identified the relationship between the

carbon to zinc ratio and prefloat mass recovery. The relationship was expected as the reactive style

strategy of increasing prefloat mass recovery during periods of poor flotation performance due to

excess carbon was a direct result of an increased carbon throughput rate, i.e. increased carbon to zinc

ratio. Therefore, it was proposed to design a controller that would dynamically target a prefloat mass

recovery based on a changing carbon to zinc ratio to manage poor flotation performance due to excess

carbon in the primary zinc circuit.

Supporting the development of a prefloat mass controller were proposals to improve the performance

of the Jameson Cell. The Jameson Cell was installed in closed circuit with the prefloat rougher circuit

in September 2005, as shown in Fig 2. Operational experience and laboratory testwork following

commissioning determined that the Jameson Cell was capacity limited at high carbon throughput rates.

The results also showed that flotation kinetics decreased as the recirculating load of slow floating

carbon particles increased between the Jameson Cell and the prefloat rougher circuit. The effect was

that target carbon grade could not be maintained for the same carbon recovery and the selectivity

between the valuable mineral (sphalerite and galena) and carbon deteriorated. Attempts to increase

the flotation kinetics of the slow floating carbon particles with reagents proved unsuccessful.

Therefore, the strategy for operating the circuit at high carbon throughput rates was to by-pass the cell

entirely thus eliminating the benefit it was designed to provide because sufficient prefloat mass

recovery could not be achieved from the cell.

To improve Jameson Cell utilisation and performance during high carbon throughput rates, it was

proposed to open circuiting the Jameson Cell tail and provide a controlled amount of prefloat rougher

concentrate to feed the cell. Providing a controlled flow of prefloat rougher concentrate to the

Jameson Cell would allow it to operate within its design parameters thus producing target carbon

grade and recovery and improved selectivity between zinc and carbon. Any remaining prefloat

rougher concentrate would be removed to final tail by an auxiliary pump thus alleviating the capacity



limitation of the Jameson Cell. The new prefloat final concentrate would therefore be a combination

of high-grade Jameson Cell concentrate and prefloat rougher concentrate. It was hypothesised that

there would be reduced zinc loss to prefloat final concentrate at high carbon throughput rate and

increased Jameson Cell utilisation by not being forced to by-pass the Jameson Cell during periods of

high carbon in the feed, because of poor Jameson Cell performance. Justification of the proposal

required testwork to determine if a reduction in zinc loss was achievable. The focus of the testwork

was on zinc only as lead contributes a small proportion of overall production.

This paper describes the development and implementation of process control initiatives that allow

dynamic targeting of a target prefloat mass recovery and improved Jameson Cell utilisation and

prefloat circuit performance during periods of high carbon throughput rates.

CONTROLLER MODIFICATION

Carbon to Zinc Ratio and Prefloat Mass Recovery Relationship To determine the relationship between carbon to zinc ratio and prefloat mass recovery, historical plant

data was analysed. Discrete carbon to zinc ratio groups were formulated to analyse the prefloat mass

recovery and combined prefloat and primary zinc circuit zinc losses. The combined zinc loss is an

important entity as optimum circuit performance relies on minimising the zinc loss to prefloat

concentrate and primary zinc tail. Determining the balance of zinc loss from the two circuits relies on

removing a sufficient amount of prefloat mass to prevent poor flotation performance down stream and

avoiding excessive removal of prefloat mass thus losing zinc to prefloat final concentrate. Therefore,

for a given carbon to zinc ratio, the lowest combined zinc loss between prefloat concentrate and

primary zinc tail represents the optimum prefloat mass recovery.

Prefloat mass recovery and combined prefloat and primary zinc loss were plotted for a range of carbon

to zinc ratio groups. Fig 3 describes data analysed for a particular carbon to zinc ratio group 0.295 to

0.305. As is shown from the plot, the optimum prefloat mass recovery was determined within a region

that constituted a minimum combined zinc loss. The region of minimum zinc loss is shown. The

variation in zinc loss may be attributed to a variety of factors such as mineralogy, reagent addition

rates, airflow rates and cell level.



Fig 3: Determination of Optimum Prefloat Mass Recovery for Carbon to Zinc Ratio Group 0.295 – 0.305

The previously described process was repeated for a range of discrete carbon to zinc ratio groups and

for each ratio, a prefloat mass recovery that minimised combined zinc loss was determined. Data was

then plotted on a graph of carbon to zinc ratio verses prefloat mass recovery as described by Fig 4.

The average carbon to zinc ratio for each group was used in the plot. This data set represents the

optimum prefloat mass recovery target for any given carbon to zinc ratio. A polynomial trend line was

established to allow the data set to be used in the plant control system. The lower and upper regions of

the plot contain limited historical data. Therefore, an upper and lower extrapolation was included to

ensure completeness through the carbon to zinc ratio range.

Fig 4: Carbon to Zinc Ratio and Prefloat Mass Recovery Relationship

y = 170.17x4 - 20.714x3 - 10.334x2 + 12.682x - 0.0112

R2 = 0.995

0.00

2.00

4.00

6.00

8.00

10.00

12.00

14.00

0.20 0.22 0.24 0.26 0.28 0.30 0.32 0.34 0.36 0.38 0.40 0.42 0.44 0.46 0.48 0.50 0.52 0.54

C:Zn Ratio

Pre

flo

at M

ass

Rec

ove

ry

Plant Data Lower Extrapolated Upper extrapolated Poly. (Plant Data)

C:Zn Ratio - Group 0.295 - 0.305

5.00

7.00

9.00

11.00

13.00

15.00

17.00

19.00

0.00 1.00 2.00 3.00 4.00 5.00 6.00 7.00 8.00

Prefloat Mass Recovery %

Co

mb

ined

Zn

Lo

ss %



Prefloat Mass Recovery Controller Design With the establishment of a relationship that allowed the determination of the optimum prefloat mass

recovery for a given carbon to zinc ratio, a circuit control strategy was developed. The prefloat mass

recovery controller is a standard PID controller that requires the target prefloat mass recovery as the

setpoint variable (SV) and actual prefloat mass recovery as the process variable (PV). Fig 5 describes

the prefloat mass controller within the prefloat circuit. The prefloat mass recovery SV is calculated by

way of the polynomial as described in Fig 4 where ‘x’ is the carbon to zinc ratio and ‘y’ is the prefloat

mass recovery SV. To ensure optimal plant performance, the carbon to zinc ratio calculation requires

accurate current zinc and carbon assay of the flotation feed. The zinc assay is provided by a

multistream analyser (MSA) and the carbon assay is provided by a three-hourly float feed grab sample

that is assayed by the laboratory and manually entered into the control system. The prefloat mass

recovery PV is calculated by online inputs of prefloat concentrate flow rate, prefloat concentrate

density and ore throughput rate. The prefloat mass recovery controller outputs the manipulated

variable (MV) as a remote set point for the prefloat air controller. The prefloat air controller then

adjusts the airflow rate to equalise the prefloat mass recovery PV with the prefloat mass recovery SV.

The prefloat air controller predominantly influences prefloat mass recovery by increasing or

decreasing the mass pull rate of the cell.

Fig 5: Prefloat Mass Recovery Controller Schematic



PLANT MODIFICATIONS

Experimental To create the correct testwork conditions, modification of the Jameson Cell tail discharge pipework

was required. A valved t-piece with two pneumatic knife-gate valves was engineered to replace a

spool piece on the Jameson Cell tail discharge pipe. This created the option to open circuit the

Jameson Cell tail, directing it to the floor during the trial and returning it to closed circuit once the trial

was complete. Jameson Cell tail discharged to the floor was subsequently sent to final tail. The

isolation of the Jameson Cell from the circuit during the trial eliminated the risk of destabilising other

parts of the flotation circuit.

The control system was also modified to enable the duty and the standby pumps on the prefloat

rougher concentrate collection sump to operate simultaneously. The role of the standby pump during

the trial was to remove the excess prefloat rougher concentrate that didn’t constitute Jameson Cell feed

to final tail and to maintain sump level. The duty pump would continue to supply fresh feed to the

Jameson Cell but at a constant flow rate.

The trial took place in December 2008. The operating parameters examined were:

Jameson Cell new feed flow rate (m3/hr),

Jameson Cell depth (cm) and

Jameson Cell feed density (%).

Data Analysis In order to minimise the impact on plant performance, the testwork on the Jameson Cell was

performed in isolation and therefore the benefit to the whole carbon prefloat circuit could not be

directly measured. Therefore, results from the trial work were used in conjunction with historical

plant data to predict the expected overall plant performance improvement. Fig 6 describes the

predicted zinc recovery improvement over historical performance. The data analysis predicted a zinc

loss reduction of approximately 0.5% if the Jameson Cell was operated in open circuit and at a

controlled new feed flow rate. The data is presented as an overall plant recovery improvement

accounting for the ensuing downstream zinc losses as a function of percent prefloat rougher

concentrate to Jameson Cell as new feed.



Fig 6: Predicted Overall Zinc Recovery Improvement

Jameson Cell Tail Reconfiguration Based on a successful outcome to the test program, permanent plant modifications were installed to

open circuit the Jameson Cell and re-direct Jameson Cell tail to the lead rougher circuit, as shown in

Fig 7. The bold line indicates the new Jameson Cell tail discharge pipe extending from the valved T-

piece. The original pipe to the head of the prefloat rougher circuit remains in place as an option if

required. The installation was completed and commissioned in August 2009.

Fig 7: Prefloat Circuit Overview – Post Prefloat Modifications, August 2009

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.00 25.00 50.00 75.00 100.00

Percent Rghr M ass Flow To Jam eson Cell (%)

Pre

dic

ted

Rec

ove

ry Im

pro

vem

ent

(%)



Plant Performance Improvements The prefloat mass recovery controller has been extremely successful in reducing the amount of time

dedicated to managing the prefloat circuit as shown by Fig 8. The prefloat mass recovery controller

was introduced in June 2009. Tuning was optimised throughout May 2009. The data between

October 2009 and December 2009 was not available due to concentrator downtime experienced

through this period. As can be clearly seen, the number of operator interactions in managing the

prefloat airflow rates has reduced on average by 84%. This indicates the operation of the prefloat

circuit is now more stable and there is less need to intervene to change the prefloat circuit’s operation

to correct poor flotation performance down stream.

Fig 8: Operator Interactions - Air Set-Point Changes, 2009 and 2010

The Jameson Cell modifications have improved Jameson Cell utilisation in periods of high carbon

throughput rate as shown in Table 1. The prefloat mass recovery limit when the Jameson Cell was by-

passed was 4.0%. As can be seen, Jameson Cell utilisation for periods of mass recovery above 4.0%

has improved from 77.9% to 89.8% or 11.9%. In addition, periods of prefloat mass recovery above

4.0% have increased from 19.4% to 32.0%. Therefore in retrospect, the 12.6% increase in prefloat

mass recoveries above 4.0% would have further deteriorated the utilisation of the Jameson Cell below

77.9% if the modification to the Jameson Cell tail were not performed.

0

500

1000

1500

2000

2500

3000

3500

Jan

2009

Feb 2

009

Mar

2009

Apr 2

009

May

2009

Jun

2009

Jul 2

009

Aug 2

009

Sep 2

009

Jan

2010

Feb 2

010

Mar

2010

Apr 2

010

May

2010

Jun

2010

Month

Op

erat

or

Inte

ract

ion

s -

Air

Set

po

int

Ch

ang

es



Table 1: Jameson cell utilisation for pre and post modification

The overall benefit of the Jameson Cell circuit modifications and the new prefloat mass control

strategy is shown by Fig 9. Plant data was analysed for the period April 2008 to December 2010, this

includes the period from April 2008 to July 2009, which was pre the plant and controller modifications

and August 2009 to December 2010, which was after the modifications were commissioned. It is clear

that there is a step change improvement in circuit performance following the implementation of the

plant and control modifications. Further analysis within the typical operating range of 4.0% - 5.0%

carbon mass recovery indicates that a performance improvement of approximately 0.5% was achieved

through the implementation of the new process control strategy and the redirection of the Jameson

Cell tail to the head of the lead circuit.

Fig 9: Prefloat Circuit Performance Benefit: May 2008 – December 2010 circuit.

CONCLUSIONS The highly complex nature of carbonaceous material in Century ore and its inter-relationship with

sphalerite have meant that managing and optimising the performance of the prefloat circuit is critical

to not only minimising the loss of zinc to the prefloat concentrate but also preventing poor

performance in down stream circuits. Historical reactive control strategies have lead to significant

0.00

0.50

1.00

1.50

2.00

2.50

3.00

3.50

4.00

4.50

5.00

5.50

6.00

2.00 2.50 3.00 3.50 4.00 4.50 5.00 5.50 6.00

Prefloat Mass Recovery %

Pre

flo

at Z

n L

oss

%

Pre-Modification - April 2008 - July 2009 Post-Modification - August 2009 - December 2010

Pre-Plant Modification Post-Plant ModificationMass Recovery > 4.0% 19.40% 32.00%

Jameson Cell Utilisation 77.90% 89.80%



periods of poor circuit performance. The introduction of the prefloat mass controller has stabilised

both the prefloat circuit and down stream circuits by proactively removing a controlled amount of

prefloat mass as determined by routine analysis of flotation feed composition (carbon and zinc

grades). Benefits such as reduced operator interactions in the prefloat circuit are attributed to the

success of the controller

A Jameson cell was installed as a prefloat rougher cleaning cell in 2005 to reduce the loss of zinc to

the prefloat concentrate, however it was determined that the cell became capacity limited and was

bypassed during periods of high carbon throughput. Test work was conducted and modifications made

to the circuit to redirect the Jameson Cell tail to the head of the lead circuit. This modification has

improved the utilisation of the Jameson Cell (from 77.9% to 89.8%) during periods of high carbon,

where previously the cell would have been bypassed.

A review of the overall carbon prefloat circuit performance before and after these modifications were

made has shown that an overall reduction in zinc losses of 0.5% has been achieved.

ACKNOWLEDGEMENTS The authors wish to gratefully acknowledge the operations and metallurgical personnel at Century

Mine who have assisted with surveys and have shown patience while the modifications were

implemented. We also wish to thank MMG for permission to publish this paper.

REFERENCES Waltho, A E, Andrews, S J, 1993a, The Century Zinc-Lead Deposit, Northwest Queensland, in

Proceedings The AusIMM Centenary Conference 1993, pp 52-54, (The Australasian Institute of

Mining and Metallurgy: Melbourne)

Waltho, A E, Allnutt, S L and Radojkovic, A M, 1993b, Geology of the Century Zinc Deposit,

Northwest Queensland, in Proceedings International Symposium – World Zinc ’93, pp120-122, (The

Australasian Institute of Mining and Metallurgy: Melbourne)



Quantifying Plant Flow Availability for an Alumina Refinery Expansion using Dynamic Simulation

B Reynolds1 and S Collins2

1. Manager, TSG, Level 11, 172 St Georges Terrace, Perth WA 6000. Email: [email protected]

2. Senior Consultant, TSG, Level 11, 172 St Georges Terrace, Perth WA 6000. Email: [email protected]

ABSTRACT The process design for green- or brown-fields metallurgical plants typically starts with two fundamental

elements: the detailed mass and energy balance and the plant flow availability. The detailed mass and

energy balance is generally developed using process flow-sheet simulation software and represents the

steady-state operation of the plant at normal operating conditions. The plant flow availability, also known as

a range of other industry-specific terms such as ‘flow availability’ and ‘operating factor’, describes the ratio

of plant production to the steady-state mass balance. These elements are then used together as the basis for

equipment selection and design.

Plant flow availability takes into account a myriad of production loss factors including planned and

breakdown maintenance, feed availability and composition changes, variability in operational performance

and the interaction of batch and continuous unit operations. The combined impact of these loss factors are

complex and as such plant flow availability is usually derived using a top-down approach such as

comparison to an existing operation, a group of similar operations or to an anecdotal ‘industry leading’

benchmark. These top-down approaches are fraught with risk however, as they cannot factor in changes to

equipment selection, flow-sheet or plant operating philosophy, and can result in over-expenditure of scarce

capital or shortfalls in plant production.

The use of dynamic simulation to provide a bottom-up alternative to quantifying the plant flow availability in

metallurgical plant design has been growing rapidly in recent years. It provides a technique to quantify the

combined impact of the loss factors described here and a tool to support process design and equipment

selection by facilitating a range of sensitivity and what-if analyses.

This paper presents a case study of how dynamic simulation has recently been used to support the process

design and equipment selection of a brown-fields alumina refinery expansion project.

INTRODUCTION The process design for green- or brown-fields metallurgical plants is typically based on two fundamental

elements: the detailed mass and energy balance and the plant flow availability.



The detailed mass and energy balance defines the operation of the plant under normal conditions. It is

typically developed using a static flow-sheet simulation package and forms the backbone of the process

design. The development of the detailed mass and energy balance for a complex metallurgical plant is a

process that can take months to complete and validate, and years to optimise and refine.

The plant flow availability defines the percentage of calendar time that the plant will operate at the normal

operating rate. While it is conceptually simple, in reality it is a complex factor to construct as it represents

the totality of the metallurgical plant design i.e. equipment selection, availability and reliability, dynamic

influences, performance variability and operability. In fact a fair proportion of the reduced availability is

wrapped up in only slightly reduced feed rates, which can make it difficult to predict accurately without a

detailed analysis.

Using these two elements of the design, the expected production from the plant is nominally the product of

the flow rate at normal conditions and the plant flow availability. Often for the earlier stages of design, the

required production is used along with an estimate of plant flow availability to calculate the required flow

under normal conditions, which then forms the basis for the process design.

While significant time and effort (rightly) goes into developing the detailed mass and energy balance and

using it to optimise the process design, the same level of rigour is rarely applied in the determination of the

plant flow availability. Quite often, high level estimates of plant flow availability are carried from initial

concept through to the feasibility and detailed stages of the design process. Indeed through most process

design phases in a project, the plant flow availability figure is the only element that is ‘sticky’ and that

doesn’t appear to change.

As the plant flow availability is contingent on the plant design, the lack of its refinement during the process

design can lead to a significant disjunct in the plant design. The later this disjunct is recognised, the more

expensive it is to fix. Even moderate changes to plant flow availability late in the project for complex, high

availability metallurgical plants typically have one of two outcomes i.e. the production and revenues fall

short or the design is forced to incorporate significant retroactive changes. Given the pressure of project

delivery timeframes and making the trade-off between the future cost of lost production and the immediate

cost of deferred production due to late project start-up, accommodating change to the plant flow availability

late in a project is not an enviable position for a project manager.

The key to mitigating this risk and improving the design process is to put more work into the estimation of

plant flow availability earlier in the design, and to incorporate its estimation into the iterative design process.

This will allow the implications of changes to the process design on plant flow availability to be evaluated,

and issues and opportunities to be identified early in the process, when there is still time available to rectify

or take advantage of them. Given the trend towards front-end loading of process design in complex

metallurgical plant projects, understanding how long the window for significant process design changes is



open for, and the implications of incorporating evaluation of plant flow availability later in the project, is

becoming crucial.

METHODS FOR DETERMINING PLANT FLOW AVAILABILITY In general there are three different methods available to determine plant flow availability. In increasing

order of detail, effort and accuracy these are:

Top-down estimate,

Bottom-up static estimate, and

Dynamic simulation.

Top-Down Estimate The top-down estimate involves benchmarking the project against an existing operation, a group of

operations or a nominal benchmark, such as ‘best in class’. While there can be varying levels of

sophistication involved, such as consideration of feed characteristics, geographical location or process

pathway, ultimately it is a technique that involves very little detailed consideration and almost always

becomes an aspirational design target rather than an achievable one or one that optimises the profitability of

the project.

The key limitation of the top-down estimate technique is that historical performance is not always a good

indicator of future performance. While minor changes to process design and plant capacity for brownfield

projects may result in similar performance and plant flow availability, for those projects where significant

changes to the location, feed composition, flow-sheet or operating mode have been made, this estimated or

experience-based assessment is not valid. This is particularly the case for brown-fields upgrades of complex

metallurgical plants, where the plant bottlenecks and operability often changes significantly, but there is a

strong tendency to assume that current plant flow availability or current operability scenarios will persist.

Top-down estimates are best used for concept level studies prior to the first serious stages of process design.

Bottom-Up Static Method While the bottom-up static estimate methodology encompasses what may seem to be a range of different

techniques, it is in reality different levels of detail premised upon the same assumptions.

The simplest technique for estimating plant flow availability using the bottom-up static method is to identify

or nominate the plant bottleneck, estimate the availability of the bottleneck itself, make some assumptions

about availability upstream and downstream of the bottleneck, and then estimate plant flow availability based

on this.



A more advanced technique is to use a reliability block modelling package to construct a reliability block

diagram at a unit operation or facility level of detail for the plant. This can then be used in conjunction with

the availability and reliability of the various components to estimate the plant flow availability.

As the bottom-up static method has essentially evolved as a tool to support the mechanical engineering

aspects of plant design, it can be a very useful tool for maintenance management and constructing bottom-up

estimates of facility reliability, but as a tool for determining process flow availability it falls short in its

ability to represent the process and its inherent operability and variability.

The first limitation of the bottom-up static method for determining plant flow availability is that when

availability events occur, such as failures or planned maintenance, assumptions are made about the ‘current

state’ of the plant and consequently the impact that the availability event will have. These assumptions may

or may not be true, depending on what else is happening in the plant at the time of the event, and the impact

of the event may not be consistent.

Some examples of factors that can influence issues of current state assumptions are the impact of

simultaneous availability events in the plant and the impact of surge capacity in the plant during an

availability event. During these events the bottom-up static method cannot factor in the current state of the

operations, such as tank levels and simultaneous failures, so an assumption is made about both. These

assumptions can lead the impact of the availability event to be understated or overstated, depending on the

current state, the specific events that occur, the process flow-sheet and the operability of the plant during

such events. A simple example of this is that when an upstream and downstream unit in a process are subject

to availability events, the impact will often be less if the events are simultaneous than if they went down at

separate times.

The other limitation of the bottom-up static method is that it is unable to factor in the impact of variability

due to the interaction of dynamic influences, or ‘plant dynamics’, in the plant. These factors include the

interaction between batch and semi-batch processes, variability in the process and the consequent shifting

bottlenecks and the impact of surge capacity, such as tanks and stockpiles, in the plant.

While bottom-up static methods are a better estimate than top-down estimates and can be very useful for

estimating reliability and analysing maintenance of facilities across the plant, it is important to recognise

their limitations, and implications for the accuracy of the plant flow availability estimate, particularly for

complex metallurgical plants.

Dynamic Simulation Method The dynamic simulation methodology involves a detailed time-series simulation of the movement of material

through the plant subject to the impacts of planned maintenance, unplanned outages, feed and performance

variability, batch and semi-batch processes, surge capacity and plant operability under specific conditions or



scenarios. It is a flexible methodology that can incorporate as little or as much detail as is required to

investigate performance. It is important to note that while dynamic simulation can have very intensive data

requirements, particularly as the design becomes more mature and the plant flow availability estimate

incorporates more operability logic, it is flexible enough to use from the concept stage through to plant

operations.

Unlike the other two methodologies, which estimate the plant flow availability directly, dynamic simulation

calculates plant throughput by simulating the full plant operations, which can then be used to back-calculate

the plant flow availability. While it is possible to capture all of the detail of the mass and energy balance in a

dynamic simulation, the inclusion of that level of detail often limits the flexibility and speed of use of the

dynamic simulation. Often, these complexities can be investigated in the detailed steady-state mass and

energy balance and then simplified for use in dynamic simulation by empirically fitting process relationship

where appropriate, for example a relationship between flow and yield for precipitation in the Bayer process.

The empirical fit maintains the speed advantages of reducing the depth of the simulation, but still captures

the process relationship adequately for the purposes of determining plant flow availability.

The use of dynamic simulation to estimate plant flow availability and the use of that figure in conjunction

with the detailed steady-state mass and energy balance is an effective and powerful way to use these

techniques. Of the two approaches discussed here, discrete-event simulation (DES) has been used in the case

study presented in this paper and is the method recommended.

Discrete-Event Simulation Discrete-event simulation (DES) is a technique that is typically used in the mineral processing industries to

model complex supply chains. Its strengths lie in the ability to represent batch and continuous systems

subject to the influence of randomness. An example applied to the design of bulk export supply chains is

provided by Glassock and Hoare, 1995. It has also been used successfully as a technique for modelling

process plant operations, for example work by Koenig et al., 2002.

The use of DES over other continuous plant modelling tools has four key benefits as follows:

1. It allows for inclusion of both batch and continuous operations, allowing impacts at the interface to

be directly captured.

2. DES has at its core a Monte Carlo simulation engine which allows for the incorporation of

randomness and its effects on the system.

3. It is a highly customisable and flexible modelling environment which allows for bespoke

representations of unit operations and processes.



4. The models can run quite fast relative to flow-sheet simulation packages running in dynamic

simulation mode. This allows for a fast response to questions asked during the discovery phase and

better integration of the tool into the design process.

Flow-sheet Simulators using Dynamic Mode Another approach that is available is to use the flow-simulation tools used to develop the detailed mass and

energy balance running in ‘dynamic mode’. Indeed the Holy Grail often described for this approach is to

have a single model that represents the full detailed plant mass and energy balance and to be able to run it in

both dynamic and static modes, depending on the task at hand.

While this may seem ideal from the perspective of having a single reconciled tool that can be used across a

design team, the reality is that dynamic simulation and mass and energy balances often require different

levels of detail and assumptions to be used effectively to support process design. Assumptions that can be

made when estimating plant flow availability, such as simplifying the energy balance, minor streams or plant

chemistry, are part of the core purpose of the detailed mass and energy balance. Similarly, assumptions

about equipment selection, operability, reliability and variability are not directly relevant to the ‘operations at

normal conditions’, the steady-state design, but make a significant difference to the plant flow availability.

Some of the disadvantages of using a flow-sheet simulator in dynamic mode to estimate plant flow

availability are that the flow-sheet simulation packages may have limitations in their customisability or

representation of variability, that they may not be able to properly represent batch or semi-batch processes,

and that the additional detail in the model required to support dynamic simulation may significantly increase

simulation run time to the point where this tool cannot be used effectively to provide decision support.

It is important to note however that the disadvantages described here for using flow-sheet simulators in

dynamic mode for determining plant flow availability do not pose a problem for other uses of these tools

such as detailed control system design and for the development of operator training packages. These

applications benefit from a high level of detail and exact correspondence between the simulation and plant

equipment.

CASE STUDY: ALUMINA REFINERY EXPANSION To support a major expansion project, a dynamic simulation model of an alumina refinery was developed

using DES. The model was developed to confirm plant flow availability, a key assumption underpinning the

project evaluation and design.

DES has been used extensively across a range of applications in the mining industry. These applications

include below-ground and above-ground mining operations, preparation plants, processing plants, railways,

bulk terminals and port facilities. Coupled with the right study process, it has proven to be a highly flexible

and valuable tool.



To facilitate the study, the team-based phased approach to simulation projects discussed by Hoare, 2007, was

used. This approach advocates a four phased approach for simulation projects namely definition,

development, discovery and reporting.

EXPANSION PROJECT The expansion project was a significant upgrade to the mining operations, alumina refinery and export

facilities. Support was also provided for the design of the bauxite and alumina transport operations using

DES.

Early in the refinery upgrade design process, a need was identified to confirm the overall plant flow

availability that had been a key assumption used in project evaluation and design. Refinery capacity and

availability would need to be examined in a whole-of-refinery context, to understand plant dynamic effects

and the implications of this for equipment requirements i.e. number of units, operating rate and spares.

Alumina Refinery Model The alumina refinery model was developed using the commercially available Rockwell Arena 11 DES

software. The model boundaries were drawn from the bauxite feed into the mills through to the delivery of

hydrate from the product filters to calcination. All of the significant units along the main alumina flow and

spent liquor loop were included, with simplifying assumptions made where appropriate. A block flow

diagram of the system is shown in Fig. 1.



Milling

Desilication

Digestion

Clarification and CCD

Polishing Filtration

Precipitation

Seed Preparation

Product Filtration

Residue Disposal

Residue to Disposal

Hydrate toCalcination

BauxiteFeed

Spent LiquorRecycle

Fig. 1: Alumina Refinery Simplified Block Flow Diagram

The variable inputs into the model included the equipment configuration, equipment performance and

process relationships. Equipment configuration included the number of pieces of equipment, equipment

spares and tank volumes and was developed from existing configuration and design estimates. Equipment

performance included the operating rates and variability, cycle times and equipment reliability, developed

from detailed analysis of historical data. Relationships between process streams were developed using

outputs from the design mass balance model in SysCAD and were used to specify chemistry and flow

relationships, such as the utilities demand and the water balance across a number of units.

As the purpose of the model was to evaluate flow availability over a long time period, the energy balance

was not considered and there were limitations placed on the level of detail of the mass balance. The balance

of solid and liquid phases over each unit was considered, but a full component mass balance was not. In

units where phase transitions occurred, such as desilication, digestion and precipitation, the mass of material

that changed phase was accounted for using conversion factors and relationships between process streams.

To ensure that the model was representative of plant performance and to capture the plant behaviour at a

level of detail that was appropriate, the detailed refinery SysCAD component mass balance was used to



develop relationships between process streams for the DES model. These relationships captured the

proposed operating philosophy in the plant under ‘normal’ conditions and allowed the representation of

issues such as seed charge and precipitation yield as a function of flow-rate, spent liquor return as a function

of bauxite feed, live steam flow to digestion as a function of flow-rate and washer overflow from the

counter-current decantation circuit to maintain the water balance in the model.

Findings While the initial focus on the model was to assess plant flow availability, sensitivity analyses quickly

highlighted a number of areas that were of interest. These included the settlers, the residue disposal pumps,

the polishing filters and the seed preparation circuit. This section outlines some of the process and findings

from the discovery phase.

Plant Flow Availability The first step in the discovery phase was the evaluation of overall plant flow availability, also known as the

operating factor or clear flow availability. This was measured as clear filtrate flow availability to

precipitation and was a key plant design parameter. The results from the model indicated a number of issues:

That nominal flow to precipitation was different to design,

That there were dynamic bottlenecks in the process that resulted in a reduced clear filtrate flow rate,

and

That the change in the system design bottleneck resulted in shortcomings in the proposed

maintenance philosophy.

While the original design of the system had assumed a particular figure for clear filtrate flow and

precipitation yield, the two figures key to alumina production, the model displayed different results.

Detailed investigation of this issue found that the flow rate to precipitation from the model was generally

below the design expectation. Production losses due to a lower average clear flow were somewhat offset by

a higher yield than expected, yield being inversely proportional to flow. This drove a higher than expected

clear flow availability requirement in order to achieve the same production.

An example output from the model showing clear flow availability versus production is shown in Fig. 2 for a

number of sample years of production from the model. This shows that the production at the design clear

flow availability was not delivering the target production capacity. In order to raise the production capacity,

a higher clear flow availability requirement would have to be set. This output also highlights the power of

the model in predicting clear flow availability and production for a given steady-state design and in

providing an understanding of the variability to be expected.



Target

Target

Avg Result

Avg Result

Cle

ar F

low

Ava

ilab

ility

(%

)

Production (Mtpa)

Fig. 2: Simulation Production and Clear Flow Availability Outputs

Subsequent investigation into the cause of the clear flow being below design expectations revealed the

presence of dynamic bottlenecks. While the design bottleneck under normal operations was digestion, the

variability inherent in the operation meant that the bottleneck was dynamic and shifted between different

process units. This shifting of the bottleneck occurred primarily when different random events disrupted the

process. These events included equipment breakdowns and planned maintenance as well as changes in plant

conditions, such as feed grade and operational performance. Investigation of the model showed that the units

that would bottleneck at different points of time were digestion, the settlers in clarification, the residue

disposal pumps and the polishing filters.

The investigation into plant flow availability also showed that as the system design bottleneck had shifted

from precipitation to digestion that the plant performance was highly sensitive to the equipment reliability

and maintenance philosophy applied in digestion and in the settlers downstream. The model was used to test

the benefits of several different maintenance schemes in this area. The model was also used to assess the

sensitivity of the refinery performance to a range of different digestion production and reliability rates,

showing that performance was sensitive and that there would likely be benefit in closely monitoring and

focusing attempts to improve reliability and performance in this area.

Clarification Settlers In the process, the clarification settlers are downstream from digestion and are the first stage in a counter-

current decantation (CCD) process separating out the solid residue, or red mud, from the alumina pregnant

liquor. The underflow from the settlers is sent into CCD and subsequently to residue disposal while the

overflow from the settlers is sent to the polishing filters and precipitation.

Some of the issues identified during the definition phase were that there were some major breakdowns in the

mid-term history of the unit, and that the settlers now sat downstream from the design bottleneck (digestion)

and as such would likely be more important for plant flow availability than in the previous configuration.



Steady-state design of the settlers had considered an operating rate and availability figure but was not able to

consider the various dynamic impacts. An example of this was that while the digesters and settlers are

normally unitised (one digestion train feeds a single settler), during longer-duration settler failures facility

exists to reconfigure the flow path to continue operation, albeit at a reduced rate. As a result, steady-state

design using a serial reliability calculation for the units underestimates reliability and overestimates the

impact on capacity. However, this assessment is not possible without a distinction between short and long

duration failures, which is also not captured in a single availability figure.

The modelling approach used for the settlers was refined over the course of the discovery phase as the

understanding of the dynamics around the settlers improved and the need for a more detailed representation

of the settlers in the model arose. As a basis, the settlers were modelled to include a bed for solids surge

capacity and control logic to maintain bed depth using the settler underflow pump, which assumed a fixed

density and solids content for the underflow stream while the balance of liquids introduced was transferred to

the settler overflow tank. Later refinements included the operating logic and limitations for reduced rate

operation of settlers during outages and the deferral of planned maintenance when another settler was down

due to an unplanned outage.

Some of the findings from investigating the settler performance in the model included an understanding of

the sensitivity to major breakdown events and occasions when reduced throughput mode was required. It

was found that while major breakdown events were relatively rare and had not occurred in the recent history

of the refinery, the potential impact of such rare events was significant enough to consider expensive risk

mitigation options, such as the provision of an additional settler or the increase of critical spares availability

for key parts on the settlers.

Residue Disposal Pumps The residue disposal pumps are downstream from the CCD circuit and pump the washed residue out into the

residue disposal area. They are positive displacement pumps that move high solids concentration slurry

against a very high pressure drop.

During the definition phase, the study team identified a pessimistic perception around the performance of the

unit. The high maintenance requirements of the pumps and the highly variable load required from the unit

led to a perception that they were a bottleneck in the process. This perception led to a call for a greater level

of redundancy and latent capacity to be provided in the upgrade than was proposed in the design.

The design of the residue disposal pumps had considered an average case of residue flow and a given

availability but did not consider dynamic impacts. Two of the key sources of variability for this unit were

identified namely:

1. There was a variable load of mud being pumped as a result of the changing feed grade, and



2. There were to be different rate machines installed to those that were currently in place, which

changed the availability of pumping capacity depending on which units were unavailable due to

planned maintenance or unplanned breakdowns.

As a result, the steady-state design was unable to consider much more than a scenario driven approach (low,

average and high mud flow) coupled with a design margin, in evaluating the capacity required. The steady-

state design approach could not consider the dynamics of feed grade, maintenance, reliability issues and the

extent to which the residue disposal pumps were decoupled from the rest of the process due to a significant

mud surge capacity in the CCD circuit.

The model used for the residue disposal pumps was refined over the course of the discovery phase, as the

need for a more detailed understanding of the plant dynamics arose. In the first instance, the residue disposal

pumps operated at a fixed capacity to pump away the solids underflow from the settlers with a given

maintenance schedule and subject to minor and major random breakdown events. Detail was later added to

look at the variability of residue flows using the historical feed grade profile. More detailed control logic for

controlling the number of pumps online was also added based on the residue accumulated within the CCD

circuit, allowing for a better assessment of the number of pumps required throughout the year and how often

an additional pump is used if required.

Some of the findings from investigating residue disposal pump performance using the model were that

conditions arose where the unit became a production bottleneck. It was shown that the combination of

random breakdown events with planned maintenance events during low feed grade periods resulted in a

bottleneck, restricting production capacity. It was also shown with the model, by comparative analysis and a

frequency analysis of pump requirements over the year, that while the variable demand load on the pumps

was significant, the number of pumps initially proposed was more than what was required. This was due to

the decoupling of the residue disposal pumps, to a certain extent, from the rest of the process by the residue

mud surge available in the CCD circuit.

The reduction in the number of residue pumps required that was identified using the dynamic simulation was

a substantial capital saving opportunity for the project.

Polishing Filtration The polishing filters are downstream from the clarification settlers and immediately upstream from the main

flow into precipitation. The polishing filters are batch operated filters that remove fine particulate matter

from the alumina bearing liquor prior to precipitation to ensure product quality. Essentially all of the liquids

flow goes through to precipitation, with a small recycle stream of cloudy filtrate drained after filtration

cycles. While the filters operate on a batch cycle, there are sufficient filters to ensure that the flow through

the unit is continuous, albeit with variability of performance.



It was identified during the definition phase and during subsequent visits to the refinery that the polishing

filters were acting as one of the main bottlenecks in the refinery. It was also identified that there was a

strong variability in the operating performance of the equipment and that filtration performance could change

significantly over short periods of time for reasons unknown in the plant. This anecdotal evidence, and the

position of the polishing filters prior to the existing design bottleneck of the refinery namely precipitation,

focused attention on the polishing filters from very early in the study.

The steady-state design of the polishing filters had assumed a pseudo-continuous mode of operation over the

unit and worked on the basis of filter availability and a required available number of filters. There was a

limited ability to understand the range of equipment performance, the implications of the interface between

the batch-operated filters and surrounding continuous processes as well as the implications of not being

directly upstream from the design bottleneck on clear flow availability.

0

5

10

15

20

25

30

35

40

Rel

ativ

e F

req

uen

cy (

%)

Average Filtration Rate (m3/m2.h)

Fig. 3: Frequency Analysis of Polishing Filtration Rate Variability

The modelling approach used for the polishing filters was refined heavily throughout the project and became

one of the most detailed components of the model. Initially, the polishing filters were modelled individually

to include each of the steps in their cycle. These steps were filling, working, dumping cloudy filtrate,

washing with spent liquor every cycle, cleaning with caustic soda after a given number of cycles and

changing the filter cloth after a specified number of operating hours. Detail later added included historical

filtration rate performance variability around a nominal rate (Fig. 3) and definition of a system curve that

defined the nominal operating rate as a function of the number of polishing filters operating (Fig. 4). The

impact of shared resources was also considered i.e. that no more than one filter could dump cloudy filtrate

simultaneously. This was an issue that would increase in significance as the number of filters online

simultaneously increased. Detailed control logic for controlling the number of filters online depending on

flow requirements was also added into the model, to allow for a better assessment of the number of units

required.



Filt

rati

on

Flo

wra

te (m

3/h

)

Number of Polishing Filters Operating

Fig. 4: Filtration Performance Input Data

Investigation into the operation of the polishing filters showed that the combination of filtration performance

variability and batch operation resulted in the unit becoming a dynamic bottleneck that was not fully

reflected in the continuous flow assumptions. The model was used to:

1. Examine the operating states of the filters (Fig. 5), and

2. Perform frequency analysis on the number of filters used under the influence of plant dynamics and

variability (Fig. 6).

Performance variability and changing the shape of the system curve were also shown to be significant,

indicating that this area should be a focus for performance improvement efforts.

0 20 40 60 80

Proportion of Time (%)

Working

Idle

Caustic Clean

U/S Tank Low

D/S Tank High

Cloth Change

SPL Wash

Cloudy Dump

Fig. 5: Simulation Polishing Filter States Analysis



0

5

10

15

20

25

30

35

40

Rel

ativ

e F

req

uen

cy (

%)

Number of Polishing Filters Operating

Fig. 6: Simulation Polishing Filter Usage Results

Importantly, the model showed that while there was an incremental benefit as quantified in alumina

production, it was not sufficient to overcome the cost of increasing the number of polishing filters for the

facility due to a step change in capital costs.

Benefits There were a number of key benefits that arose from the use of dynamic simulation to support the project,

both tangible and intangible namely:

1. Demonstrating that the additional polishing filters and additional residue disposal pumps were not

required was a significant capital saving and more than recovered the cost of the simulation study on

the basis of installed capital alone.

2. Supporting discussions between the project team and operations staff, and to provide them with

objective analysis about the performance of different scenarios and articulate clearly why the

additional equipment was not required.

3. A less visible benefit of using the simulation is the saving made in the cost of making the wrong

decision during the design process. This benefit often outweighs the benefits of the simulation

model in making capital reductions, as not meeting the design target and producing under design

generally has significant operating revenue implications.

4. Its use in identifying those areas where large operability ranges are required and the implications of

reducing the operability ranges on the total plant. This is particularly important for trim streams,

such as evaporators, or utilities, such as power and steam, where an operability range reduction can

yield significant cost savings, or where operability limitations can introduce unanticipated

bottlenecks.



5. As the dynamic simulation evaluates plant capacity and throughput using equipment operating rates

rather than the ‘normal’ steady-state operating rates, it can be used to evaluate plant design margins

on equipment, or the difference between installed capacity and the ‘normal’ flow-rate from the

detailed mass and energy balance. This is of particular benefit in facilities where design margin

(such as a nominal 10%) comes at a very high cost.

6. Improving the understanding of the impact of dynamics on plant operation and the critical sources of

variability that can affect capacity. This understanding can be invaluable for the engineers and

operating staff that go on to later commission and operate the project, and it provides them with a

different perspective when considering variability and plant capacity issues in the future.

7. Finally the consideration of inter-facility level control schemes during the process design and

identification of those areas where improved control can yield a high-value outcome. In this case

study, at least one project based on inter-facility control was identified due to the design team

thinking from a dynamic perspective.

These points clearly demonstrate the return on investment on the use of dynamic simulation for equipment

selection and decision support.

INCORPORATING DYNAMIC SIMULATION INTO THE PROCESS DESIGN Using dynamic simulation to evaluate plant flow availability and using this in conjunction with the detailed

mass and energy balance can significantly improve plant design over the top-down or bottom-up estimates

commonly used. The key issues that remain to be addressed are:

When is the best time to introduce dynamic simulation, and

What can be expected from using this technique?

When to Use Dynamic Simulation The question of when to use dynamic simulation depends on:

Consideration of the data and effort required to develop the simulation,

When key issues in the process design will be investigated and scope reduced, and

The cost of re-working design or re-opening design pathways at a later stage.

In process design for a metallurgical plant, significant decisions for the plant are often made very early in the

project to avoid carrying forward too many options. Effectively by feasibility study and detailed design,

most of the options have been decided upon and there is a single design which is then optimised and refined

over the subsequent project phases. By this stage, the cost of re-work due to a re-evaluation of the process



flow availability and a change in the required plant flow under normal conditions is very significant, as it

now involves re-work at a higher level of detail and in conjunction with more team members.

With an increase in the front-loading of process design in projects, the best time in the project lifecycle to

begin the use of dynamic simulation is at the pre-feasibility stage. During this phase, the first HAZOPs and

design reviews begin to occur, and the benefits of considering the dynamic operation of the plant and its

operability concerns can have a real impact on the design.

Ideally, the simulation would start during the pre-feasibility study and be refined, with increasing levels of

detail as the project progresses through its subsequent phases. This would allow the dynamic simulation to

evolve with the design and for changes to the plant flow availability to be incorporated into the iterative

design process.

Expectations from Dynamic Simulation While the outcomes for dynamic simulation depend on the level of detail and what is required, there are a

number of outcomes that dynamic simulation is very useful for. The analysis of plant throughput / capacity

and estimation of plant flow availability, which has been the focus of this paper, is a key outcome from the

dynamic simulation exercise. Other outcomes include using the simulation to evaluate the impact and

requirements of design margins, the use of the simulation to evaluate ‘creep’ or ‘stretch’ scenarios and the

identification / confirmation of design bottlenecks in each of these cases.

It can also be invaluable in assisting the engineers and operating staff that go on to commission and operate

the plant.

CONCLUSIONS In summary, dynamic simulation can be used to support the design process in continuous metallurgical

process plants by quantifying the plant flow availability and improving the understanding of plant dynamics

and variability. Dynamic simulation has distinct advantages over other techniques such as top-down and

bottom-up static reliability estimates, including:

The detailed representation and understanding of the implications of batch and continuous interfaces,

The effects of randomness and dynamics in the plant, and

The implications for equipment selection, capacity and spares.

The development of a separate dynamic simulation using DES to use in conjunction with the detailed plant

steady-state mass and energy balance is an effective and powerful way to use these techniques. The time and

effort required to develop a single model that answers both the steady-state and plant flow availability



questions is prohibitive, and it is not an approach that is tailored to provide effective decision support in the

earlier process design phases when there are more options and the design is rapidly changing.

In the alumina refinery expansion case study presented, DES has been used to complement the detailed mass

balance employed for process design. It has been used to quantify plant flow availability and plant

throughput as well as evaluate equipment selection and spares requirements.

From a practical perspective, dynamic simulation is flexible enough to be used effectively at pre-feasibility

stages in the design and to be refined as the project progresses. It is important to incorporate dynamic

simulation earlier rather than later from a process design perspective, as the number of options available

quickly diminishes and the cost of re-work due to process flow availability changes significantly increases as

the project progresses.

By using dynamic simulation to evaluate plant flow availability, the risk can be reduced for over-

capitalisation of equipment and high capital costs or, alternatively, under-performance relative to the design

target and significant losses in operating revenues. The cost of the dynamic simulation to minimise these

risks has substantial payback in terms of project certainty and profitability.

REFERENCES Glassock, C and Hoare, R, 1995. An application of dynamic simulation to a bulk solids handling facility, in Proceedings 5th

International Conference on Bulk Materials Storage, Handling and Transportation, pp 291-297 (Institution of Engineers:

Barton).

Hoare, R, 2007. The role of simulation modelling in project evaluation, in Proceedings Project Evaluation Conference 2007, pp

121-126 (Australasian Institute of Mining and Metallurgy: Melbourne).

Koenig, R, Marinopoulos, N and Olsson, B, 2002. Using reliability modelling to confirm plant design capacity, in Proceedings

Metallurgical Plant Design and Operating Strategies 2002, pp 248-260 (Australasian Institute of Mining and Metallurgy:

Melbourne)



Secondary Copper Processing using Outotec Ausmelt TSL Technology

J Wood1, S Creedy2, R Matusewicz3 and M Reuter4

1. Senior Process Engineer, Outotec Ausmelt Pty Ltd, 12 Kitchen Road, Dandenong Vic 3175. Email: [email protected]

2. Marketing Engineer, Outotec Ausmelt Pty Ltd, 12 Kitchen Road, Dandenong Vic 3175. Email: [email protected]

3. Technical Development Manager, TSL Smelting, Outotec Ausmelt Pty Ltd, 12 Kitchen Road, Dandenong Vic 3175. Email: [email protected]

4. Director – Technology & Product Management, Outotec Ausmelt Pty Ltd, 12 Kitchen Road, Dandenong Vic 3175. Email: [email protected]

ABSTRACT Steady depletion of the world’s primary copper reserves coupled with a meteoric rise in electronic waste (e-waste) generation has led to secondary copper processing assuming an ever-increasing importance within the global copper industry. Treatment of secondary copper feeds at existing smelting operations has traditionally been performed in the Blast furnace, Peirce-Smith converter and/or anode furnace. In the past decade or more, bath smelting technologies have been preferred for the processing of these materials due to their superior environmental performance and flexibility to operate under a wide range of conditions. This paper focuses on the pyrometallurgical processing of copper secondaries in the Outotec Ausmelt TSL Furnace and some of the unique issues facing secondary copper smelters and downstream gas handling/cleaning equipment.

INTRODUCTION Primary copper smelting has long been the dominant processing route for copper production. In the past decade the proportion of global copper production derived from copper secondaries has hovered around 35% (Fig. 1) (International Copper Study Group, 2010).

Fig. 1 – Global Copper Production (International Copper Study Group, 2010)



More recently, the copper industry has witnessed significant growth in the recovery of copper and other metal values from secondary materials (Reuter and Van Schaik, 2008), a number of which are listed and illustrated below (Fig. 2):

Metallurgical wastes - low grade slags, residues, anode slimes, oxide residues etc.;

industrial wastes - copper sheeting, bars, pipes, wire, ship screws, etc.;

consumer wastes - brass and bronze applications;

electrical and electronic waste (e-waste) - domestic electrical, audio-visual, computer and telecommunication appliances.

Fig. 2 – Various Secondary Copper Materials

A significant fraction of global copper production is derived from copper secondaries (International Copper Study Group, 2010), with key factors for this being:

Tightened supply of copper concentrates;

a drive towards improved energy efficiency during copper production, with copper recycling providing energy savings of up to 85% compared with primary smelting (Bureau of International Recycling, 2010);

the desire to enhance the environmental performance and sustainability of the copper industry, with the recycling of copper scrap reducing CO2 emissions by 65% (Bureau of International Recycling, 2010);

increased availability of secondary copper materials, particularly e-waste, for which global production has soared to more than 40 million tonnes per year as a result of the rapid growth of electronic markets and short lifespan of electronic products (United Nations Environment Program, 2009);

legislation mandating the treatment of e-waste (European Parliament and Council Directive, 2003);

increased profitability of e-waste processing due to the high levels of copper (typically 5 - 30 wt. %), precious metals and platinum group metals (PGMs) compared with sulphide concentrates (United Nations Environment Program, 2009);

SECONDARY COPPER PROCESSING

The recovery of copper and other metal values from secondary copper feeds has traditionally been carried out via:

The ‘Knudsen Process’, patented in 1915 and its variants, whereby an impure ‘black copper’ product is generated under reducing conditions (typically in the blast furnace) and



subsequently oxidised to raw copper using Peirce Smith or Hoboken converting (European Commission, 2009);

re-melting/alloying copper and copper based alloy scrap (brass mills and wire rod plants);

remelting in existing Reverberatory, Peirce-Smith converting or anode furnaces.

In the last 15 years however, there has been a shift towards secondary copper processing using bath smelting technologies such as the Outotec Ausmelt Top Submerged Lance (TSL) and the Outotec Kaldo Top-Blown Rotary Converter (TBRC) processes. A number of hydrometallurgical processing routes have also attracted some interest in recent times (Cui and Zhang, 2008), although there are ongoing concerns over the handling of intermediate products such as sludges, precipitates and other solutions from these operations. Table 1 summarises a number of smelting technologies recently adopted by selected global operations for the processing of secondary copper feeds.

Despite growing interest in the treatment of secondary copper feeds, the pyrometallurgical processing of these materials is characterised by a number of issues not prevalent in primary smelting. Furthermore, the presence of elements/compounds not typically contained in concentrate feeds have implications on the design and operation of equipment and processes, presenting difficulties in the treatment of secondary copper materials within existing smelter operations.

To address these factors and the increasing availability of secondary copper feedstocks, a number of facilities have been established in the last decade or more for the dedicated processing of these materials. These operations have focussed on achieving the necessary flexibility to adapt to changes in feed composition and availability in combination with the tight control of key process variables to ensure stable operation for the effective recovery of metal values while ensuring best practice in environmental control.

Aspects of Secondary Copper Processing

One of the primary issues faced during the pyrometallurgical processing of secondary copper materials and end-of-life consumer goods lies in the highly variable physical and compositional nature of feeds being treated.

In primary concentrate smelting, despite differences in the composition of feeds treated by the world’s copper smelters, process flowsheets and operating conditions at these facilities are largely the same. Variations in concentrate grade, gangue content and impurity levels, intrinsically linked to the geology and mineralogy of the parent ore body are typically managed via appropriate changes to the smelting operating conditions, without the need for extensive modification of plant equipment. Well known ore mineralogies have made it possible to optimize smelting processes over many years, enabling relative robust process control of the smelting process. In contrast to primary concentrates, the availability, physical characteristics and composition of secondary copper feeds is subject to both significant and rapid variation.

With reference to concentrates, secondary copper smelter feedstocks are often made up of many different materials obtained from a wide variety of sources. Hence the quantity and characteristics of the overall feed blend are strongly dependent on material availability, economics and sourcing practices employed. E-waste feeds in particular are influenced by factors such as consumer trends, product designs, technological developments and government legislation (e.g. replacement of lead with other elements in solder, increased antimony content as a flame retardant in resins/plastics and replacement of CRT devices with LCD units which contain for example indium-tin-oxide etc).



Table 1 – Selected Pyrometallurgical Secondary Copper Processing Technologies (Laney and Peterson, 2006, Cui and Zhang, 2008, European Commission, 2009).

Process Technology Location Feed Materials Metals Recovered Process Summary

TSL

Outotec Ausmelt TSL

Dowa Mining Co. Ltd. Kosaka Copper Smelter

Akita, Japan

E-waste Cu residues

Cu, Ni, Ag, Au, PGMs, Pb, Zn,

Sb, Sn

Multi-stage batch smelting and reduction operation to produce raw Cu and recover Pb, Zn and other metal values.

Outotec Ausmelt TSL

Global Resources & Materials (GRM) Danyang, Korea

E-waste Cu residues

Cu, Ni, Ag, Au, PGMs, Pb, Zn

Continuous smelting under reducing conditions to produce a ‘black copper’ product treated in downstream operations.

Xstrata ISASMELT®

Umicore Precious Metals Hoboken, Belgium

E-waste Industrial wastes

Cu residues

Cu, Ni, Ag, Au, PGMs, Pb, Sb,

Sn, Bi, Se, Te, In

Batch smelting and converting to produce blister Cu. Part of the overall Base Metal and Precious Metal Operations Flowsheets.

Xstrata ISASMELT®

Aurubis Lünen, Germany

Cu scrap Cu Residues

Cu, Pb, Sn, Zn, Batch smelting (reductive) to ‘black Cu’ and converting to produce raw Cu. Part of the overall Kayser Recycling System (KRS).

TBRC

Outotec Kaldo

Boliden Rönnskår Smelter Rönnskår, Sweden

E-waste Cu, Ni, Ag, Au, PGMs, Se, Te

Batch converting of e-waste to produce a mixed Cu alloy treated in downstream converting operations.

- Metallo Chimique N.V.

Beerse, Belgium ‘Black copper’ Cu, Ni, Pb, Sn

Processing of ‘black copper’ sourced from external operations.

Noranda Process

- Horne Smelter

Quebec, Canada E-waste

Cu concentrate Cu, Ni, Ag, Au, PGMs, Se, Te

Continuous smelting of primary and secondary feeds to produce Cu matte treated in downstream converters.



To address variability in feed composition, feed size, impurity levels and plastics/organics content, preparation and blending operations are often utilised prior to feed introduction to the smelting vessel. Crushing, grinding, drying and/or agglomeration stages are commonly used to transform ‘raw’ secondary copper materials into a form amenable to smelting. Consumer and e-waste materials usually undergo some form of manual, semi-automated or automated dismantling, mechanical pre-processing and/or sorting.

For example E-wastes are often manually dismantled to separate batteries, printed circuit boards (PCBs) and hazardous materials prior to shredding and mechanical sorting to separate copper and recover ferrous, aluminium and plastic fractions which are recycled via other means. Due to the complexity and functionality of electronic products, the extent to which liberation/separation of desired metal values can be achieved through feed pre-treatment is driven by economics (e.g. the inherent value of contained precious metals and PGMs in the recyclates). Smaller items such as mobile phones, portable MP3 players and cameras are consequently introduced to the smelter with minimal pre-processing (United Nations Environment Programme, 2009) due to these economic considerations especially as every additional processing step could increase the loss of valuable materials/metals.

Despite use of such feed pre-treatment and blending practices, establishing a long-term homogeneous feed blend is still extremely difficult to achieve. Consequently, secondary copper smelters must be flexible and sufficiently versatile to adapt their operating conditions so as to account for variability in feed morphology and composition. This is especially so given that the physical and compositional characteristics of secondary copper materials have design and operation implications to both the smelting vessel and downstream offgas handling and cleaning systems. Furthermore, tight control of process operating conditions during secondary copper smelting is required for the separation of impurities and desired metal values between the various product streams.

Secondary Copper Processing in the Ausmelt TSL Furnace Outotec’s Ausmelt TSL Furnace is ideally suited to the processing of secondary copper materials given that the physical characteristics of feeds introduced to the furnace are not overly critical to its operation (unlike technologies such as Flash Smelting). Feed materials able to be treated in the Ausmelt TSL Furnace include:

Heavy, bulky items and/or lumpy materials (e.g. fittings and crushed metallurgical wastes);

fine, ‘fluffy’ and/or dusty materials (e.g. shredded materials and metallurgical dusts);

complex waste electrical and electronic equipment (WEEE);

irregular sized materials (e.g. scrap off-cuts and fittings);

high moisture content materials (e.g. residues).

The complex and variable composition of copper secondaries are also easily handled by the Ausmelt TSL Furnace through precise control of the process chemistry, temperature and bath oxygen potential (pO2). Operating conditions within the bath are regulated via the injection of fuel, air and in some cases, oxygen directly into the slag phase using a submerged lance (Fig. 3). Conditions above the bath ‘splash’ zone are regulated independently of the bath through the addition of air via a dedicated lance ‘shroud’ system. Shroud air also provides for the efficient recovery of heat generated from the combustion of volatile components of the feed without influencing the bath oxygen potential.

The precise level of control achieved in the Ausmelt TSL Furnace enables the recovery of metal values to targeted product phases from which they can be economically recovered whilst impurities and gangue components are directed to a discard slag or by-products from which can be safely treated (Fig. 3).



Fig. 3 – Possible Distribution of Elements for Secondary Copper Smelting in the Ausmelt TSL Furnace Under Neutral/Oxidising Process Conditions

Typically, secondary copper materials are processed in the Ausmelt TSL Furnace using one of two basic flowsheets (Fig. 4), both of which are variations of the Knudsen process.

Fig. 4 – Two possible Ausmelt TSL Furnace Secondary Copper Flowsheet Options



In Flowsheet 1, a continuous ‘reductive smelt’ stage is used to generate an intermediate black copper product and low copper content discard slag. Treatment of this black copper to produce a raw copper product may subsequently be achieved via:

An oxidative converting stage in the same Ausmelt TSL Furnace (two-stage batch process);

converting in separate units (e.g. Peirce-Smith converters).

This flowsheet provides benefits in instances where an Ausmelt TSL Furnace is used to replace a blast furnace within an existing secondary copper operation, given that downstream converting and refining operations typically already exist. This flowsheet may also be favoured when treating lower grade materials, as the generation of a discard slag in the smelting stage eliminates the need for downstream processing (copper recovery) of this material.

Conversely, in instances where feed materials are characterised by large variations in copper, precious metal and/or impurity levels, it is often beneficial for smelting to be carried out using a multiple stage batch process (Flowsheet 2). Selective introduction of customised feed blends in each stage and the precise control of furnace operating conditions (particularly pO2) in this flowsheet provides for the maximum recovery of valuable metals and elimination of impurities with the slag.

In the first oxidative smelting stage, the high-grade copper product generated acts as a collector for valuable minor elements (PGMs, precious metals, cobalt etc.) which are recovered using sophisticated downstream refining operations. Slag from this stage is subsequently processed under reducing conditions in Stage 2, providing for the recovery of desired metal values to a ‘dirty’ black copper product, retained in the furnace for next smelting cycle. Impurities and gangue components of the feed meanwhile are distributed to the discard slag.

Additional benefits of this flowsheet include:

Desired metal values may be recovered in separate product streams;

an ability for the thermal duty in each stages to be optimised;

slag chemistry adjustment in each stage, reducing the overall process flux requirement.

Ultimately, flowsheet selection is based on the grade of secondaries being treated, the type and concentration of impurities in the feed, the nature and capacity of existing processing/refining infrastructure and client’s preferred product stream/s. Furthermore, the inherent versatility and flexibility of Ausmelt TSL Technology allows for the addition and/or removal of extra stages, even if not included within the original design.

In addition to control of the furnace bath conditions (temperature, pO2, slag composition etc.), handling and cleaning of process offgas is also considered as part of the overall Ausmelt TSL Plant. Secondary copper materials are commonly associated with various types of plastics and ceramics, many of which contain elements which produce toxic/hazardous substances if processed under certain conditions.

As a result, gas collection and abatement systems typically located downstream of the Ausmelt TSL Furnace (Fig. 6), incorporates capabilities to handle:

Nitrogen oxides (NOx) and carbon monoxide (CO);

sulphur dioxide (SO2) generated from the treatment of sulphidic materials such as residues;

volatile metallic species (Pb, Zn, Hg, Cd etc.) and dusts;

halides (Cl, F, Br etc.);

polychlorinated dibenzoparadioxins (PCDD) and polychlorinated dibenzofurans (PCDF) generated from the processing of organic and chlorine containing compounds.



Fig. 6 – Schematic of a Typical Ausmelt TSL Furnace Secondary Copper Gas Handling Configuration

In spite of various feed treatment and sorting initiatives, which can mitigate the formation of harmful dioxins, furans and other compounds; tight control of operating conditions both within the smelting furnace and offgas handling system is still required.

CONCLUSION

The recycling of secondary copper materials and processing of ‘end-of-life’ consumer goods is a significant contributor to global copper production. Flowsheets and technologies employed in secondary copper operations differ considerably from those in primary copper smelting operations due to variability and complexity in the composition and physical characteristics of materials being processed. Consequently, secondary copper smelters must incorporate technology with the flexibility to vary operating conditions and practices with the ever-changing nature of feed materials being treated. Outotec’s Ausmelt TSL Technology offers precise control of key process parameters and versatility in flowsheeting options, necessary for secondary copper smelting. This is complemented by the Outotec range of offgas handling solutions which may be customised for a particular process or application to ensure that the technology operates well within best available practice and strictest environmental legislation.

REFERENCES Bureau of International Recycling, 2010. Recycling Facts – Copper [online]. Available from <http://www.bir.org/industry/non-ferrous-metals/> [Accessed 23 February 2011]. Cui, J and Zhang, L, 2008. Metallurgical recovery of metals from electronic waste: A review, Journal of Hazardous Materials, 158: 228-256. European Commission, 2009. European Dioxin Inventory - Secondary Copper Production [online]. Available from: <http://ec.europa.eu/environment/dioxin/pdf/stage1/seccopper.pdf> [Accessed 23 February 2011]. European Parliament and Council Directive 2002/96/EC of the European parliament and of the council of 27 January 2003 on waste electrical and electronic equipment (WEEE). Official J. OJ L 37 of 13.02.2003, Brussels Belgium. International Copper Study Group, 2010. The World Copper Factbook 2010 [online]. Available from: <http://www.icsg.org/> [Accessed 02 March 2011]. Laney, M and Peterson, P, 2006. Current Status of Secondary Copper Production in the United States. Prepared for U.S. Environmental Protection Agency (EPA Contract 68-D-01-073). Reuter, M and Van Schaik, A, 2008. Thermodynamic Metrics for measuring the “Sustainability” of Design for Recycling, Journal of Metals, Vol. 60(8): 39-46. United Nations Environment Programme, 2009. Recycling – From E-waste to Resources [online]. Available from: <http://www.unep.org/PDF/PressReleases/E-Waste_publication_screen_FINALVERSION-sml.pdf> [Accessed on 25 February 2011].



Solvent Extraction of Uranium – Towards Good Practice in

Design, Operation and Management

P Bartsch1 and S Hall2

1. MAusIMM, Manager Technology and Projects, Mitsui Uranium Australia, Level 16, 91 King William Street, Adelaide SA 5000. Email: [email protected]

2. MAusIMM, Manager Metallurgy, Uranium SA Ltd, 32 Beulah Road, Norwood SA 5067. Email: [email protected]

ABSTRACT

Uranium solvent extraction, USX has been applied commercially for recovery and concentration for over 50 years. Uranium in acidic liquor, which is prepared following ore leaching, solid/liquid separation and clarification, can be treated through a sequence of operations; extraction-scrubbing-stripping, to obtain purified liquor. USX has dominated the primary uranium industry as the preferred technological route to converter grade yellowcake.

The practices of design and operation of USX facilities has found renewed interest as new mines are developed following decades of industry dormancy. This article seeks to outline principles of design and operation from the practitioner’s perspective. The discussion also reviews historical developments of USX applications and highlights recent innovations. This review is hoped to provide guidance for technical personnel who wish to learn more about good practices that leads to reliable USX performance.

INTRODUCTION

This paper provides a brief summary of practical knowledge and lessons learnt in the field of uranium recovery by solvent extraction, (SX) from testwork, through design to production. These suggestions, clarifications and observations are offered on the basis of the authors’ experience which are hoped to encourage constructive discussions in the hydrometallurgical community, and thereby improve solvent extraction performance in current and future facilities.

FOUNDATION CHEMISTRY

Uranium solvent extraction, (USX) involves exchange of dissolved metals, as complex ionic components, from an aqueous phase into an organic phase or solvent. The organo-metal complexes may be purified by scrubbing, before they are subsequently stripped into a more pure, concentrated aqueous phase. The soluble, metal complex, or ion-pair, is chemically loaded by an extractant, which is dissolved in an hydrocarbon diluent, (Bosman 1980) which together form the solvent. An exchange of metal ion pairs will allow a charge balance to be maintained across the transfer interface, (cations or anions will be exchanged depending on the selected extractant). Interested readers can learn more detailed chemistry of USX in useful texts (IAEA 1993, Ritcey and Ashbrook 1979, Merritt 1971, Brown et. al. 1958).

SX FEED CONTROL

The most important feature for good SX practice is stable stream flow, particularly pregnant leach solution, (PLS) and hence steady process chemistry. Every effort of designers, and the eternal vigilance of operators, must be aimed at consistent rates and composition of feed streams, and thereby internal and recycle streams. Realistic sizing of PLS storage, as well as sensible make-up capacity for reagent streams are essential aspects for reliable performance of a SX facility. Operators must expect



that crud, which is a generic term for dispersed solid phases that accumulate in settlers, (Ritcey 1980) will occur. Adequate systems and procedures must be installed and maintained to collect and treat this nuisance material so that reliable flows and hence performance can be maintained.

The relative performance of extraction, scrubbing and stripping can be characterized by coefficients for uranyl complexes which will shift with the overall sulphate, SO4 and Cl concentrations in the respective aqueous phase, (Soldenhof and Davidson 2005). Testwork for process specification must be conducted with representative solutions’ chemistry and temperatures, and transfer results interpreted for process criteria. Interpretation of kinetic and equilibrium effects will be important. For example SO4 levels in PLS may rise to over 100 g/L, (Thiry and Roche 2000), while few uranium leaching operations have operated with 'high' Cl concentrations, e.g. above 5 g/L. The Honeymoon Project, (Dobrowolski 1983), will operate in-situ recovery in a saline aquifer with about 8 g/L Cl in PLS, and use solvent extraction for uranium purification, (Mayfield and Aker 2006).

PLS may contain significant nitrate which can depress uranium transfer. Nitrosyl in sulphuric acid, (NOHSO4 reported as NOx) at levels of above 50 ppm may be present in delivered acid that is manufactured from smelter gases, (Lyne et. al. 2002). Hydrolysis to nitrates, together with contamination from residual explosives in mined ore means nitrates can accumulate in PLS depending on the closure of the process water balance, (Munyungano et. al. 2008).

For each impurity that is partially transferred through various scrubbing and stripping stages, potential flowsheet additions such as selective precipitation, or re-leach, can be investigated to ensure the specifications of customers, (i.e. uranium converters) are met for marketed products, (Bellino and Mackenzie 2010). In this respect detailed quantification of PLS chemical speciation is interesting for process optimization, but largely superfluous compared to the leach chemistry in respect to selecting an economic or preferred flowsheet unless precipitate specifications are threatened.

PROCESS STAGES - EXTRACTION, SCRUBBING AND STRIPPING

Solvent Extraction

Cationic or neutral extractants, i.e. long-chain alcohol, phosphoric derivatives, were first proposed for uranium recovery from PLS, (Grinstead et. al. 1958) but were superseded by amines due to greater selectivity. These reagents are no longer used alone for uranium recovery from PLS, but are popular for yellowcake purification at converter operations. Organic phosphate derivatives are commonly used for purification of uranium concentrates or oxides following digestion with nitric acid where the purpose is the removal of anion impurity fractions, e.g. silicates, sulphate and zirconia.

Mixed solvent, i.e. cationic and anionic extractants applied in the same transfer stage, was first described in reports from Colorado, (Rosenbaum et. al. 1958). Operations at the Durango facility with mixed solvent were described. Application of mixed solvent, together with a neutral modifier, (tri-butyl phosphate, TBP or di-butyl-butyl phosphonate, DBBP) were investigated at Honeymoon for treating PLS that contains elevated Cl in process waters, (Phillips 2001). This project will include the first application of solvent extraction combined with in-situ leaching for recovery of uranium, (LaBrooy et. al. 2009).

Anionic extractants such as proprietary tertiary amines, (i.e. organic, nitrogen compounds with three alkyl chains in liquid form) have become favoured for uranium recovery from acidic PLS because base metal impurities are efficiently rejected into raffinate. Silica and halides, (F, Cl, Br) can form simple anions, and some transition metals form soluble complexes, e.g. V or Mo as oxy-anions, Co as cyanide, Bi as oxy-chloride, or Fe in the presence of chloride, are also extracted by amines, (Bender et. al. 2010).

Certain additives tend to inhibit third phase formation that can arise as semi-miscible fluids of extractant-metal salts which separate and accumulate at the phase interface within settlers, (Kertes1966). Solvents that contain amines often utilize modifiers such as long chain alcohols, iso-decanol or tri-decanol, at about half the extractant concentration, (%v/v).



Dosing the solvent with mixed polarity modifiers can influence the uranium extraction co-efficient, hence lowering extraction of uranium in amine systems. Modifiers in cationic solvents include TBP and DBBP, which can have synergistic influence to improve uranium extraction, (Blake et. al. 1958). Appearance of a third phase is more prevalent during high extractant or uranium concentration due to the potential formation of organo-metallic salt.

Management of solvent degradation can become an operational challenge in SX circuits with high nitrate or ferric concentrations in PLS. Declining extraction performance, or extended phase break time are indicators that solvent properties require improvement. Where SX facilities have high organic entrainment in raffinate, or include solvent bleed streams that dispose of excess crud, the degradation products may not reach nuisance levels. In this respect the control of PLS chemistry is a good strategy, while uranium leaching rates and extent are managed, with acid and oxidant dosing for pH and Eh control, to reach economic levels. Alternatively PLS treatments to lower its chemical potential are also available, e.g. dosing with metallic iron or sparging with pure SO2 gas.

Bisulphate loading at high PLS acid level can suppress uranium extraction when leaching solution contains above 10 g/L free acid in SX feed. Controlled acid in PLS will offer benefits of economic leaching and SX performance, (George and Ross 1967). Lower acid levels for leaching are recommended, consistent with reaching uranium dissolution and mass transfer targets, as well as balancing silica dissolution, and accounting for downstream PLS and raffinate or tailings treatment costs.

Given the rapid kinetics of sulphate, bisulphate and uranyl sulphate exchange on solvent in the extraction stage, the relative loadings of U will quickly reach equilibrium. While the PLS has pH between 1.5-2.5 the stage-wise extraction equilibrium is usually available within 1-2 minutes of mixer retention, depending on mixer box design and mass transfer gap, (Faure and Tunley 1971). Stage design with twin or triple sequenced mixer boxes may be cost effective to obtain efficient mass transfer and adequate residence time distribution with low temperature PLS.

Loading and distribution of uranium on amine extractant will generally improve with increasing acidity of PLS. Adjustment of PLS with dilute sulphuric acid to lower pH to improve extraction performance has been mentioned at ERA’s Ranger operation, (Nice and Banaczkowski 2000). The solvent performance can also shift with PLS sulphate and chloride concentration. Each system must determine the competition of uranium and other anions in highly saline PLS with operating solvent, and allow for cost of reagent inputs.

Solvent Scrubbing

Solvent scrubbing can be applied advantageously between extraction and stripping to displace anionic impurity from amine based solvents, or metal cations from phosphate based solvents. Water scrubbing can wash out entrainment of PLS from the advancing solvent.

The number of scrub stages and make-up of scrub liquor will depend on the impurity type and relative loading on solvent compared to uranium. Dilute sulphuric acid, low chloride water, or strip liquor bleed, i.e. ammonium sulphate solution, can be used in 2 or 3 sequential scrub stages, (Hardy 1978).

Generally, spent scrub liquor will be recycled to the PLS pond, or directly into the extraction feed stream, to permit recovery of the contained uranium. Process stability and entrainment are also important criteria for the choice of scrub stages and stream recycle destinations.

Solvent Stripping

Strong ammonium or acid sulphate solutions, ca > 150-300 g/L SO4 can be used for uranium stripping from tertiary amine extractants, (Ritcey and Ashbrook 1979). In this process step the sulphate concentration is the dominant chemical which displaces the uranyl sulphate into the aqueous phase, while ammonia gas or solution is applied to regulate pH. Stripping with strong sulphate solution is preferred above dosing with ammonium hydroxide, (aqua-ammonia) due to more stable operation.



Sulphate strip chemistry can avoid localized, high pH which can cause premature precipitation of yellowcake in settlers and hence subsequent crud generation.

Stripping kinetics can be accelerated at marginally higher pH or increased solution temperature. Stage-wise control of pH rise across the stripping must be graduated. Tightly regulated stripping chemistry may be countered by pH regulation difficulties, or the presence of excessive crud.

Stripping with strong ammonium sulphate solution is preferred if the yellowcake precipitation process uses ammonia gas to make ammonium di-uranate as a marketable product. Residual ammonium in tailings liquor, which arises from the stripping process, cannot be released off the mine site. Hence ammonia recycle via lime boiling, or production of ammonium sulphate from evaporation, will be a necessary flowsheet inclusion.

Strong acid stripping can be attractive for flowsheets where acid, and caustic or magnesia for neutralization, are low cost and high quality. Acid stripping can avoid the need for a solvent regeneration stage and potentially simplify the overall SX circuit design. Carbonate stripping will be preferred when uranyl peroxide is the desired form of yellowcake or sulphate stripping is undesirable for environmental reasons, (Edwards 1992). The maximum uranyl concentration in strip liquor will be determined by the extractant concentration, and impurity loading, to ensure product specification, or avoid third phase formation in the strip stage.

PROCESS STABILITY

SX facilities will meet the design criteria when the feed flow and concentrations, particularly the uranium tenor, have low variability. Changes to PLS flow control set-point once per day are recommended, while efforts to minimize swings in uranium and impurity tenors, and hence mass rates, will be beneficial to overall production. Otherwise the SX operator will need to vary the PLS flow proportionally to the uranium tenor to maintain productivity, i.e. constant ‘metal’ feed rate.

PLS, scrub or strip aqueous phase flow variations will induce the respective settler interface level to rise or fall progressively through the sequence of settlers. A stepped flow increase can displace crud from the settler interface to the next mix box where more crud will be generated due to the turbulent intensity at the impeller tip. The transmission of excess crud in sequential mix boxes can promote crud generation through the entire system, i.e. ‘crud begets crud’.

Large storage volume can subdue fluctuations of PLS uranium tenor. An added benefit of thoughtful design will be suspended solid removal, which can be available from a generously sized PLS pond. Minimization of PLS storage volume is poor design practice that is often driven by mis-guided efforts of project managers to cut facility capital costs prior to start-up, without accounting for overall life-cycle costs of the asset. Zealous cost cutting can also compromise equipment selections for cleaning PLS such as filters or clarifiers. Dynamic process simulation can be helpful to optimized sizing of process storages for hydrometallurgical facilities, (Smith 2005)

Stable PLS tenor is an important ally for operators to control impurity transfer. If the loaded solvent carries high impurity loads then scrub liquor concentrations can be adjusted to suit anion transfer rates. Progressive process start-up, shut-down and changes to phase ratios of each stage can be managed carefully to avoid unplanned crud displacement, and transient impurity transfer though to final product.

For high uranium concentration in PLS, the SX plant could be run with alternative series-parallel extraction configuration, while the raffinate is recycled to ore leach. This design option will be determined by mine plans but is unlikely to become favoured for new plants that have low ore grade and do not recycle process waters or discharge raffinate to tails.



SX EXCHANGE EQUIPMENT

Process engineers will apply results from relevant extraction, scrubbing and stripping investigations to prepare ‘isotherms’ which are derived from shake-out test data, or pilot plant testwork, and obtained with varying ratios of PLS and solvent. The relevant test techniques and interpretation procedures are generally known in technical literature, (Gupta and Singh 2003).

McCabe-Thiele diagrams are constructs of uranium distribution or concentration profiles across various ratios of equilibrated aqueous and organic phases, using results from laboratory or pilot tests run at pre-determined conditions and constant temperature, i.e. graphical isotherms, (IAEA 1990). Such interpretations are used to generate process criteria such as flow ratios of solvent to aqueous. Design criteria can then be specified and applied to equipment sizing and selection, including columns, for extraction, (Bartsch and Lawson 2003).

Initial SX investigations for pre-feasibility, and consequent preliminary designs may be done with synthetic solutions. Experimental programs based on actual leach solutions can be used for confirmation of process criteria during the various study stages, and hence detailed plant design. Fresh solvents can be used for preliminary shake-out tests, but aged or used solvent is recommended when process designs are used for final feasibility studies. The choice to run continuous mini-rig or pilot solvent extraction investigation will depend on the degree of innovation of process chemistry, flowsheet, equipment or unusual impurity occurrence. Testwork must be conducted by experienced technicians, and guided by practitioners who can translate the isotherms into design criteria and equipment sizing

Two major types of equipment are used for extraction of uranium from PLS; namely mixer-settlers and agitated columns. Materials of construction for hydrometallurgical plants and equipment include lined or stainless steel or lined concrete, (Robinson et. al. 2003). Fibre reinforced plastic bodies and linings have become popular for small mixer settlers, and most columns for primary uranium extraction applications.

Columns evolved from designs in the nuclear industry for purification of numerous metals including uranium. (Desson et. al. 1983). Column size and capacity have increased in recent installations, e.g. Olympic Dam. Columns are suited to multiple stages of extraction in a single unit consisting of active section with top and bottom decanters, and throughput is limited by hold-up of the dispersed phase. Columns can provide effective performance when mass transfer of uranium is relatively rapid, e.g. in extraction stages. Multiple columns, which operate in parallel, to accommodate larger flows, can be installed as determined by total flux, i.e. total SX feed plus advance organic flow rates. Columns may also be advantageous in comparison to mixer settlers due to lower, overall solvent inventory, or when advance phase ratios are widely different, (Movsowitz et. al 2000).

Mixer settlers were also originally developed in the nuclear industry, (Royston and Burwell 1972), while design improvements and scale-up was driven by copper extraction. These units have become popular for all transfer stages during the late 20th century, (Glasser et. al. 1976)

For extraction mixers the turn-down ratio of organic recycle over a wide range is useful for maintaining organic phase continuity. Likelihood of phase inversion arises if the mixer O:A occurs below 0.9:1.0 for extended durations. The authors experience suggests that organic continuity in extraction mixers is essential for management of emulsion break time and maintaining lower solvent losses by entrainment, i.e. avoidance of phase inversion. The effect of phase ratios and continuity on transfer of anions can be optimized with trials during commissioning and operation, and will depend on mixer retention time and potential for unwanted transfer by entrainment.

Overall process operation, and interface control in columns, can be easier than mixer settlers, particularly with respect to flow regulation and crud removal. Metal loading characteristics of columns in terms of approach to equilibrium are usually rapid with changes to PLS. Operators’



control of uranium loading in columns, or control of O:A holdup, have little influence on uranium transfer kinetics or equilibrium if the columns designs and scale-up are reasonable.

Extraction columns may run at O:A phase ratio at up to 1:10, whilst maintaining organic continuity of the hold-up emulsion, depending on PLS conditions, particularly suspended solids and temperature. Operators’ immediate response to changes of PLS flow or tenor in columns will be start-up or shut-down of individual units.

Columns can be idled while full to permit rapid re-start and avoid excess tank inventory of solvent. Sufficient solvent storage is needed outside the columns to allow individual column drainage of fluids for maintenance.

Columns for extraction will generally use more power than mixer settler configurations for the same duty if agitation is driven by pulse air, although direct comparison will depend on location and design criteria for the proposed operation, (Taylor 2007). The installed power for mixer impellers is more efficient than supply of pulsing air from blowers, but the number of mixers is often more than the number of blowers. Mechanical or hydraulic agitation designs for columns are not commercially available for PLS applications, but smaller installations for yellowcake purification and nuclear fuel reprocessing are available.

Columns are designed as closed vessels, (except for pulse air vents), with low roof space. Given good design columns can have greater fire suppression attributes, particularly if solvent purge or dump is arranged with due consideration of piping corridors. Columns will require more attention to piping detail during design to ensure adequate support and maintenance access are available.

SOLVENT CONDITION

A separate protonation or acidification stage is redundant while SX feed contains about 5-10 g/L acid. The recycle of spent scrub liquor to join PLS will lower its pH, which will assist process control if the upstream leach conditions are variable. In-line dosing of recycle or acid streams are a design opportunity if intermediate, sectional storages are available. Direct dosing of ‘neat’ acid into solvent streams must be avoided due to localized over-heating and oxidation.

Solvent scrubbing and stripping of uranium also tend to be relatively rapid, but retention may be dictated by dispersion and stability of reagent dosing for pH control. Equilibrium transfer of impurities during stripping can require 1-5 minutes retention per stage depending on solution chemistry temperature and mass transfer gap.

Regeneration of solvent is recommended if degradation products or interfering long-chain organic anions, e.g. carboxylates, accumulate in the running solvent. Such contaminants can be scrubbed from solvent in a separate stage with highly alkaline liquor, ca > pH 10, (MacDonald et. al. 1978).

The effectiveness of regeneration can vary with the recipe of the wash liquor, which will depend on contaminant type and load, by systematic experimentation. Strength of the regeneration liquor and campaign intensity will depend on extent of solvent deterioration as measured by break time. The Olympic Dam facility applied a three component mix, (Hall and Reed 1996), but operators will necessarily develop or procure their own recipe.

A routine, laboratory, maximum-loading test can assess the capability limit of the working solvent. Such tests which trend solvent performance over extended time frames. Tracking phase break results from testing the plant solvent are useful to determine when regeneration is required.

PHASE DISENGAGEMENT

The specific settling rates for uranium solvent extraction plant design are determined by design SX feed and solvent flow rates, with an allowance for crud interference. Experimental tests on pure fluids tend to over estimate the rate of phase break. Useful methods to assess phase disengagement in



operating plants includes regular, standardized break tests, e.g. separation of emulsion in 1L cylinder, on loaded and stripped solvent during each shift. A simple procedure can be established for use by plant personnel.

Specific rates for conventional settlers are generally set at 3-5 m3/m2/hour depending on extractant concentration, ionic strength, solvent temperature and expectation of crud make and accumulation, (Cheng et. al. 2004). Recent settler designs are proposed to enhance specific flow rates up to 5-15 m3/m2/hour, but potential for increased entrainment loses are not quantified. The flow pattern of emulsion from the mixer overflow to the settler will influence separation and solvent circulation. Installation of ‘picket’ fences within the settler box, located in proximity to the incoming flow, and spreading of organic-continuous emulsion from the mixer overflow into the solvent layer, is recommended.

Specific settling rates for pulse columns, when treating ‘clean’ PLS can be between 20-50 m3/m2/hour for the active, pulsing section. The total flux will depend on extractant concentration, solvent condition and system temperature.

Without diligent or automated crud removal a conservative settling rate is recommended. Recent design for scrub and strip stages have applied comparatively low settler rates due to the prevailing high crud content. Scrub and strip settlers do not generally receive suspended ore particles and silica solids, so crud management, and hence design settling rates will depend on precipitation of silica, zirconia or other impurities as observed during pilot trials.

CRUD FORMATION AND HANDLING

Suspended solids in PLS will enter USX following poor clarification of PLS. Dissolved species, such as silica and zirconia, have the potential to generate crud during solvent extraction processes.

Soluble silica in PLS becomes a problem when it polymerizes and converts to sub-micron colloidal, or larger particulate form, (Moyer and McDowell 1981) which can carry into the extraction, scrubbing and stripping sections. The silica colloids change in size as the silica polymerises and different size colloids can cause crud generation problems in different parts of the SX circuit, (Ritcey 1980).

Slimed ore particles and floating floccules from clarification can contribute to PLS solids loading and hence crud make in extraction settlers. These colloids and particles will tend to build up at the settler fluid interface, which causes stable emulsion build-up, and can restrict hydraulic throughput or lead to increased solvent losses, (Saruchera et. al. 2010).

Silica content of PLS, as a quantitative forward measure of potential crud make, can be determined quantitatively by laboratory filtering at defined filter pore sizes. Visual field observation of PLS solids, e.g. sighting in a clarity wedge, can be useful for rapid operator intervention. Field operators can collect ‘dip’ samples from each settler with a clear plastic tube, and thereby track crud build-up.

Maintenance of organic phase continuity in all SX mixers for uranium will limit crud formations and so limit crud build-up in settlers. In extreme circumstances bulk crud flows between settler banks can occur, otherwise known as 'crud-runs', which leads to higher solvent losses and poor metallurgical control.

Crud handling and treatment is important to recover solvent and lower metal losses, (Hartmann 2010). All uranium SX plants will experience crud formation and good plant designs will accommodate its removal and treatment. The design philosophy for PLS preparation as SX feed can best focus on prevention as compared to cure or remedial action. While less silica is dissolved during lower intensity leaching, and if more silica is precipitated and flocculated during PLS clarification, then less crud can be expected to accumulate in extraction settlers. Crud in downstream scrub or strip settler is often related to impurity transfer during unstable pH control.



Careful observation and impurity tracking during pilot testing can help designers and operators control of anion transfer and set stage pH targets. Preventative measures to avoid the deleterious effects of silica and crud include;

a) Control leach temperature, retention and acidity;

Extended retention time at lower temperature can ensure effective uranium dissolution;

Higher acid levels will tend to dissolve more silicate minerals and consume more acid;

Two stage leaching can keep final leach temperature down and conserve acid but at higher capital cost for installation of a second leach thickener or cyclone bank.

b) Minimize slime formation during grinding particularly of ores with clays or foliated silicates.

c) Operate effective PLS clarification; note that one stage of sedimentation will have difficulty reaching below 50-100 ppm suspended solids in SX feed when process upsets occur. A second clarifier with steady flow and deep bed operation can polish PLS to about 20-50 ppm.

d) Provide a PLS pond with large retention volume; e.g. one day minimum, (not recommended), three days is good practice while one week will be preferred by operators. Capability to remove settled or flocculated particulates from floor of pond will provide benefits by lower solvent losses.

e) Heat the clarified PLS before SX feed to remove super-saturation of silica and gypsum. Direct steam injection of PLS, or allow heating of the recycled, spent scrub liquor will tend to re-dissolve colloidal silica.

Silica polymerization and aggregation can be modified by anion or cation concentrations in PLS, (Terry 1983). Many solutes will form ion-pairs, e.g. Al & F, which can influence the relative effects of dissolved species including silica. Variable solute concentrations in PLS can complicate diagnosis of silica deportment and subsequent clarification and crud control measures.

TOWARDS THE FUTURE OF USX

Solvent extraction has dominated the uranium production industry for over 50 years, but during the hiatus in new uranium projects, the technology did not advance, (Ritcey 2003). USX is currently utilised in Australia two largest uranium mines, as shown in Table 1. The future of USX will reflect the expected trends of new projects that include lower grade ores, tighter controls on water supply or discharge. The trend of higher proportion of production via in-situ recovery will increase the application of ion exchange technology. To remain competitive USX will need to achieve greater stage recovery, lower solvent losses, and higher tolerance of salinity.

Table 1 Australian uranium extraction plants at March 2011,

Projects with feasibility study ongoing or complete are shown in italics, (after IAEA/OECD 2010).

Operation & Location Ranger, Northern Territory

Olympic Dam South Australia

Beverley South Australia

Honeymoon South Australia

Four Mile South Australia

Ore Category Unconformity Breccia Sandstone Sandstone Sandstone

Leach Chemistry Acid sulphate Acid sulphate, 3 g/L chloride

Acid sulphate, 4 g/L chloride

Acid sulphate, 5-8 g/L Cl

Acid sulphate

Recovery Technique Amine SX Amine SX Strong base IX Mixed SX Strong base IX



Nominal Capacity, tU 4660 3820 850 340 1150

Operation & Location

Kintyre, Western Australia

Wiluna, Western Australia

Yeelirrie, West Australia

Lake Maitland West Australia

Ranger Expansion

Ore Category Unconformity Surficial/ calcrete

Surficial/ calcrete

Surficial/ calcrete

Stockpiles

Leach Chemistry Acid sulphate Alkali carbonate

Alkali carbonate

Alkali carbonate

Acid sulphate

Recovery Technique DP DP DP DP Amine SX

Nominal Capacity, tU 2000 680 3000 850 TBC

REFERENCES

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Desson, M, Duchamp, C, Henry, E, Jouin, J P, and Michel, M, 1983. Industrial experience with pulse columns in the processing of uranium ores. AICHE Symposium Series. ISEC 1983, Vol. 80, No. 238, pages 162-169. Dobrowolski, H J, 1983. The Honeymoon Project – An alternative method of uranium mining. Proceedings of a symposium; Uranium Technology in South Australia, Australian Mineral Foundation, Adelaide. Edwards, C R, 1992. Uranium extraction process alternatives. CIM Bulletin, Vol., 85, No 958, pages 112 136. Faure, A, and Tunley T H, 1971. Uranium Recovery by Liquid-Liquid Extraction in South Africa. IAEA-SM-135/30 Symposium on the Recovery of Uranium, Sao Paulo, Brazil, pages 241-251. George, D R, and Ross, J R. 1967. Recovery of uranium from mine waters and copper ore leaching solutions. Panel Proceedings on Low Grade Uranium Ore, STI/PUB/146, IAEA, Vienna. Glasser, D, Arnold, D R, Bryson, A W, Vieller, A M S, 1976. Aspects of Mixer Settler Design. Minerals Science and Engineering, Vol. 8, No. 1, pages 23-45. Grinstead, R R, Shaw, K G, and Long, R S, 1958. Solvent extraction of uranium from acid leach slurry and solution. Proceeding of United Nations International Conference on the Peaceful Uses of Atomic Energy, Geneva, 1955, Vol. 8, Production Technology, pages 71-76. Gupta, C K, and Singh, H, 2003. Uranium Resource Processing. Springer Verlag, 2003, pages 172-175. Hall, S, and Reed, M, 1996. Cross contamination of ODC solvent extraction circuits. SME Annual Meeting, Phoenix. Hardy, H J, 1978. The chemistry of uranium milling. Radiochemica Acta, 25, pages 121-134. Hartmann, T, 2010. Crud treatment with 3 phase centrifuge in heap leach uranium process. Third International Conference on Uranium, Saskatoon, pages 741-749. Canadian Institute of Mining Metallurgy and Petroleum.. IAEA 1990. Technical Report No. 313, Manual on Laboratory Testing for Uranium Ore Processing, figure 27. IAEA 1993. Uranium Extraction Technology, International Atomic Energy Agency, Vienna. IAEA/OECD 2010. Uranium 2009; Resources, Production and Demand. Kertes, A S, 1966. The chemistry of the formation and elimination of a third phase in extraction systems. Solvent Extraction Chemistry of Metals, MacMillan, London, 1966, pages 377-400. LaBrooy, S, Spratford, D, Middlin, B, and Otto, K, 2009. Differentiating the honeymoon uranium process flowsheet, ALTA Conference, Perth Lyne, E G, Berryman, A B, Evans, C M, Sampat, S, and Jensen-Holm, H, 2002. Advances in NOx removal in smelter acid plants. International Seminar for New Initiatives in the Mining Sector, Chilean Copper Commission, Santiago 2002. http://www.topsoe.com/business_areas/flue_and_waste_gas/~/media/PDF%20files/Scr_denox/Advances%20in%20NOx%20removal%20in%20smelter%20acid%20plants.ashx



MacDonald, J B, Mattison, P L, and Mackenzie, J M W, 1978. Technical service in uranium solvent extraction. Practical problem solving in South Africa and the United States. Henkel Corporation Guidebook. Mayfield and Aker 2006. Honeymoon Uranium Project, Summary of Feasibility Study. http://www.sedar.com/CheckCode.do;jsessionid=0000Y-uard_JxXVHI0qodIl22P2:-1 Merritt R C, 1971. The Extractive Metallurgy of Uranium, Colorado School of Mines Research Institute. No. 389. Movsowitz, R L, Kleinberger, R, Buchalter, E M, and Grinbaum, B, 2000. Comparison of the performance of full scale pulsed columns vs. mixer-settlers for uranium solvent extraction. Uranium 2000 International Symposium on the Process Metallurgy of Uranium, MetSoc, pages 181-192. Canadian Institute of Mining Metallurgy and Petroleum. Moyer, B A, and McDowell, W J, 1981. Factors influencing phase-disengagement rates in solvent-extraction systems employing tertiary amine extractants. Symposium on separation science and technology for energy applications, Gatlinburg. Munyungano, B, Feather, A, and Virnig, M, 2008. Degradation problems with the solvent extraction organic at Rossing uranium. Proceedings ISEC 2008, September, Tucson, pages 296-274. Nice, R W, and Banaczkowski, M, 2000. Expansion of the ore treatment plant at Ranger uranium mines. Uranium 2000 International Symposium on the Process Metallurgy of Uranium, MetSoc, pages 163-180. Canadian Institute of Mining Metallurgy and Petroleum. Phillips, R, 2001. Honeymoon Uranium Project, Metallurgical Testing – Laboratory and Plant Trials, 1998 to 2000, Southern Cross Resources Pty Ltd, May 2001. Ritcey, G M, 1980. Crud in solvent extraction processing - A review of causes and treatment. Hydrometallurgy, Vol, 5 1980, pages 97-107. Ritcey, G M, 2003. State of the art and future directions in solvent extraction, Proceedings of the third international solvent extraction workshop. Digby, Nova Scotia, Canada. http://www.solvent-extraction.com/Proceedings/2003Proc.pdf Ritcey, G M. and Ashbrook R A, 1979. Solvent extraction : principles and applications to process metallurgy. Elsevier. Robinson, T, Sandoval, S, and Cook, P, 2003. World copper solvent extraction plants; Practices and design. Journal of the Minerals, Metals and Materials Society, Vol. 55, No.7, pages 24-46. Rosenbaum, J B, Borrowman, S R, and Clemmer, J B, 1958. P/501. Proceeding of Second United Nations International Conference on the Peaceful Uses of Atomic Energy. Geneva, 1958, Vol. 3, Processing of Raw Materials, pages 505-509 Royston, D, and Burwell, A, 1972. The design and performance of pump-mix and gravity flow mixer settlers. Australian Atomic Energy Commission, Lucas Heights. Saruchera, T, Jarvi, J, and Moldovan, B, 2010. Crud separation and SX Optimization at Key Lake. Third International Conference on Uranium, Saskatoon, pages 731-740. Canadian Institute of Mining Metallurgy and Petroleum.



Smith, H, 2005. Process simulation and modelling. In Developments in Mineral Processing. Chapter 5, Vol 15. Ed. Mike Adams, Elsevier. Soldenhoff, K and Davidson, J, 2005. Uranium recovery from highly saline in situ leach solutions by ion exchange. First Extractive Metallurgy Operators Conference, Brisbane, AusIMM. Terry, B, 1983. The acid decomposition of silicate minerals part II. Hydrometallurgical applications. Hydrometallurgy Volume 10, Issue 2, May 1983, Pages 151-171 Taylor, A, 2007. Review of mixer-settler types and other possible contactors for copper SX. ALTA Conference Perth. Thiry, J. and Roche, M, 2000. Recent process developments at the Cominak uranium mill. Uranium 2000 International Symposium on the Process Metallurgy of Uranium, MetSoc, Canadian Institute of Mining Metallurgy and Petroleum.


Fosterville Gold Mine Heated Leach Process

M Binks1 and P Wemyss2

1. MAusIMM, Processing Manager, Fosterville Gold Mine, McCormicks Road, Fosterville Vic 3557. Email: [email protected]

2. Processing Manager, Stawell Gold Mine, Leviathan Road, Stawell Vic 3380. Email: [email protected]

ABSTRACT

The Fosterville Gold mine is located approximately 20km to the East of Bendigo in central Victoria. The gold at Fosterville occurs as solid solution within disseminated arsenopyrite and pyrite. The Fosterville ore bodies contain various amounts of native carbon in the form of bituminous coal. This carbon (Non-Carbonate Carbon or NCC) has been the predominant mechanism for gold loss from the processing facility through”preg-robbing”.

Processing of the Fosterville ore is achieved initially through crushing, grinding and flotation to extract the sulphides. The sulphide concentrate is oxidized using Bacterial Oxidation, before being leached in a conventional CIL circuit.

A high portion of the native carbon (NCC) in the mine ore is naturally hydrophobic in nature, and subsequently reports to the flotation concentrate stream, and ultimately onto the CIL circuit. This NCC has a notable preg-robbing ability. Treatment of Black Shale ores, which have elevated NCC levels, has historically resulted in CIL recoveries as low as 35%, with around 60-80% of the gold loss from the leach circuit attributed to preg-robbing.

Standard technologies for dealing with a carbonaceous leach feed have been trialed with limited success. The discovery of the significance of heat on leach recoveries, triggered extensive testwork on a range of leach feed and tails samples. Pilot plant testing demonstrated that “Heated Leaching” of the CIL tails was the best process with an average recovery increase of 7.5% being achieved.

Following a successful pilot study, an engineering feasibility study was completed by Minerva Engineering, and the project economics determined. With strong recovery gains evident and a predicted project payback of 1-1.5 years, approval for the installation of the full scale plant was granted. Detailed design commenced in September 2008 with installation commencing in January 2009. The circuit was successfully commissioned in April 2009.

Following commissioning, the Heated Leach circuit has achieved recovery gains of 4-14%, and has proven itself to be a significant contributor to the overall plant performance.

INTRODUCTION

Fosterville Gold Mine is wholly owned by Northgate Minerals, a Canadian company. Mining at Fosterville has taken place intermittently since 1894. Contemporary exploration and heap leach operations commenced in the 1980’s and up to 2001 produced a total of 240 000 troy ounces (oz) of gold from these operations. Following the completion of a successful deeper drilling program in 2001/02, a detailed feasibility study was undertaken based on mining and processing refractory ore at a nominal rate of 800,000 metric tonnes


per annum. Engineering for the ore processing plant commenced in November 2003 with plant construction commencing in March 2004. Installation was completed and the first sulphide gold bar was poured in May 2005.

The Fosterville mineralization occurs within Lower Ordovician sediments comprising of interbedded sandstones, siltstones and shales. The predominant feature in the area is the Fosterville Fault System, a north-south striking, steep westerly dipping reverse fault comprising of numerous sub parallel faults. Gold is typically located in disseminated arsenopyrite and pyrite forming a halo to veins in a quartz – carbonate veinlet stockwork, which is in turn controlled by late brittle faults. The arsenopyrite occurs as fine-grained acicular needles with no preferred orientation. The disseminated pyrite associated with gold mineralisation occurs as crystalline pyritohedrons (Hitchman et al).

The Fosterville ore bodies contain various amounts of native carbon in the form of bituminous coal. This carbon (referred to as Non-Carbonate Carbon or NCC) occurs through hydrothermal alteration and as a significant sedimentary structure along the Fosterville fault line. The carbonaceous minerals in the ore are the predominant mechanism for gold loss from the processing facility through”preg-robbing”.

Processing of the Fosterville ore is achieved through a simple single stage jaw crushing circuit followed by a SAG mill. The sulphides are separated using a 3 stage flotation circuit. The sulphide concentrate is oxidized using bacterial oxidation, with the residue washed through a counter current decant (CCD) circuit, before being leached in a conventional CIL circuit. The tails from the CIL circuit are heated (Heated Leach) to recover a portion of the preg-robbed gold, before being pumped to a CIL residue storage dam. The flotation residue and the neutralised liquor from the CCD circuit are combined and pumped to an In-Pit residue storage facility.

A high portion of the native carbon in the mine ore is naturally hydrophobic in nature, and subsequently reports to the flotation concentrate stream, and ultimately into the CIL circuit. This NCC has been shown to have a high a preg-robbing ability, with gold loss of 6 to 7% from even the cleanest of ores. Treatment of black shale ores, which have elevated NCC levels, has historically resulted in CIL recoveries as low as 35%, with around 70-80% of the gold loss from the leach circuit attributed to preg-robbing (Wemyss, 2007).

Defining Fosterville “Preg-robbing”

Both plant and laboratory leaching demonstrate that the leach rate of the gold in oxidized Fosterville ores is extremely fast. Kinetic leach testing however, demonstrates a continued drop in solid grades over a 48 hour period and indicates a steady release (desorption) of weakly preg-robbed gold back into solution for adsorption onto activated carbon. The rate of gold desorption from the preg-robbing NCC is, as expected, solution grade and time dependent.

Diagnostic analysis (Table 1) and size by size assays (Table 2) provides an indication as to which size fraction the gold reports, and mechanism by which it is reporting. The historical average level of preg robbed gold in CIL tails is approximately 62%, with 36% locked in sulphides and the balance in silicates or precipitates.

Analysis of the size by assay shows that on average 70% of the gold and 90% of the native carbon is below 9 micron in size. The fine nature of the gold and native carbon limits the ability to study the surface of the mineral particles to more firmly establish the


gold/NCC association. The gold/NCC association is subsequently indirectly determined through diagnostic testwork.

Figure 1, is a plot of the gold and NCC levels obtained from composite samples taken over a two year period. It can be clearly seen that a reasonably strong correlation exists between the level of native carbon and gold within the CIL residue stream.

Table 1 CIL tail diagnostic analysis.

Diagnostic Average

CIL gold feed grade (g/t) 58.5

CIL tail grade gold grade (g/t)

8.9

Preg-robbed gold (%) 61.7

Native Carbon (NCC) (%) 1.2

Gold grade (g/t) 5.5

Sulphide locked gold (%) 35.7

Sulphide sulfur (%) 0.9

Silicate locked gold (%) 2.5

Table 2

Average elemental and weight distributions within the CIL tails stream.

Size CIL tail distributions

Wt (%) Au (g/t) S2- (%) NCC (%)

+ 45 micron 12 14 47 3

+ 20micron 8 10 35 2

+ 9 micron 8 4 1 5

+ 2 micron 30 26 1 46

< 2 micron 42 46 16 44

Total (%) 100 100 100 100


Figure 1. CIL tail gold grade versus %NCC.

Combating “Preg-robbing”

During the prefeasibility study for the Fosterville project, the presence of native carbon and losses through preg-robbing were identified, but the degree to which the ore bodies contained native carbon was underestimated. During the pre-feasibility, initial tests on the carbonaceous samples used the standard method of Kerosene blanking in a bid to address the pre-robbing loss, however this proved relatively unsuccessful.

Following commissioning of the Fosterville processing facility, CIL gold recoveries struggled to achieve the levels predicted in the prefeasibility study. Although leach gold losses were initially exacerbated by low oxidation levels, the loss of gold to preg-robbing was identified as the key recovery influence.

A myriad of testwork was conducted in a bid to establish a treatment method, including but not limited to:

Elevated Carbon Concentrations; Low Density Leaching; Kerosene Blanking; Synthetic Blanking; Pressure Oxidation; Thiosulphate Leaching; Chlorination; Thiourea Leaching.

Outside of the treatment of the leach feed to combat preg-robbing, other tests focused on the flotation circuit and the rejection of native carbon at this stage of the process. Chemical suppression was unsuccessfully trialed several times in the plant. Through the trials however, the use of Guardisperse (naphthalene sulphonate) was found to assist in the management of tenacious carbonaceous froth experienced during the blending of black shale ores. Gardisperse continues to be used periodically to assist in managing tenacious froth and not specifically NCC rejection.


Due to the fine nature of the native carbon and its lower density relative to the sulphide minerals, physical separation from the secondary circuit flotation concentrate was assessed. The use of 1” Mozley cyclones was evaluated (Binks, 2006) and ultimately applied to remove fine native carbon from the stream. This project proved successful in its own right with 60% carbon rejection achieved with an associated 5% gold loss from the flotation circuit. The clean cyclone underflow produced vastly improved leach recoveries. However this circuit was superseded with the introduction of the Heated Leach process and the need to blend the high NCC ores to maximize flotation gold recovery.

Roasting is seen as the best method for the treatment of Fosterville flotation concentrates. However, the location of the site in a highly populated area, coupled with the local weather conditions meant that the process was considered environmentally unsound and therefore not a viable proposition.

A Left Field Approach

Standard technologies and known methods for dealing with the Fosterville carbonaceous leach feed have been trialed with limited success. The lack of success from laboratory tests provided evidence that a portion of the gold must already be associated with the native carbon prior to entering the leach circuit. Mozley cyclone classification testwork on flotation concentrate (Binks, 2006) demonstrated that a gold/carbon association is not present within the flotation concentrate and subsequently indicates that the “preg-robbing” is initiated in the bacterial oxidation circuit.

This understanding moved the focus of testwork to reversing or breaking the gold/NCC association developed within the bacterial oxidation circuit. The application of heat to the pulp was tested as a method to break the gold/NCC bond. Initial leach temperatures of 40 degrees Celsius were adopted with the view to use waste heat from the bacterial oxidation circuit which operates at a similar temperature. “Heated Leach” tests conducted on CIL feed at 40 degrees Celsius immediately showed benefits with a step increase in CIL recoveries over both plant and previous testwork results and provided the foundation for further work.

Heated Leach Pilot Plant Studies

To fully evaluate the Heated Leach process, establish a potential processing circuit flowsheet and to enable identification of any idiosyncrasies of operating such a circuit, a pilot plant was constructed. Como Engineering of Perth Western Australia were engaged to construct the plant based on site design. The pilot plant (Figure 2) consisted of 2 trains of 6 tanks. Each train was different in volume to enable a comparison of residence time and to conduct two tests simultaneously.

Heating of the pulp was achieved by circulating hot water through jackets around the tanks. Each tank was fitted with a basket to retain the carbon and allow for easy carbon transfer. The circuit was also set up with chemical, oxygen and air dosing to the tanks.

Standard carbon management in the pilot plant was designed around having 20g/l carbon concentration in each tank, with basket movements in each train timed to approximate carbon movement rates and dwell times achieved in the existing CIL plant. Free cyanide levels were kept at historical plant levels and pH was kept in line with the plant. Any variation from the above was performed specifically to test different treatment scenarios.


Sampling of streams was performed every 6 or 12 hours in order to give good variation around the 9hr and 18 hour basket movement regime. This way assay results were available from the start, middle and end of basket dwell duration in order to determine variation effects.

Figure 2. Heated Leach Pilot Plant.

The Heated Leach pilot plant was operated for a period of 4 months on a continuous roster to evaluate the effects of:

Heated Leach process on CIL Feed; Heated Leach process on CIL Tails; Various Temperature Profiles; Carbon addition to the circuit including location, activity and rate; Cyanide addition; Caustic addition.

Pilot plant testing clearly demonstrated that Heated Leaching of the tailings from the CIL plant was the best method for increasing overall gold recovery from the ore, with an average recovery increase of 7.5% being achieved. The pilot work also identified that the recovery increase through Heated Leaching is possible without any further additions of cyanide, and can be performed in the absence of air or oxygen injection into the pulp (Wemyss, 2007).

Pilot studies were later extended to evaluate the ability of the process to recover gold from historic leach tails which had an average gold grade of 10 g/t. The testwork indicated that 35-50% of the gold within the tail residue could be recovered, further supporting the implementation of the process. Additional capacity was factored into the design to allow for reclaimed tails to be processed con-currently.


Board Approval for the project was received in October 2008. Minerva Engineering of Melbourne were engaged to complete the detailed design. Site construction commenced in January 2009 with commissioning underway by April of that year.

Full Scale Operation

The six stages used during the pilot plant studies on the CIL tail were ultimately applied to the full scale plant. The initial three stages of the Heated Leach circuit are operated up to 70 degrees Celcius with the remaining three stages cooled to aid the adsorption of gold from solution.

Heating of the CIL tails stream is achieved through inline injection of steam which is provided by a 4 megawatt LPG fired boiler. Tanks on the first three stages are rubber lined and insulated to reduce radiant heat loss. Stainless steel lids are fitted to four of the six stages to retain heat and reduce steam entering the operator working zone on the top of the tanks.

The detailed design identified the necessary cooling of the secondary stages (tanks 4-6) could be achieved through simple aeration of the tanks. This cooling method however, proved unsuccessful during the commissioning phase and ultimately cold recycled water from the CIL residue dam was adopted to achieve the temperature adjustment, at the sacrifice of density and tank residence time.

In stages 1-3, recessed impeller pumps are used for carbon transfer to prevent unnecessary cooling of the pulp that would occur if using air lifts. Air lifts are used for carbon transfer in stages 4-6 where lower operating temperatures are required. In all cases, the carbon advance slurry is screened to allow the carbon only to progress forward and the slurry to return to the tank from which it came.

Barren carbon is continuously added to both stages 1 and 6 in the circuit. The loaded carbon from the Heated Leach circuit is recovered to a carbon column, regenerated and then returned to the CIL circuit. Figure 3 below depicts the carbon movement at nominal rates and grades.

Figure 3. Carbon Movement through the Heated Leach and CIL circuits.

The circuit operates continuously with an availability of 95-98%. In 2010, the Heated Leach process (Fig 4.) provided an average recovery gain of 7.4%. This is deemed an excellent result particularly given the level of circuit downtime and reduced efficiency associated with poor availability of the steam boiler. During the year, a peak monthly gold


recovery gain of 14% was achieved during the processing of a blend containing high NCC levels.

Figure 4. Fosterville Gold Mine Heated Leach circuit

Conclusion

The Heated Leach circuit is the most notable addition to the Fosterville processing facility that has been made since the operation of the plant commenced. The Heated Leach circuit contributes significantly to improved process economics and the long term future of the site. Continued work on the optimization of the process and improvement of the equipment availability will further enhance the benefits obtained.

Acknowledgements

The authors would like to acknowledge Northgate Minerals for the support and confidence in firstly piloting the process and then establishing a full scale operation.

References

Binks, M, 2006. Cyclone Classification Report February 2006, Fosterville Gold Mine.

Hitchman, S, Holland, I, and Evans, B, 2008. Technical Report on Fosterville Gold Mine, Victoria Australia.

Wemyss, P, 2007. CIL Tail Heated Leaching, Fosterville Gold Mine.



Leach Residue and Pregnant Liquor Separation – Process and Capital Comparison of Counter-Current Decantation and

Counter-Current Washing with Vacuum Filtration

R Klepper1 and P McCurdie2 1. Global Manager Hydromet, FLSmidth, 7158 S. FLSmidth Drive, Midvale Utah 84047-5559, USA.

Email: [email protected] 2. Filtration Process Manager, FLSmidth, 7158 S. FLSmidth Drive, Midvale Utah 84047-5559, USA.


ABSTRACT

Separation of leach residue suspended solids from pregnant liquor solution is an important step in hydrometallurgy flowsheets. Alternative flowsheet options are counter-current decantation (CCD) thickeners or counter current washing (CCW) horizontal belt vacuum filters. Extensive sedimentation and vacuum filtration simulations have been completed by FLSmidth on leach discharge slurries to indentify equipment capacity and performance.

These data were used to compare Capex and Opex for CCD versus CCW alternatives. In addition, Capex and Opex for large horizontal belt vacuum filters with filtration areas greater than 250m2 are compared to current state of the art sizes of 150m2. This paper presents this information and discusses various advantages and disadvantages.

INTRODUCTION

Filtration with washing can recover soluble metals using less water than thickening and washing. However the overall filtration rate must be at a rate where the capital cost and operating cost is less than thickening and washing in a CCD circuit.

Solid-liquid separation equipment selection for separating leach residue suspended solids from pregnant liquor solution (PLS) in leach discharge slurry is controlled predominately by the particle size distribution of the suspended solids. The quantity of particles less than 37 micron diameter controls the rate of gravity separation and permeability of the packed bed of suspended solids or filter cake formed during filtration. The contents of the paper will identify the parameters and physical properties that must be considered during feasibility evaluations and in the detailed design of hydrometallurgical plants.

A review of the history of gravity thickeners and vacuum filtration is also presented to better understand existing plant design compared to future plant designs now being proposed. Examples of capital cost and operating costs will also be presented as a basis for various options.

THICKENER HISTORY

One item has dominated thickener development - flocculation! Fig. 1 (Klepper ALTA 2009) is a pictorial timeline of thickener development. The development of polymer chemistry has lead to the reduction of the diameter of thickeners and allowed throughput to increase significantly. Originally thickener performance was controlled by gravity and the settling rate of the smallest particle. These thickeners are defined as conventional thickeners.

In the 1970’s flocculants were added into thickeners at various locations and collected the finest particles into larger groups or floccules increasing the gravitational settling rate. Some design philosophies proposed the feed slurry exit the feedwell at a position under suspended solids within the thickener to enhance capture of unflocculated particles. These smaller diameter thickeners are defined as high rate thickeners.



In the 1980’s knowledge of optimizing the use of flocculants was learned along with combining gravity and compression forces created by the height of flocculated suspended solids within the thickener. These smaller and taller thickeners are defined as deep cone paste or high density thickeners depending on the rheological properties of the underflow slurry.

Fig. 1 CCD Thickener Development History Timeline

Optimizing flocculation was achieved by dilution of the feed slurry into the thickener feedwell to some optimum suspended solids concentration that allows flocculants to be evenly dispersed into the feed slurry. Another phenomenon that occurs at the optimum feedwell suspended solids concentration is larger particles become part of the floccules acting as ballast. Consequences are a measurable maximum settling velocity and a homogeneous particle size distribution within the thickener contents.

Additionally CFD modeling of various feedwell designs, pioneered in the AMIRA International project “P266 Improving Thickener Technology” (AMIRA 2011), identified feedwell geometries that created more shear than other geometries. Shear or turbulence breaks floccules in the feedwell resulting in lower settling velocities and lower suspended solids concentrations in the underflow slurry. Fig. 2 illustrates a CFD model of the consequence of thickener feedwell design changes minimizing shear and floccule destruction.

Fig. 2 CFD Modeling of Feedwell Development

Paste and high density thickeners are also designed with physical features, that can be seen in Fig. 3 (Klepper COM2009), to enhance liquor release from the slurry within the thickener:

sidewall height is increased to create more compressive force, add slurry residence time for liquor release and creates resistance to short circuiting of liquor into the underflow discharge,



rotating vertical members or pickets is used to alter the permeability of the slurry within the thickener allowing more rapid liquor release,

rake geometry is changed to minimize cross-sectional area reducing resistance or torque development,

tank floor geometry is changed to a steeper slope and a discharge cylinder is included to allow discharge of a homogeneous slurry exhibiting non-Newtonian rheological properties such as an apparent yield stress of >300 Pascals, and

greater installed unit torque in the rake drive is used so the thickener can operate at higher suspended solids concentrations in the underflow slurry without stopping.

Fig. 3 High Density Thickener Features

Fig. 4 illustrates the impact on thickener diameter and CCD footprint at the same throughput for conventional, high rate and high density thickeners determined for a hydrometallurgical copper project.

In summary high density or paste thickeners can produce underflow slurries with higher suspended solids concentrations in smaller volumes that can be controlled automatically.

70m Conventional Thickener

50m High Rate Thickener

31m High Density Thickeners

Fig. 4 Comparison of Thickener Diameter by Thickener Type

VACUUM FILTER HISTORY

Vacuum drum filters were initially used for solid-liquid separation and washing of various leach discharge slurries because soluble metal recovery as these can use mush less water than in CCD thickener circuits. However the capacity per filter is limited by the size (maximum area 140m2) and



geometry of the drum thus restricting the size of each filtration zone. Fig. 4 illustrates a vacuum drum filter and maximum size of filtration zones during one rotation is also listed below:

filter cake formation zone composed of leach residue suspended solids captured on the filter cloth as the PLS flows through,

filter cake washing zone to displace the residual PLS from the pores of the filter cake,

filter cake dewatering or drying zone where air is drawn through the pores of the filter cake, and

filter cake discharge zone before the sequence is repeated with the next revolution.

Fig. 5 Drum Filter & Filtration Cycle Zones

Again, polymer chemistry played an important role in the expansion of vacuum filtration capabilities leading into rubber belt technology advancements and as a result, the first horizontal belt vacuum filter (HBVF) was invented in the 1960’s. The advantages of the HBVF’s over drum filters are complete flexibility of the duration of all segments of the filtration cycle and greater capacity per filter. The rubber carrier belt and filter cloth rotate around two large rolls and the time the filtration area is under vacuum is infinite. The feed slurry does not require mechanical suspension because the slurry is poured on top of the filter cloth. Advances in understanding the mechanical design of HBVF components such as rubber carrier belts has allowed the total filtration area to increase to a maximum of 250m2/filter with a carrier belt width of 6.2m. The capacity per filter is therefore increased.



Fig. 6 Horizontal Belt Vacuum Filter Geometry

There are five filtration zones in a HBVF filter cycle:

filter cake formation zone or time (θf) it takes to form a given weight per area (w),

filter cake wash zone or time (θw) to takes for wash to pass through the filter cake,

filter cake dewatering or drying zone time (θd) it takes to obtain minimum residual moisture content of filter cake,

filter cake discharge zone or time to discharge filter (θdc) cake from the filter cloth, and

filter cloth washing zone where residual suspended solids and/or precipitated solids are removed from the filter cloth to maintain permeability.

The overall filtration rate is the weight per area of the dry suspended solids in the filter cake (w) divided by the sum of filter cake form time (θf) plus the filter cake wash time (θw) plus filter cake dry time (θd) plus cake discharge time (θdc).

The overall filtration rate is controlled by several variables which are illustrated in general in Fig. 7:

pressure differential or obtainable vacuum,

permeability of the filter cake or suspended solids particle size distribution,

permeability of the filter cloth,

quantity of liquor to be removed from the feed slurry to form the filter cake, and

quantity of wash liquor used to recover soluble metals.

Pressure differential is a variable due to the atmospheric pressure decrease as plant elevation increases. At sea level there is a theoretical 101 kPa pressure differential available and this reduces to about 84 kPa at 1500m elevation. The installed vacuum pump capacity controls how close to theoretical pressure differential can be achieved. The quantity of air the vacuum pump must remove is controlled by the permeability of the filter cake during the cake drying portion of the filter cycle. The permeability of the filter cake can increase significantly if the filter cake shrinks and cracks during the dry portion of the filter cycle.



Fig. 7 Parameters Controlling Overall Filtration Rates

Filter cake permeability can be maximized by concentrating the suspended solids in the feed slurry or by using flocculants to prevent fine particles from accumulating on the top of the filter cake thus creating an interfering thin layer of low permeability. The cake permeability must also be maintained and not allowed to shrink and crack between cake formation and cake wash or the cake liquor will by-pass the filter cake pores in favor of going through the cracks.

Filter cloth permeability must be maintained by providing continuous filter belt washing to remove suspended solids remaining on the belt after filter cake discharge. In addition intermittent acid washing of the filter cloth with dilute HCl must be provided to remove an accumulation of precipitated solids deposited on the filter cloth as a consequence of a change in solubility. The liquor immediately flowing through the filter cloth evaporates water due to the lower pressure on the vacuum side of the filter cake. Deposition of small quantities of precipitated solids onto the filter cloth is inevitable reducing the open area of the filter cloth or permeability so that more pressure drop occurs across the filter cloth than the filter cake thus reducing flow of liquor over time.

The quantity of pregnant liquor in the filter feed slurry must be minimized by using a thickener upstream of the HBVF to concentrate the suspended solids sufficiently so that the feed slurry can be evenly distributed across the width of the filter.

The quantity of wash liquor can be optimized by using counter current washing with multiple stages. Barren wash liquor is applied at the back end of the HBVF passing through the filter cake and collected as filtrate and reapplied to the top of the filter cake repeatedly. The reason for counter current washing is the fact that the remove of liquor containing soluble metals is a combination of displacement and diffusion of soluble metals from non linear channels inside the filter cake. The amount of diffusion is a direct function of the quantity of flocculated fines or particles less than 10 microns in the filter cake.



THICKENER CCD SCOPE OF SUPPLY

Now that a brief history and review of operating parameters has been completed, the Scope of Supply associated with the cost to install CCD thickeners will be discussed. Fig. 8 illustrates a typical CCD circuit used for a copper hydrometallurgy flowsheet. Fig. 9 illustrates a typical P&ID for a CCD Thickener.

Fig. 8 CCD Circuit PDF

The CCD circuit system includes multiple thickeners in series including the following components:

thickener (including tank, feed system and rake mechanism),

flocculant preparation and control of distribution system to each thickener,

flocculant dilution and distribution system into each thickener feedwell,

underflow pumps and control,

overflow tank and pumps (optional versus gravity flow through the CCD circuit)

interstage mix tank and agitator, and

wash liquor supply.

The dominant variables affecting the installed cost of any CCD thickener circuit are the number of stages and thickener diameter. These two variables define the land area of the plant or footprint required for the equipment. Thickeners smaller than 50m diameter can be designed as elevated tanks with a bridge spanning the tank diameter with the rake drive mechanism mounted on the bridge. Thickeners larger than 50m are on-ground tanks with a central column supporting the rake drive mechanism and a bridge spanning from the perimeter to the rake drive for operation and maintenance access.



Fig. 9 CCD Thickener P&ID

Leak detection and secondary containment of thickener contents with an elevated tank is positive and relatively easy to design. However leak detection or secondary containment of thickener contents for an on ground tank requires a design solution specific to the materials of construction.

The smallest diameter thickeners available for a CCD circuit should be based on optimized flocculation occurring at some optimum suspended solids concentration of slurry into the feedwell. The current state of the art uses High Density thickeners to obtain the smallest foot print. However, many major mining companies go through a conservative risk versus current state of the art analysis and may not know of differences between High Rate and High Density thickeners. Because of this lack of understanding, consequence can be using older “proven technology” in High Rate thickeners with a larger foot print rather than the current “state of the art” High Density Thickeners with their reduced footprint.

In hydrometallurgical applications the next item affecting capital cost is the materials of construction of the thickener rakes and tank to provide corrosion resistance to the acidic PLS. The choices are suitable alloy metals or carbon steel or concrete lined with one of several suitable non metallic linings or coatings.

Selecting the proper material of construction requires defining the physical chemical conditions in the CCD circuit while keeping in mind that due to washing, the conditions are most severe in CCD1 and less severe in CCDn. This means there is potential to change materials of construction at CCD3 or CCD4 where the concentrations of corrosive components in the liquor are lower. The physical chemical variables are:

temperature,



acid concentration (g/l H2SO4 or HCl),

pH,

redox potential,

ionic chemical composition (g/l Al+3, Ca, Cl-1, Co+2, Cu+2, Fe+2&+3, Mg+2, Mn+?, Ni+2, Si+6, Zn+2), and

organic content (g/l).

Temperature is a fundamental parameter to define suitable alloy metals or non metallic linings. Temperatures greater than about 90oC require consideration of both personnel safety and corrosion resistance. Personnel safety can be achieved by installing a roof over the thickener to contain water and acid vapors and to prevent injury, insulating tanks and pipe wherever contact is a possibility. Most process designs recover heat after leach and lower the temperature of the leach discharge slurry to the CCD circuit to 50-70oC. However, if the temperature is 100-110oC then brick or fiberglass reinforced plastic lining of the thickener tanks are some options available. Brick acts as an insulator and is actually placed over a rubber membrane that provides the corrosion protection.

Generally at temperatures less than 90oC suitable nonmetallic lining systems can be rubber sheets, fiberglass reinforced plastic panels, flaked glass filled resin coatings, or thermal plastic sheets. The effectiveness of almost all non metallic lining systems is dependent on proper surface preparation of the tank structural material and application of the lining under environmental conditions such that the properly prepared surfaces do not degrade.

By partially neutralizing the residual acid in the leach discharge liquor, some process designs alter the leach discharge liquor chemistry favorably for less costly materials of construction to be used.

The choice of metallic materials of construction for thickener rakes and tanks are stainless steels, duplex stainless steels, lean duplex stainless steels, super duplex stainless steels, austenitic stainless steels, and other specialty alloys. Selection should be based on measured corrosion rates at the physical chemical conditions existing in the PLS recovery system.

A lesser cost but paramount to optimizing the flocculation on minimizing thickener diameter is the feed dilution system. Several feed dilution options exist with varying degrees of flexibility to maintain the optimum suspended solids concentration in the feedwell slurry. There are:

self diluting feedwells designed based on relatively constant differential densities of feed slurry, feedwell slurry and overflow liquor producing a relatively constant dilution rate,

systems that use kinetic energy of the feed slurry to create an eductor affect drawing thickener overflow liquor into the feed slurry as the diluent at a relatively constant dilution rate, and

low-head pumps installed in the feed pipe with variable speed drives that pump the appropriate volume of overflow diluent into the feed slurry to obtain practically any optimum suspended solids concentration in the feedwell for all types of ore.

Thickener diameter also is an important parameter affecting the selection of the rake drive and the torque required to rotate the rakes. High Density thickeners will have the smallest diameter compared to High Rate or Conventional thickeners. Torque can be defined as T=KxD2 where K is unit torque factor (N/m) and D is thickener rake diameter (m). Minimizing the diameter of the thickener means less torque is required to rotate the rakes. High Density thickeners are designed to produce higher suspended solids concentrations in the underflow slurry resulting in more resistance for the rakes to rotate through the slurry. For the thickener rakes to continue to rotate in more dense underflow slurry a drive with greater K value must be utilized than would be necessary in a High Rate thickener.



Flocculant preparation for CCD circuits is required to supply sufficient flocculant for each thickener in the CCD circuit to function correctly. The flocculant preparation system consists of a storage area to store un-hydrated flocculant keeping it dry prior to use. A conveying system is usually used to add a quantity of flocculant into an agitated tank of water so the long flocculant molecules can totally dissolve into a neat solution at a concentration of 0.5-2 wt%. Once the flocculant is totally dissolved, the flocculant solution can be metered to each thickener using a variable speed positive displacement pump at a dose sufficient to properly flocculate the suspended solids.

The neat flocculant solution is viscous and to allow efficient blending of the flocculant with the feedwell slurry the viscosity must be reduced. Viscosity reduction of neat flocculant solution is achieved by blending overflow liquor with the neat flocculant solution in a static mixer. Overflow liquor is used rather than water to avoid diluting the pregnant liquor and to conserve fresh water usage. The overflow liquor can be pumped from within the thickener using a submersible pump or the overflow liquor can be pumped external to the thickener using a horizontal pump. Taking the dilution liquor from within the thickener minimizes the actual overflow rate and carryover of suspended solids in the overflow liquor.

Centrifugal underflow slurry pumps with variable frequency drives are typically used to control the inventory or residence time of suspended solids in each thickener and transfer the underflow slurry to the next CCD interstage mix tank. The inventory of suspended solids is measured indirectly by measuring the hydrostatic pressure at the bottom of the thickener and/or by measuring the height of the bed of suspended solids. In some CCD circuits the plant water balance requires minimizing the quantity of PLS dilution and double mechanical seals may be used rather than packing gland seals. The quantity of seal water can be reduced to zero if the quality and quantity of underflow slurry allows the use of peristaltic hose pumps. The quality depends on rheology and particle size distribution.

Overflow liquor can be pumped using centrifugal pumps with variable frequency drives or flow by gravity from each thickener into the interstage mix tank for blending with underflow slurry. The plant topography and geotechnical composition define the cost of civil excavation for the CCD circuit and whether gravity flow is feasible.

The fundamental design basis of a CCD circuit is mixing or blending the liquor in the underflow slurry of the upstream CCD stage with the overflow or wash liquor from the downstream CCD stage so a homogeneous soluble metal concentration is obtained. The interstage mix tank is where this occurs. The mixing task in the interstage mix tank is defined as blending, but the blending efficiency or degree of homogeneity cannot be 100% because there is always some short circuiting in a single tank. It is more difficult to blend streams with great differences in rheology. HDT underflow slurry has greater suspended solids concentrations exhibiting non-Newtonian rheological characteristics than produced by HRT’s. Therefore the interstage mixing system must be designed to provide adequate shear to disperse the underflow slurry into the overflow liquor so that and efficiency or degree of homogeneity of about 95% can be achieved.

The wash liquor rate or wash ratio, defined as mass of wash liquor per mass of leach residue suspended solids, should be controlled at a rate proportional to the suspended solids feed rate into the CCD circuit. The wash liquor quality used in the CCD circuit should have minimal dissolved pay metals to ensure maximum recovery of the pay metals.

HORIZONTAL BELT VACUUM FILTER SCOPE OF SUPPLY

The scope of supply of CCW vacuum filter systems will be discussed next.

Filtering leach discharge slurry requires the use of at least two filters in parallel to be able to maintain production during maintenance or to be able to have adequate filter area to filter total



production. The filtration system includes multiple HBVF’s in parallel and each HBVF includes the following components:

filter (including rubber carrier belt with roller deck, filter cloth support rolls, vacuum pan with filtrate collection pipes and belt drive as seen in Fig. 10),

filter feed slurry thickener (required to minimize liquor concentration in leach slurry and minimize separation of fine particles during filter cake formation),

feed slurry storage tank and distribution system,

flocculant preparation and distribution system (optional),

filtrate receivers and pumps (single wash or counter current wash scheme),

vacuum pumps, seal water, and moisture traps,

cloth wash pump for suspended solids removal,

cloth wash pump for scale removal, and

filter cake discharge handling system.

Fig. 10 HBVF Components

HBVF’s can be designed with filtration areas under vacuum ranging from 11m2 up to 250m2 as illustrated in Fig. 10. The filter consists of a frame to mount numerous rollers that support, guide and drive the rubber carrier belt during filtration and return to the front of the filter. Additional rollers are used to support and guide the filter cloth at discharge and return to the front of the filter. Also attached to the frame is a vacuum pan to collect filtrate and air, rubber wear belts to produce a seal between the rubber carrier belt and the vacuum pan, a feed distribution box, wash liquor distribution boxes and cloth cleaning mechanisms. The rubber carrier belt is driven by either one or two variable speed drives so the cake thickness or filtration rate can be controlled to match the physical/chemical characteristics of the solids and liquid in the feed slurry.



Fig. 11 Largest and Smallest HBVF

A filter feed slurry thickener is used to reduce the liquor content minimizing the time it takes to form the filter cake. Filter cake form time is a major component of overall filtration rate. An additional benefit of thickening the feed slurry is elimination or minimized segregation of fine particles from coarse particles which would inevitably lower the permeability of the filter cake. For operation control it is very important to have a filter feed tank between thickener underflow discharge and filter feed pump. The storage tank provides surge capacity for the thickener underflow slurry discharge rate to be independent of the filter feed slurry rate. The storage tank also provides time to stop either piece of equipment for short periods of time without stopping both or even total plant production.

Filter feed slurries can also be flocculated also to eliminate segregation of fine particles from coarse particles and so avoiding a reduction in filter cake permeability created by the layer of interfering fine particles. However, the cost of flocculant can produce an uneconomical Opex.

The feed slurry distribution system is important as it is necessary to have equal filter cake thickness across the width of the filter so wash liquor flows through the filter cake evenly and does not short circuit through thinner filter cake areas with less resistance.

Filtrate receivers as illustrated in Fig. 12 collect filtrate from the manifold connecting the filter cake formation zone, filter cake wash zones and filter cake dry zone and are used for preliminary separation of liquid from air. The filtrate from the filter cake form zone is PLS and the filtrate from the wash zone is combined with the PLS or if there are multiple wash stages can be reapplied to the filter cake.



Fig. 12 HBVF Layout

Double mechanical seals are used on some filtrate pumps to minimize PLS dilution. These filtrate pumps require a barometric leg equal in hydrostatic height to the vacuum used in the filter to prevent flashing of PLS that may lead to seal damage or scale formation. In processes where the dissolved solids concentration is not near saturation in the filtrate and dilution of the PLS is not a concern, barometric legs are not required as there are open filtrate pumps that operate at reduced suction head.

Vacuum pumps evacuate air from within the filter creating the pressure differential for flow of liquor through the filter cloth and formed filter cake, flow of wash liquor through the filter cake and flow of air through the filter cake to remove residual moisture or dry the filter cake. A moisture trap as illustrated in Fig. 12 must be installed between the filtrate receivers and the vacuum pump to ensure no process liquid goes through the pump causing damage. Vacuum pumps require seal water for the shaft seal, temperature control and to lubricate the rotors. Temperature of the PLS must also be considered with regard to the amount of water evaporation that will occur at the reduced pressures. Additional water may be required to condense water vapor so the pumps remove more non-condensable air rather than water vapor.

Cloth washing is required to maintain permeability of the filter cloth, prevent suspended solids from accumulating under the filter as the filter cloth returns to the front of the filter and to keep the filter cloth support rolls clean so the cloth remains in the desired location. In an effort to maintain a water balance in the plant and conserve fresh water use, seal water from the vacuum pumps can be used for cloth wash water and ultimately as filter cake wash water. Seal water from the vacuum pumps can be collected in a tank. A pump is used to create pressure for the cloth wash water to go through spray nozzles positioned across the width of the filter cloth on both sides of the cloth after the filter cake has been discharged.

In almost all hydrometallurgical applications scale forms over time on the individual fibers of the filter cloth progressively restricting the flow of liquor until ultimately the filter cloth is totally blinded or impermeable, restricting the flow of liquor. At or before the filter cloth is “blinded”, the filter must be taken out of service and a separate acid wash applied to dissolve the scale to restore the filter cloth to near original permeability. This cloth wash liquor will contain sufficient acid concentration to dissolve the scale, therefore the wash liquor system must use suitable materials of construction for corrosion resistance. On an OH&S basis and to prevent injury the wash spray nozzles are inside an enclosure that the filter cloth passes through.



As an example large tonnage of leach residue require multiple filters. There are two rubber carrier belt widths (4.2m and 6.5m) available to determine the total number of filters. When there are more than four (4) filters required, a savings of about 15% lower Capex can be achieved by using fewer larger filters. Fig. 13 illustrates a layout of eight (8) 170 m2 HBVF with 4.2 m rubber carrier belts providing a total of 1360 m2 of total filter area. Fig. 14 illustrates a layout of five (5) 273m2 HBVF’s with 6.5 m rubber carrier belts providing a total filtration area of 1365 m2.

It should be noted that HBVF development has kept pace with other minerals processing equipment development into larger size to take advantage of economy of scale.

Fig. 13 136m2 Total Filtration Area Using 8x170m2 Filters 4.2m Wide

Fig. 14 1365 m2 Total Filtration Area Using 5x273m2 Filters 6.5m Wide

CAPEX & OPEX COMPARISON FOR CCD THICKENERS AND HORIZONTAL BELT VACUUM FILTERS



The outlines below lists the items to consider for capital expenses (Capex) and operating expenses (Opex) for a CCD thickener circuit and a horizontal belt vacuum filter installation.

CCD Thickeners Capex and Opex

1. Capex 1.1. Civil Site Preparation

1.1.1. Sufficient area for all tanks 1.1.2. Potential terracing for gravity flow of overflow liquid from stage to stage 1.1.3. Spill containment ponds

1.2. Equipment 1.2.1. CCD Thickener

1.2.1.1. Number of thickeners to recover pay metals 1.2.1.2. Type & size of thickeners to produce maximum underflow suspended solids

concentration at plant production 1.2.1.3. Materials of construction for submerged components and tank suitable for corrosion

resistance and potential scale removal 1.2.1.4. Leak detection 1.2.1.5. Rake drive with sufficient torque suitable for underflow slurry physical properties,

variable speed and torque transmitter 1.2.1.6. Rake lift for escape from abnormal operation with rake height transmitter 1.2.1.7. Feedwell slurry dilution system for efficient flocculation (low-head dilution liquor

pump optional) 1.2.1.8. Flocculant dilution system

1.2.1.8.1. Dilution liquor pump and shaft seal 1.2.1.8.2. Dilution liquor / flocculant static mixer 1.2.1.8.3. Diluted flocculant distribution piping and valves

1.2.1.9. Suspended solids bed level transmitter 1.2.1.10. Suspended solids bed hydrostatic bed mass transmitter

1.2.2. CCD Thickener Ancillaries 1.2.2.1. Flocculent supply system

1.2.2.1.1. Flocculent dry storage 1.2.2.1.2. Flocculant dissolution tank and mixer 1.2.2.1.3. Flocculant solution storage tank 1.2.2.1.4. Flocculant metering pump 1.2.2.1.5. Flocculant dose control

1.2.2.2. Feed slurry mass measurement 1.2.2.2.1. Feed slurry density transmitter 1.2.2.2.2. Feed slurry flow transmitter

1.2.2.3. Underflow slurry pumps variable speed, operating and stand-by 1.2.2.3.1. Mechanical seal water supply system 1.2.2.3.2. Underflow slurry piping 1.2.2.3.3. Underflow slurry density transmitter 1.2.2.3.4. Underflow slurry flow transmitter 1.2.2.3.5. Underflow slurry density control 1.2.2.3.6. Suspended solids inventory control

1.2.2.4. Interstage mixing system tank and agitator 1.2.2.5. Overflow liquor pump variable speed, operating and standby (optional without

gravity flow) 1.2.2.5.1. Overflow liquor pump sump 1.2.2.5.2. Pump sump level transmitter

1.2.2.6. Interconnecting pipe and pipe rake steel 1.2.2.7. Stairs and platforms between and around thickener tanks

2. Opex 2.1. Power

2.1.1. Rake drive motors



2.1.2. Rake lift motors (intermittent and emergency power) 2.1.3. Feedwell dilution liquor low-head pumps (optional) 2.1.4. Flocculant metering pumps 2.1.5. Flocculant dilution liquor pumps 2.1.6. Underflow slurry pumps 2.1.7. Overflow liquor pumps (optional without gravity flow) 2.1.8. Interstage mixers 2.1.9. All controls

2.2. Consumables 2.2.1. Flocculant (100-400 g/t total) 2.2.2. Process water

2.2.2.1. Wash liquor 2.2.2.2. Flocculant dissolution 2.2.2.3. Pump seal water

2.2.2.3.1. Underflow slurry pump 2.2.2.3.2. Overflow liquor pump (optional without gravity flow) 2.2.2.3.3. Flocculant dilution pump

2.3. Maintenance 2.3.1. Gearbox oil

2.3.1.1. Rake drive 2.3.1.2. Feedwell dilution liquor low-head pump (optional) 2.3.1.3. Interstage mixer

2.3.2. Hydraulic oil and filters with hydraulic rake drive (optional) 2.3.3. Descaling periodically 2.3.4. Lining or coating repair if damaged during abnormal operation or descaling 2.3.5. Pumps & pump shaft seals

Horizontal Belt Vacuum Filters

1. Capex 1.1. Civil Site Preparation to accommodate total filtration area

1.1.1. Filtrate splash or drip containment 1.2. Equipment

1.2.1. Leach discharge or filter feed thickener 1.2.2. Filter feed slurry storage tank with mixer 1.2.3. Horizontal belt vacuum filter

1.2.3.1. Number of filters to recover pay metals 1.2.3.2. Size of filters to produce minimum liquor content at plant production 1.2.3.3. Materials of construction for corrosion resistance during production and filter cloth

acid wash to remove scale 1.2.3.4. Cake thickness transmitter

1.2.4. Filter ancillaries 1.2.4.1. Filter feed slurry pumps variable speed 1.2.4.2. Flocculant supply system (optional)

1.2.4.2.1. Flocculent dry storage 1.2.4.2.2. Flocculant dissolution tank and mixer 1.2.4.2.3. Flocculant solution storage tank 1.2.4.2.4. Flocculant metering pump 1.2.4.2.5. Flocculant dose control

1.2.4.3. Feed slurry mass measurement 1.2.4.3.1. Feed slurry density transmitter one out of filter feed storage tank 1.2.4.3.2. Feed slurry flow transmitter one per filter

1.2.4.4. Vacuum system 1.2.4.4.1. Vacuum pumps operating and standby 1.2.4.4.2. Moisture traps

1.2.4.5. Filtrate system



1.2.4.5.1. Filtrate receives 1.2.4.5.2. Filtrate pumps

1.2.4.6. Cloth wash water system to remove suspended solids 1.2.4.6.1. Wash water storage tank 1.2.4.6.2. Wash water pump

1.2.4.7. Cloth acid wash system to remove scale (intermittent) 1.2.4.7.1. Acid wash storage tank 1.2.4.7.2. Acid wash pump

1.2.4.8. Interconnecting pipe 1.2.4.9. Stairs and platforms between and around filters 1.2.4.10. Filter cake conveyer

2. Opex 2.1. Power

2.1.1. Filter feed slurry tank mixer 2.1.2. Filter feed slurry pumps 2.1.3. Flocculant metering pumps 2.1.4. Rubber drainage belt drive motors 2.1.5. Vacuum pumps 2.1.6. Filtrate pumps 2.1.7. Cloth wash water pumps 2.1.8. Acid wash pump (intermittent) 2.1.9. Filter cake conveyors 2.1.10. All controls

2.2. Consumables 2.2.1. Flocculant (20-200 g/t range) 2.2.2. Filter cloths (2 per year per filter) 2.2.3. Vacuum pan wear belts ( 2 sets per filter per year) 2.2.4. Rubber drainage belts (1 per 10 years per filter) 2.2.5. Roller deck rollers and bearings

2.3. Process water 2.3.1. Vacuum pump seal water 2.3.2. Filtrate pump seal water 2.3.3. Cloth wash water 2.3.4. Acid wash 2.3.5. Vacuum pan wear belt lubricant water

2.4. Maintenance 2.4.1. Bearing lubrication 2.4.2. Cloth replacement 2.4.3. Vacuum pan wear belt replacement 2.4.4. Roller and bearing replacement 2.4.5. Gearbox oil

2.4.5.1. Filter feed slurry tank mixer 2.4.5.2. Rubber drainage belt drive

2.4.6. Rubber drainage belt replacement

The comparison of the total Capex and Opex for either CCD thickeners and Horizontal Belt Vacuum Filters can be made using solid/liquid separation rates and physical properties measured in lab scale simulations. The most cost effective pregnant liquor recovery option can then be selected.

CONCLUSIONS

Equipment development continues to keep pace with the demand to process larger capacities in the hydrometallurgical process market to recover copper, cobalt, nickel, gold, and uranium. Competing process developments require consideration of the feasibility to recovering pregnant liquor solution using either a CCD thickener circuit or a number of horizontal belt vacuum filters. One must consider the range of physical chemical properties of leach discharge slurry and select the



appropriate solid/liquid separation rates for thickeners and filter using the suitable materials of construction. Obtaining costs specific to the plant location for the Capex and Opex items identified in the paper will provide a meaningful economic comparison to be made.

ACKNOWLEDGEMENTS

The authors wish to acknowledge Jared Quilter for his expertise in preparing some of the graphics used in this paper and other peers within FLSmidth has contributed expertise in preparing this paper.

REFERENCES

AMIRA International, 2011. Improving Thickener Technology, Available from: http://www.amira.com.au/.

Klepper R.P., 2009, The Evolution of Thickeners, in Proceedings ALTA 2009 Nickel/Cobalt Conference, Perth, 25-27 May.

Klepper R.P., 2009, High Density Thickeners in CCD Circuits: Case Study, in proceedings of the International Symposium, Hydrometallurgy of Nickel and Cobalt 2009, 39th Annual Hydrometallurgical Meeting, Held in Conjunction with Nickel & Cobalt 2009, pp 195-207



Agglomeration – The Key to Success for the Murrin Murrin Ni Laterite Heap Leach

D Readett1 and J Fox2

1. FAusIMM(CP), Group Manager Project Development, Minara Resources Limited, Level 4, 30 The Esplanade, Perth WA 6000. Email: [email protected]

2. Heap Leach Superintendent, Minara Resources Limited, Level 4, 30 The Esplanade, Perth WA 6000. Email: [email protected]

ABSTRACT The most critical variable in the successful establishment of the Murrin Murrin Heap Leach plant was the effective agglomeration of the Ni laterite ore. This paper describes the test work undertaken to establish the key agglomeration parameters including state of the art testing and analysis procedures. These parameters were then utilised to establish the key design and operational concepts for the world’s first commercial scale Ni laterite heap leach. Following the technical and economic success of the operation, the actual operational data is compared back to the original test work data.

INTRODUCTION Agglomeration has been a key operating practice for both the gold and copper heap leach. McClelland et al., (1983) provided an historical perspective on precious metals heap leaching. The use of copper heap leaching has now become an industry standard. In 2005, Bouffard (2005) provided a review of agglomeration practice both for gold and copper heap leach operations. Bouffard (2005) concluded that ”agglomeration was a breakthrough technology...with ore of high fines or clay content. Achieving up to 80-90% metal recovery from ores at first thought to be heap-unleachable.” All of the operations that were reviewed used some form of chemical binder which was dependent on the leach environment that was subsequently going to be utilised. For copper heap leaching typical agglomeration was conducted in a drum agglomerator with the addition of 15-25kg/ sulphuric acid per tonne of ore and 60-100kg of water. Readett and Miller (1996 and 1997) and Readett (1997) highlighted that for copper bio-heap leach applications agglomeration using acidified leach liquor is required to ensure adequate inoculation of the ore with bacterial culture. As the leaching of Ni laterites is an acid based process (Agatzini-Leonardou and Dimak 1994), the data from gold leaching is of limited use. However the data generated from acid based copper heap leaching is of relevance. Ni laterite ores are also known to have unique characteristics and materials handling of these ores can be problematic. The agglomerated laterite ore can exacerbate these materials handling characteristics therefore these characteristics need to be considered in the detailed design of the proposed materials handling system.

LABORATORY TESTING Testing of Murrin Murrin ore types commenced in 2004. Initially “bottle roll” tests were performed on samples to determine their amenability to potential heap leach. A number of ore types achieved in excess of +50% Ni recoveries with the addition of ~300kg/t of acid.



Preliminary agglomeration testing indicated highly variable results. It was possible to add up to 200kg/t of acid solution in agglomeration, due to the very high inherent moisture content of the laterite ores, which ranged from 15 to 30%. However results from 1m column tests showed permeability issues at the very high acid solution addition rate to agglomeration (200kg/t). Acceptable column leaching was achieved with agglomerates from 50-150kg/t of acid solution. Unagglomerated ore was impermeable in column testing. Agglomeration and column testing with water and acidified leach liquor indicated that it was necessary to ensure optimal overall moisture addition for the best agglomeration result and that acidified leach liquor was superior to water as the moisture agent. In laboratory 1m columns the Ni rate and extent of extraction was found to be solely a function of acid solution addition, with more acid solution added in agglomeration resulting in earlier release of Ni. Each ore type however appeared to exhibit its own unique optimal agglomeration condition leading to optimal column performance. These preliminary results highlighted that it was necessary to try and establish laboratory testing techniques that allowed for the measurement of the effectiveness of agglomeration prior to conducting column leach testing. Additionally these tests were also utilised to try and establish if the effectiveness of agglomeration (including maximum potential heap height and maximum potential leach solution application rate) could be sustained over the leach cycle after the consumption of 300-500kg/t of acid.

A summary of the testing techniques undertaken is provided below

Kappes Percolation Test This is an empirical test attributed to Kappes, Cassidy and Associates and is a laboratory based single pass/fail test to establish if agglomerated gold ores would perform in an industrial heap leach application. The test was adopted for the Ni ores (Table 1). The test measures the drain down permeability of a saturate column of ore whereby empirical data indicates if the sample achieves a percolation rate in excess of 10,000L/hr/m2 and a slump of less than 15% then it will perform acceptably in an industrial heap leach application.

Table 1 – Kappes Percolation Results

Ore Type

Agglomeration Addition kg/t Bulk Density

t/m3

Drain Rate

L/m2/hr

Slump

% Water Acid Solution

Scats 34 0 1.14 4,278 1.1

0 25 1.00 41,559 1.1

13 50 1.11 43,545 1.0

0 100 1.17 23,529 1.5

0 150 1.28 1,757 1.1

BB 0 246 1.00 6,723 3.3

277 50 1.07 9,397 3.6

BD 0 421 1.11 7,181 0.0

388 50 1.15 16,119 0.5

B Blend 0 242 1.10 1,451 3.7

269 50 1.18 7,105 1.1



Load Permeability Test The load permeability test was used to establish the likely saturated permeability of a sample and can measure how the permeability may change with heap height (Fig. 2). It provides an indication of the relationship between heap height and the leach liquor application rate that would result in saturation of the ore. The results from this test are considered limited as they are influenced by the loading and saturation history of the sample tested (Williams, 2005). Results are likely therefore to be conservative.

Fig. 1 – Load Permeability Results

Soil Water Characteristic Curve (SWCC) The SWCC is used to describe the unsaturated hydraulic properties of soil and can be used to estimate the unsaturated permeability of a material with respect to the degree of saturation (Fig. 2).



Fig. 2 – SWCC Results for Scats and Scats Residue

Load Percolation Test This test is aimed at determining the maximum loading that can be applied to a sample, at a constant leach solution application rate, before flooding (saturation) occurs (Fig. 3). The maximum load can then be converted to an equivalent heap height for design purposes.



Fig. 3 – Typical Load Percolation Results

Saturated Hydraulic Conductivity Test The saturated hydraulic conductivity of a porous medium defines the maximum hydraulic capacity to transmit solution once all of the void space within a sample is occupied with water (i.e. saturated). To simulate a stacked heap the saturated hydraulic conductivity tests were conducted under various loading conditions (Fig. 4). These tests were designed to establish the short term upper percolation capacity (Guzman, 2005). The test however does not quantify the transient behavior as a result of the interaction of acid solution with the ore.

Fig. 4 – Scats Saturated Hydraulic Conductivity

Stacking (or Stacked Density) Test This test protocol was recommended by Guzman (2005) based on experience that had shown that the hydraulic performance of a heap is strongly controlled by the density of the ore. Samples are subjected to various static loads to define the relationship between heap height and dry bulk density, and hence provide the density profile throughout the proposed stacking depth. These tests are conducted under partially saturated conditions on the basis that this better represents the actual



density profile developed during ore stacking and also allows determination of the hydraulic conductivity without the destruction of the agglomerated product (Fig.5).

Fig. 5 – Stacked Density Test Results for Ore

Summary of Tests No single test was found to provide a definitive answer to the ultimate question regarding true performance of an agglomerated material in a commercial operation. Each was found to provide indicative data on the optimal agglomeration conditions. In general the test work indicated that the use of acid solution, such as CCD leach solution, as the moisture agent in agglomeration was more effective that water. A total acid solution addition of 50-100kg/t was found to be optimal. The optimal condition for each ore type was unique. However, even when combined with results from an extensive column leach program a definitive answer on the optimal heap height and solution application rate required for the successful heap leaching of Ni laterite ores through to completion could not be established (Scheffel 2005). As an example, for the acid solution agglomerated scats

Kappes tests indicated the material was heap leachable. Load permeability showed ranges of maximum heap height of:

a) 4m at 5 to 45L/hr/m2, or b) 8m at 3 to 25L/hr/m2.

Load Percolation tests showed ranges of maximum heap height of: a) 19m at 30L/hr/m2, or b) 7m at 60L/hr/m2.

Saturated Hydraulic Conductivity was 0.14 to 0.24 cm/s. Stacked Density was:

a) 0.9 g/cm3 at 4m, and b) 0.94 g/cm3 at 8m.

Column tests indicated ranges of maximum heap height of: a) 4 at 60L/hr/m2, to b) 8m at 0 to 30L/hr/m2.

On the other hand, results for leached scats were:. Kappes tests indicated the material was heap leachable. Load permeability showed ranges of maximum heap height of:

a) 4m at 2 - 14L/hr/m2, or b) 8m at 1 - 7L/hr/m2.

Load Percolation tests showed ranges of maximum heap height of: a) 20m at 30L/hr/m2, or b) 8 - 20m at 60L/hr/m2.



Saturated Hydraulic Conductivity was 0.04 to 0.005cm/s. Stacked Density was:

a) 0.6 g/cm3 at 4m, and b) 0.68 g/cm3 at 8m.

Column tests indicated ranges of maximum heap height of: a) 4 at 60L/hr/m2, or b) 8m at 0 to 15L/hr/m2.

An added complication with the results was that in general if the agglomerated material was exposed to excessive mechanical impact or saturated leach conditions its resultant permeability approached zero (ie un-heap leachable). As expected the testing programme highlighted the unique nature of the laterite ores and also the unique nature of the acid heap leach process. Laterite ores are naturally high in clay content, high in moisture content and have very low density. The stacked ore can have a dry bulk density of less than 1.0 g/cm3 so the impact of solution application (with a leach solution with an SG of 1.1 to 1.3 g/cm3) and solution saturation can be extreme. The ore exhibits extremely high acid demand (350-500kg/t) to maximize Ni recovery. As a consequence, heap leaching demands high solution application rates over extended leaching cycle times of 1 to 2 years. The stacked ore exhibits a high degree of initial slump upon commencement of leaching and then during the leach cycle the acid can consume 15-30% of the mass of the ore. Ultimately the lack of a definitive answer led the project team to recommend the construction of demonstration industrial/commercial scale heaps. Then upon successful operation of the demonstration heaps the demonstration plant could be expanded to a full commercial plant.

AGGLOMERATION AND MATERIALS HANDLING DESIGN The test work indicated that well controlled effective agglomeration was the key to success. Additionally it indicated that the ore and subsequent agglomerates exhibited very poor materials handling characteristics and rapid deterioration of permeability under saturated conditions, both of which could adversely affect the heap leach performance of the agglomerated ores. Additionally the agglomerates were relatively low in strength and agglomeration effectiveness could be destroyed with excessive handling. This combination of characteristics required a very high degree of attention to detail in the design of the agglomeration and materials handling system. The key design areas were bins, agglomeration drum and chutes/transfer points. The ROM feed bin and subsequent surge bin were designed with very steep angles and were to be lined with a teflon based wear system. This would ensure minimal likelihood of material compaction and hang up within the bins. Some of the key design features of the agglomeration drum included

a steeply angled and teflon lined feed chute, an effective control system for moisture and acid addition, the design and placement of moisture and acid addition points within the drum to ensure

maximum mixing and minimal potential impact with the drum lining system, and the addition of a scraper bar to allow for removal of compacted ore from within the drum.

Post agglomeration, the design incorporated the minimum number of transfer points to minimize the potential for mechanical degradation of the agglomerates. Each transfer point was also designed to allow the agglomerated ore to free fall onto the proceeding conveyor. The permeability and subsequent column testing indicated that, depending on ore types and agglomeration conditions, the optimum heap height should be in the range of 4 to 8m. As a consequence the stacker design allowed for a variable stack height of ore from about 3m up to 9m. Heap drainage was also a key design parameter. The testing had highlighted that saturation of the heap could result in the destruction of the agglomerates and effectively reduce permeability to a minimum. As such it was necessary to ensure heaps were adequately drained. The need for drainage



was exacerbated by the high application rates required. To ensure effective drainage was achieved a 500mm layer of free draining acid resistant coarse rock was placed at the base of each heap. Within the drainage material a “half herring bone” arrangement of drainage piping was also installed. The drainage piping connected to header drainage pipes that ran the full length of the pad and exited into the collection sump at the front of each pad.

OPERATING EXPERIENCE The attention to detail in the design phase ensured minimal commissioning issues and minimal longer term operational problems. To date there have been no significant changes to the overall design of the agglomeration and materials handling system (Readett et al 2006). The agglomerator feed chute did cause some preliminary compaction and blockage issues however the addition of a small water spray directly onto the chute impact surface overcame this. The ore type that has been predominantly stacked is “scats”. Testing of this material indicated stack heights of 2 - 10m could support application rates of 15-30 L/hr/m2. Column leach testing indicated potential for heap saturation at 8m and 15-30 L/hr/m2. The actual operational data has tended to the conservative end of the test results. Typically optimal heap heights are 3-5m with sustainable initial application rates of 15-20 L/hr/m2 and longer term rates degrading down to 5-10 L/hr/m2. Solution drainage at the base of the pads has worked effectively and no drainage related side slope failures have occurred. Leaching of fresh laterite ore has also been implemented using the optimal agglomeration conditions established in the test work program. Again the actual operational data has tended to the conservative conditions determined in testing. From a metallurgical perspective the heaps have performed satisfactorily and results of the heaps performance has been provided in a number of papers (Readett and Fox, 2009 and 2010).

CONCLUSIONS The most critical variable in the successful establishment of the Murrin Murrin Heap Leach plant was the effective agglomeration of the Ni laterite ore. The test work conducted allowed the optimal agglomeration conditions to be established and in general provided an optimistic indication of the likely performance in the full scale operation. Results also provided a focus for key design criteria and necessary materials handling characteristics that required detailed attention. In general the test results overstated the ultimate optimal heap height and application rate that could be sustained in a full scale operation over the entire leach cycle. Use of all of the available data ultimately helped to ensure a successful project. In hindsight a conservative approach was warranted considering the unique nature of this project.

ACKNOWLEDGEMENTS The authors would like to thank Minara Resources Limited for permission to publish and acknowledge

the contribution of Michael Rodriguez, David Williams, Randy Scheffel and Amado Guzman.

REFERENCES Agatzini-Leonardou, S. and Dimak, D, 1994 Heap Leaching of Poor Nickel Laterites by Sulphuric Acid at Ambient

Temperature, International Symposium Hydrometallurgy ’94, pp193-208

Bouffard, S C, 2005 Review of Agglomeration Practice Fundamentals in Heap Leaching, Mineral Processing &

Extractive Metall. Rev.26: pp 233-294, 2005.



Guzman, A, 2005. Personal communication.

McClelland, G E, Pool, D L and Eisele, J A, 1983. Agglomeration-Heap Leaching Operations in the Precious

Metals Industry, USBM IC1945, 1983.

Readett, D J and Miller, G M, 1996. Practical Aspects of Heap Leaching, Copper Hydrometallurgy Forum, ALTA,

Brisbane, Australia, 1996.

Readett, D J and Miller, G M, 1997. Engineering and Process Developments associated with Industrial Scale

Copper Bioleaching, IBS Biomine 97 Conference Sydney, Australia.

Readett, D J, 1997. Engineering and Operational Aspects of Copper Heap Bioleaching, AMF Application of

Biotechnology to Economic Recovery of Metals from Ores and Concentrates, Sydney, Australia, 1997.

Readett, D J, Meadows, N E and Rodriguez, M, 2006 Murrin Murrin Heap Leach Project, AusIMM MetPlant

Conference Perth, September 2006

Readett, D. J and Fox J, 2009. Development of Heap Leaching at its Integration into the Murrin Murrin

Operations ALTA Nickel Cobalt Conference, Perth, Australia, May 2009

Readett, D J and Fox J, 2010. Commercialisation of Ni Heap Leaching at Murrin Murrin Operations, IMPC 2010,

Brisbane, Australia, September 2010

Scheffel, R E, 2005. Personal communication.

Williams D, 2005. Personal communication.



Scale Suppression using Swirl Flow Technology

J Wu1, D Nairn2, B Nguyen3, J Farrow4 and D Stegink5

1. CSIRO, Advanced Processing Technologies, MDU, Process Science and Engineering Division, Graham Road, Highett Vic 3190. Email: [email protected]

2. CSIRO, Advanced Processing Technologies, MDU, Process Science and Engineering Division, Graham Road, Highett Vic 3190. Email: David. [email protected]

3. CSIRO, Advanced Processing Technologies, MDU, Process Science and Engineering Division, Graham Road, Highett Vic 3190.

4. FAusIMM, CSIRO, Advanced Processing Technologies, MDU, Process Science and Engineering Division, Graham Road, Highett Vic 3190.

5. Queensland Alumina Ltd, Parsons Point, Gladstone Qld 4680.

ABSTRACT Swirl Flow Technology (SFT), a novel approach for agitation in large industrial slurry tanks, was jointly

developed and patented by CSIRO and Queensland Alumina Ltd (QAL). Currently, SFT is used in 16

alumina precipitation tanks at QAL, with the first installation in January 1997. This paper provides insight to

the underlying fluid dynamics, which were investigated through pilot-scale experiments. A key finding was

that SFT produces tank wall velocities approximately four times that of conventional draft tube agitator

systems for the same power input, whilst for solids suspension, much less power is needed. SFT is energy

efficient and practical in scaling suppression as demonstrated by the full-scale experience at QAL.

INTRODUCTION Many mineral processing operations such leaching, digestion, and precipitation require the effective agitation

of large volumes of slurry. The tanks used for these operations can be as large as 30 m high and 15 m in

diameter. Often a large number of these slurry tanks are installed for continuous chemical reactions. This

makes it possible to deliver multi-million tons per year of refined metal or concentrated ore products, even

though the reactions may be slow, and require ore particles to be suspended in chemical solutions for many

hours or even days.

In low viscosity Newtonian slurry mixing tank operation, agitators are often designed on the basis of

achieving off-bottom solids suspension as other mixing processes satisfactorily follow once the solids are

suspended. For example, in fully suspended Newtonian slurry the time required to mix liquid and solids or

liquids is typically minutes, which is an order of magnitude shorter than the hours to days required for

reactions such as leaching or crystallization.

Typically, axial flow impellers pumping downward with vertical baffles are used for solids suspension

operations. It has been generally accepted that axial flow impellers are more energy efficient than radial

turbines in fully baffled tanks (Nienow 1992; Ibrahim and Nienow 1996; Wu et al 2001), and that the energy

efficiency for off-bottom solids suspension is sensitive to the impeller off-bottom clearance and impeller

diameter, Nienow (1992); Chapman et al (1983); Wu et al (2001, 2002). Zwietering (1958), Nienow (1992)

and Wu et al (2001, 2006, 2007 and 2010) have described the basis of solids suspension in mixing tanks.



Such conventional agitation designs are mostly developed from chemical industry experience, where

reactions are fast and the tanks used are relatively small compared to the tanks used in the minerals

industry, where the reactions are often slow.

Scale growth is a major cause for loss of production through tank downtime in minerals processing.

Conventional agitators are not designed to address growth of scale, which is probably a more serious

problem in the minerals industry than in most other process industries.

In 1997, CSIRO and Queensland Alumina Limited (QAL) developed Swirl Flow Technology (SFT) (Welsh

2002) to address this problem. In contrast to conventional agitation approaches, SFT employs a radial flow

impeller located at the top of the tank to draw the slurry up the centre of the tank, and discharge the slurry

from the impeller with a swirling motion (Fig. 1). Radial flow impellers have Power numbers (a measure of

volumetric pumping efficiency) close to one, while axial flow impellers typically have Power numbers of about

0.3. This implies that for the same mixing flow, SFT impellers have greater flexibility in the tip speed vs

capacity trade-off compared to axial flow impellers.

The highly structured swirling flow patterns can suspend solids at lower power inputs than conventional

agitators, or can deliver tank wall velocities an order of magnitude higher than the wall velocities in

conventionally agitated tanks. These high velocities are effective in reducing the growth rate of scale on the

walls.

Fig. 1 Swirl Flow Technology, showing the intense inner vortex and high wall velocities.

This paper outlines the full-scale operational experience and benefits that have been achieved at QAL using

SFT. In addition the paper also provides experimental data from an application outside of the alumina



industry that demonstrates the features of SFT and its superior performance compared to conventional

agitation technologies.

INDUSTRIAL EXPERIENCE- ALUMINA Since the first installation in 1997 QAL has converted 16 precipitation tanks (24 m high x 11 m diameter, ~3

ML) to Swirl Flow Technology. QAL experience of the benefits of SFT includes:

Approximate doubling of service life between de-scaling operations compared to the draft tube

agitators;

Superior re-suspension capabilities after power failures, namely a few hours for SFT vs. several

days, if at all, for a conventional draft-tube tank;

Lower capital cost due to simplified support structures;

Health and Safety benefits due to ability to perform agitator maintenance without entering the tank;

Reduced repair and maintenance cost due to the simpler configuration;

Improved service life for the impeller compared to agitators in draft tube systems.

Fig. 2 Industrial application of SFT in an alumina precipitator vessel at QAL.

OTHER APPLICATIONS When considering a new application for SFT, laboratory pilot-scale testing is required to size the agitator

system and to give confidence that the desired performance will be achieved. The experiments and results

obtained in various pilot-scale tests for one given application from the gold industry are described below.

EXPERIMENTAL SET-UP AT CSIRO Two mixing rigs (Fig. 3(a) and 3 (b)) one of 390 mm diameter and the other 1000 mm diameter were used.

To minimize optical distortion both were placed inside water-filled, transparent rectangular tanks. Test

impellers were mounted on the central shafts, which were equipped with Ono Sokki torque transducers and

speed detectors.



(a) (b) Fig. 3 Test Facility at CSIRO, (a) schematic of rig, and (b) mixing tank.

Fig 4 & 5 shows laboratory test configurations based on a study for an application of SFT within the gold

industry. The study compared the conventional configuration proposed by a mixer vendor with the CSIRO

SFT.

The conventional configuration would have required that the existing cone bottom tanks be modified to have

a flat floor, whereas SFT works well in cone bottom vessels.

Because this will be the first SFT application outside alumina, a comprehensive laboratory study was

undertaken to compare the conventional design with SFT, to ensure that the technical risks to the plant were

minimized. The test solids were glass particles of SG = 2.5 at a concentration of 600 g L-1, and the test

liquor was water.



Fig. 4 Pilot-scale Conventional Multiple Impeller Configuration

Fig. 5 Pilot-scale CSIRO SFT Configuration

Velocity Velocity in the liquid was measured using a Laser Doppler Velocimetry system (Wu et al, 2006).

Speed, Torque and Power The speed and torque were logged using a personal computer equipped with a data acquisition board, thus

providing on-line analysis of power draw.

Mixing Rate The mixing rate was quantified by measuring the time required to homogenize the liquid phase. A salt tracer

solution was injected into the tank through a dip tube submerged below the liquid surface, and the salt

concentration recorded using conductivity probes placed at the mid height of the tank (Wu et al 2011). The

time taken for the salt concentration to reach steady state was defined as the mixing time(tM).

Solids Suspension The liquid and solids flows in the tank bottom were observed through the transparent tank walls and the tank

floor. The sedimentation bed height (HB) was recorded as the impeller speed was reduced from a full solids

suspension condition. Refer to Wu et al (2010, 2011) for more detailed description on the experimental

methods.



Solids Dispersion The solids dispersion through the tank volume was measured by observing the interface height (HS) between

the slurry and the clear supernatant liquid.

RESULTS

Solids Suspension Fig. 6 shows the normalized sedimentation bed height (HB/H, where HB is the 100% settled bed height) as a

function of the shaft speed N. As is expected, when the shaft speed increases from zero, sedimentation bed

height decreases.

The impeller speed required to just achieve off-bottom solids suspension (Njs) is defined as the speed where

HB = 0. In this example Njs is 125 rpm for the SFT design, and 150 rpm for the conventional multiple impeller

design.

0.00

0.10

0.20

0.30

0.40

0.50

0 50 100 150 200 250 300

Agitator Speed N (rpm)

No

rmal

ized

Sed

imen

tati

on

Bed

Hei

gh

t H

B/H

Conventional Mixer, 3 X baffles_ 3 X A310

CSIRO SFT, S36B, 60 Deg cone bottom

Tank dia = 1.0mSolid concentration = 600 g/L

Fig. 6 SFT suspends solids at lower agitator speed

An important design parameter is the power needed to obtain off-bottom solids suspension. In Fig. 7 the

normalized sedimentation bed height is plotted against the shaft power for both designs. To suspend solids

off the tank bottom (i.e. HB=0), the conventional mixer design requires 125 W whilst the SFT design requires

25 W, one fifth the power.



Fig. 7 SFT suspends solids at lower agitator power

Solids Dispersion A full dispersion of solids corresponds to good utilisation of the tank volume for processing operations.

Dispersion of solids is conveniently quantified by the solids suspension height, Hs, measured in the

laboratory by recording the visual interface between the dense slurry suspension and the clarified liquid layer

at the tank top region. Fig. 8 shows that the normalized slurry height produced by the SFT was higher than

that achieved by the conventional mixer over the test power range, suggesting a superior solids dispersion

performance with the SFT design.

Fig.8 SFT delivers superior solids dispersion

Wall Velocities Fig. 9. Shows the velocity distributions measured using Laser Doppler Velocimetry for a SFT impeller and for

a typical draft tube agitator configuration. The velocities shown are normalized by the impeller tip velocity

and the radial position is normalized by the tank radius.

0.00

0.20

0.40

0.60

0.80

1.00

0 40 80 120 160 200 240 280 320 360 400

Agitator power (W)

No

rmal

ize

d S

lurr

y H

eig

ht

Hs/

H

Conventional Mixer, flat bottom, 3 X baffles_ 3 X A310

CSIRO SFT-S36B, 60 Deg Cone bottom


0.00

0.10

0.20

0.30

0.40

0.50

0 40 80 120 160 200 240 280 320 360 400

Agitator power (W)

No

rma

lized

Sed

ime

nta

tio

n B

ed

He

igh

t H

B/H

Conventional Mixer, flat bottom 3 X baffles_ 3 X A310

CSIRO SFT, ST36B, 60 Deg cone bottom




-1

-0.5

0

0.5

1

0 0.2 0.4 0.6 0.8 1

(normalized radial position) r/RNo

rmal

ized

Vel

oci

ty

SFT Dia 285mm, LDV at z/H = 0.6

Draft tube C110 agitator dia 285mm, LDV at z/H=0.75

draft tube

SFT rotor tip

Fig. 9 Radial Variation of Velocity for SFT and for a Draft tube agitator. The velocities are the tangential component for the SFT rotor, and the axial component for the draft tube design (positive

for upward direction). Other velocity components are negligible.

The higher velocities towards the centre correspond to the strong spinning at the vortex core. In the case of

the draft tube agitator design, the velocity changes direction, being downward inside the draft tube (negative

sign), and upward in the annulus.

For suppression of scale growth, the velocity near the tank wall surface immediately adjacent to the

boundary layer is a critical factor, based on operational experience at QAL. From Fig. 9:

,16.0tipU

V for SFT ,03.0

tipU

V for draft tube agitator,

where V (m s-1) is the velocity near the wall (r/R~1) and Utip (m s-1)is the impeller tip velocity.

Swirl Flow provides approximately five times the velocity at the wall compared to a draft tube agitator design,

at the same impeller tip velocity.

To compare the velocity on a power input basis a velocity circulation efficiency parameter as described in Wu

et al (2006) can be used:

,17.1)/( 3/12

TP

V

for SFT, and 27.0 for draft tube agitators

Where is the non-dimensional velocity efficiency parameter, P is the agitator power input (W), is the

liquid density (kg m-3), and T is the tank diameter (m).

SFT can provide ~4.3 times the velocity over the tank wall surface for the same power input compared to the

draft tube agitator design. Since power varies as V3, a conventional draft tube agitator would need to operate

at ~4.33 = 80 times the SFT power, to match the wall velocity performance of the SFT rotor.



Mixing Times For some applications, the liquid mixing time is important. Figure 10 shows mixing times for generic axial

flow impellers, one in a baffled tank, the other unbaffled, obtained from a laboratory study.

0

40

80

120

160

200

0 50 100 150 200 250 300 Time (Sec)

Sa

lt C

on

cen

trat

ion

(p

pm

)

No Baffles

4x Baffles

t=t m , (mixing time)

Fig. 10 Mixing time measurements: the effect of baffles, reproduced from Figure 11 in Wu et al 2011.

This is not specifically a comparison between SFT and conventional impellers, but it is general enough to

draw conclusions about the order of magnitude of mixing times. One might expect that the mixing rate would

be lower due to the whole body rotation of the slurry upon removal of the baffles. Apparently there is still

significant internal motion within that rotation. The mixing time without baffles is increased by 50%

compared to that with baffles. What is significant is that the mixing time is seconds (or minutes in full-scale

tanks) rather than hours, and the difference between the two configurations is not likely to be relevant when

residence times are several hours.

ADVANTAGES OF SFT

Mechanical Features SFT is installed in the top third of the vessel. The short cantilevered shaft design means less complexity with

no bottom bearing, no associated mechanical seal and support structure.

SFT requires that the vessel be free of baffles and other intrusions in the tank. Not needing baffles allows for

potentially lower capital cost and reduced areas for scale to grow.

The total weight of an SFT installation can be significantly less than that of a conventional installation for the

same process duty, possibly allowing for lighter support structures.

Access is from above, eliminating the need for tank entry for agitator maintenance.

Maintenance time is also reduced due to simpler access, lower weight and less scale and erosion.



High Wall Velocities Scaling at a material surface can often be reduced with increased shearing over the surface, although the

relationship between shear stress and the scale growth rate is not well understood. Industrial experience in

the QAL precipitator tanks suggests that scale growth rates are halved with SFT compared to conventional

mixers. Laboratory tests have shown that SFT produces wall velocities approximately five times higher than

does the conventional draft tube agitator system, for the same power input.

Re-start As the SFT impeller is located in the top third of the tank, there is no “bogging” issue as can occur with

conventional agitators, particularly after a power failure. Full-scale experience at QAL suggests that SFT

impellers restart easily and re-suspend all the solids in much shorter time than a draft tube agitator system.

Process Intensification The ease of re-starting SFT after power failure provides scope for operating at very high solids loading which

may enable throughput increases. Wu et al 2007 and 2011 showed that the production throughput in a

slurry reactor could be increased by operating at high solids concentration or by operating the slurry

suspension with stratification, such that the solids residence time is made larger than the bulk flow residence

time.

Tip Speed & Erosion The SFT impeller operates at a lower tip velocity than an axial flow impeller for the same power input. This

has several advantages including lower potential for impeller erosion and potentially reduced attrition to

particles, which might be important in processes such as carbon in pulp leaching of gold or when leaching

friable materials.

Full-scale experience at QAL suggests that SFT impellers operating in alumina precipitation tanks do not

usually have erosion problems; by contrast industry standard agitators (e.g. draft tube agitators) in such

systems often experience serious erosion damage.

LIMITATIONS OF SFT

Mixing Time A “side effect” with using SFT is increased mixing time. This is usually not a problem in the mineral

processing industry, where the time scale for reactions, and hence the slurry residence time are typically

many hours compared to minutes for liquid mixing.

Non-Newtonian systems SFT is usually not suitable for high viscosity slurries, e.g. non-Newtonian with shear-thinning rheology

properties. More research is required to quantify this limitation, in terms of the viscosity or Reynolds number.



Gas Dispersion SFT can disperse moderate flows of gas as used for oxygen injection in CIL. SFT however is not suitable for

dispersing large flows of gas such as used in base metal leaching.

CONCLUSIONS SFT is a novel approach for solids suspension and mixing applications commonly encountered in

hydrometallurgical processing. Laser Doppler Velocimetry measurements, pilot scale measurements of

solids suspension and tracer measurements confirmed the performance of SFT compared to conventional

agitator systems.

The full-scale experience at QAL in 3 ML precipitation tanks demonstrates the practical advantages of SFT

compared to draft tube agitators. These are:

better off-bottom solids suspension energy efficiency,

lower rates of scale growth,

easier re-starting after power failure,

lower maintenance costs, and

lower capital cost.

ACKNOWLEDGEMENTS The authors wish to acknowledge the valuable long-term technical collaboration with QAL in developing Swirl

Flow Technology. The kind permission of Gold Fields Australia to quote the data from the testwork for their

application is acknowledged.

REFERENCES Chapman CM, Nienow AW, Cook M, Middleton JC. Particle-gas-liquid mixing in stirred vessels, Part I:

Particle-Liquid Mixing. Chem Eng. Res Des. 1983; 61: 71-81.

Ibrahim S, Nienow AW. Particle Suspension in Turbulent Regime: the Effect of Impeller Type and

Impeller/vessel Configuration. Trans IChemE. 1996; 74: 679-688.

Nienow AW. The Suspension of Solid Particles. In: Hamby N, Edward MF and Nienow AW. Mixing in

the Process Industries. London: Butterworths 1992.

Welsh Martin C, “Method and Apparatus for Mixing”, US Patent 6,467,947 B1, Oct 22, 2002.

Wu J, Zhu Y, Pullum L. Impeller geometry effect on velocity and solids suspension. Trans IChemE.

2001; 79(A): 989-997.

Wu J, Zhu Y, Pullum L. Suspension of high concentration slurry. J. AICHE. 2002; 48(6): 1349-1352.

Wu J, Graham LJ, Nguyen B, Mehidi MNN. Energy Efficiency Study on Axial Flow Impellers. Chemical

Engineering and Processing. 2006; 40: 625-632.

Wu J, Graham LJ, Mehidi MNN. Intensification of Mixing. Journal of Chemical Engineering of Japan.

2007; 40(11): 890-895.



Wu J, Nguyen B. and Graham LJ. Energy Efficient High Solids Loading Agitation for the mineral

Industry. Can. J. Chem. Eng. 2010; 88: 287-294.

Wu J, S. Wang, L. Graham, R.Parthasarathy and B. Nguyen, High Solids Concentration Agitation for

Minerals Process Intensification, AICHE Journal, in press 2011.

Zwietering TN. Suspension of Solids in Liquid by Agitators. Chem. Eng. Sci. 1958; 8: 244–253.


On-Belt Analysis at Sepon Copper Operation

T Arena1 and J McTiernan2

1. Process Engineer, MMG LXML Sepon, Savanakhet Province, Lao PDR. Email: [email protected]

2. Superintendent, MMG LXML Sepon, Savanakhet Province, Lao PDR. Email: [email protected]

ABSTRACT To maintain copper recoveries and maximise production at the Sepon Copper Operation it is crucial to control the amount of copper entering the leaching circuit. This requires close control of blending of the multiple ROM stockpiles to achieve the desired head grade. Grade variability in the ROM stockpiles can lead to significant variation in the processing plant copper feed grade.

Head grade sampling from the milling circuit at Sepon is complicated due to milling in leach solution to preserve the water balance. Copper is present in both liquid and solid forms and copper in the added milling solution has to be accounted for. This leads to difficulties in manual head sampling which also introduces delays due to sample preparation and assay turnaround. On-stream analysis is complex because of the solution and solid copper and the need to account for recycled copper solution.

This paper describes the complexity of feed grade management at Sepon and the application of on-belt analyser technology to assist in the control of copper metal throughput.

INTRODUCTION The Sepon Copper project is located approximately 40 km north of the township of Sepon in the Savanakhet province in southern Laos and is operated by MMG LXML (“MMG”) in which the Lao government holds a 10% interest with MMG owning the remaining 90%. The copper plant was commissioned in 2005 and designed to produce 60,000 t/year of LME A grade copper cathode. In 2010 an expansion was completed to increase the design copper production to 80,000 t/year.

Run-of-mine ore is crushed, milled and acid leached at 80°C for seven hours to extract copper into solution. The residue is thickened, and the copper rich clarified pregnant leach solution (PLS) is treated by solvent extraction and electrowinning to produce copper cathode. Thickened leach residue is washed with low grade copper raffinate solution and water in a counter-current decantation (CCD) circuit and the residue sent to flotation for pyrite recovery. The flotation residue is neutralised and sent to a tailings storage facility. The flotation concentrate is oxidised in an autoclave at 220 °C to generate heat, soluble iron and acid required for copper leaching (Keokhounsy S et al., 2006). Fig. 1 Sepon Process Overview


Fig. 1 Sepon Process Overview Feed grade control is critical to ensure that the atmospheric leaching section operates at maximum capacity. The maximum capacity is set by the regeneration rate of ferric ion (the active lixiviant for the chalcocite) and by the tankhouse. Treating over this maximum capacity will result in lower recoveries. A nominal leach feed grade of 4.6% copper is treated at 230 dry tph to the leach plant resulting in a copper metal feed rate of 10.6 tph. At design circuit conditions 9.8 tph of copper is leached into solution representing a recovery of 92%.

Post mining feed grade control starts at the copper ROM pad where trucks collect ore from multiple stockpiles varying in both grade and mineralogy. The majority of the copper ore extracted from the mine is secondary copper, predominantly chalcocite (Cu2S), with the remainder carbonate mineralisation containing significant quantities of azurite (Cu3(CO3)2(OH)2) and malachite (Cu2CO3(OH)2). The chalcocite deposit varies throughout the orebody from 1% through to 10% copper. Large malachite and azurite rocks in the carbonate ore also result in a highly non uniform grade distribution throughout these stockpiles Fig. 2 Azurite & Malachite Samples in the Sepon Carbonate Stockpiles. The blend ratio from each stockpile is determined by the ROM management team in conjunction with the copper processing engineers based on geological block model data of the stockpiles and the requirement of the processing plant.

Fig. 2 Azurite & Malachite Samples in the Sepon Carbonate Stockpiles The reclaimed ore is fed to a ROM bin and is crushed through a sizer to a nominal size of 100 mm. The crushed ROM ore is then slurried with acidic CCD overflow liquor and milled to 80% passing


106 microns. Cyclone overflow, at 37.5% w/w solids, flows by gravity to a linear screen that provides protection for the downstream leach process from oversize material. The screened slurry then reports to three leach surge tanks with a combined capacity of 4,800m3. The surge tanks provide approximately eight hours surge capacity to enable routine maintenance to be completed in the milling area. The estimated copper grade, slurry density and slurry flowrate are then used to provide the optimum copper metal throughput of the atmospheric leaching plant. Fig. 3 Sepon Crushing & Milling Circuit

Fig. 3 Sepon Crushing & Milling Circuit Design feed grade control was based on a calculated leach feed assay utilising data from multiple streams in the milling circuit. On-stream analysis was not included as part of the original plant design due to the complexity involving copper in both solution and solid streams as well as the need to account for recycled copper solution (CCD overflow). Spot sampling of the mill feed conveyor was not a viable option owing to the coarse mill feed size and the high grade variability. The sample size required to obtain accurate representation of the feed would be too large to collect routinely.

Manual sampling incurs significant time delays to obtain assay results and therefore deviations in the expected mill feed grade could continue for many hours. The variability of the feed and the delay in receiving data resulted in significant challenges in maintaining optimum metal units to the leach plant.

FEED GRADE CONTROL . The use of acidic CCD overflow liquor as grinding solution is driven by the need to limit the amount of water introduced into the process with the feed solids. This solution contains approximately 20 g/l sulphuric acid and 10 - 15 g/l copper. The carbonate minerals in the feed are readily soluble in acid and typically begin the leaching process in the mill prior to discharge into the leach surge tanks. Chalcocite also partially leaches in this solution due to residual ferric in the grinding solution. The slurry stored in the leach surge tanks therefore contains copper in the solid phase from fresh ore, copper in the solution phase that has been leached from the ore and copper in the solution phase that was present in the CCD overflow liquor.

The leach feed copper is calculated on a four hourly basis using a number of measurements from the milling circuit and leach feed tanks. Spot samples are taken from the leach feed tanks and CCD overflow solution. The leach feed slurry is then filtered. The solution phases and washed solids are


analysed for copper and the solution densities and % solids determined for the relevant streams. Leach feed copper flow-rates and mill feed grades are then calculated using mill tonnages for the relevant period after deduction of contained copper in the CCD overflow solution.

Results can take up to two hours for solution assays and seven hours for solid assays owing to the time necessary for sample collection and preparation. For the majority of the process streams this is not a problem owing to the large circuit inventory. However, the grade and composition of the ROM stockpile material can vary significantly in this time resulting in large changes in the composition of the leach surge tank material.

At the time of commissioning this assay data was the only information available on which to base feed blend or leach throughput changes. The infrequency of sampling and the delay in obtaining assay data could result in multiple shifts before a change could confidently be made to either the feed blend or the leach feed tonnage. On-line analysis was identified post commissioning as a technology that could provide valuable process control data if a suitable technology could be identified.

ALTERNATIVE METHODS FOR HEAD GRADE DETERMINATION Slurry analysers are traditionally used in mineral processing plants throughout the world. Initial consideration was given to the installation of a unit to analyse the slurry between the cyclone overflow and the leach surge tanks. However, it was uncertain how the analyser would respond to the copper being in both the solid and solution phases, as this application had not previously been tested.

Discussions were also held with suppliers regarding a custom technology solution. It was proposed that the solid and solution phases could be separated, the solids then analysed separately to the solution, with a separate stream analysing CCD overflow liquor to enable the copper feed grade calculation. There was considerable risk associated with an untested custom application involving a high degree of automation, and this option was ultimately cost prohibitive.

The third option considered was on-belt analysis Fig. 4 On-belt Analyser Schematic. This technology had not previously been implemented for base metals but its use is common in the iron ore, coal and cement industry. The technology enables analysis to take place whilst the ore is still in solid form, prior to contact with acidic copper containing CCD overflow liquor.

On-belt analysis uses a phenomenon known as Prompt Gamma Neutron Analysis (PGNAA). The ore sample is exposed on a conveyor belt to a beam of neutrons emitted from a Californium 252 source. When a neutron contacts an element it results in a gamma ray being emitted with an energy level specific to that element. A detector collects the intensity of the energy and software is used to correlate this with lab data to provide meaningful elemental assays every two minutes. The belt-analyser is capable of analysing a large suite of elements although the major element of interest in this project was copper.


Fig. 4 On-belt Analyser Schematic

To evaluate the feasibility of this technology nine 200 L samples of Sepon ore from various stockpiles and belt cuts were analysed using a factory on-belt analyser. Sub samples of the ore were then analysed in the assay laboratory and both sets of data used to develop calibration equations Fig. 5 Analyser Calibration Data. The feasibility test-work indicated that the calibrated instrument was capable of predicting the copper grade with a standard deviation of 0.5%.

Fig. 5 Analyser Calibration Data Although not accurate enough to eliminate the need for routine samples the data would enable early identification of unscheduled variations in the mill feed and assess the effectiveness of blend changes almost instantly. The on-belt analyser project was approved in August 2007 and commissioned in May 2008.


INSTALLATION AND OPERATION Physical installation of the analyser was challenging owing to the lack of available space at low elevation on the copper mill feed conveyor, the weight of the analyser unit, the requirement for access to the equipment and safety considerations associated with the neutron source.

The analyser envelope around the conveyor belt and it was necessary to remove part of the conveyor structure to accommodate the unit. This required new supports to be installed so that the conveyor frame could be split into two sections and provide the necessary space for the analyser installation.

The on-belt analyser unit weighs approximately 4.5 tonnes, mainly because of the large amount of shielding required for the neutron source. A mount was built to hold the analyser in place independent of the conveyor belt structure. The new support also included a weatherproof enclosure and platforms to allow access to either side of the analyser for maintenance and source isolation Fig. 6 On-belt Analyser Installation at Sepon Copper Operation.

Although radiation sources are common in mineral processing plants in density gauges, level probes and slurry analysers, they typically contain sources that emit gamma radiation only. The californium 252 source present in the on-belt analyser predominantly emits neutron radiation, although significant quantities of gamma radiation caused by neutron radiation can be detected also. The introduction of this new hazard required extensive radiation surveys to be conducted around the analyser with the source both in the on and off position. The analyser is fitted with an automatic source drive that can move the source into an area of heavy radiation shielding when not operating. A restricted area was created around the equipment, new procedures for area access were introduced and an intensive education program was conducted for all personnel working in this area. New dosimetry badges for measuring neutron radiation exposure were also introduced to ensure that systems were working effectively.

Fig. 6 On-belt Analyser Installation at Sepon Copper Operation Assay data provided by the on-belt analyser was linked to the Process Control System (PCS) and the Plant Information (PI) system for monitoring in conjunction with other process data.

The installed analyser unit had been factory calibrated on static samples of Sepon ore. The challenge was now to verify that the analyser was producing data that accurately represented the composition of the material feeding the copper mill. In other applications the analyser is typically used in conjunction with an automatic belt sampler that can be used to collect a representative sample of conveyor belt feed over a fixed time period, however a belt sampler was not available at


Sepon. A standard manual conveyor belt sample was not suitable as the goal was to verify the dynamic data rather than the analysis of a static sample. Alternatively spot samples were taken of cyclone overflow and CCD overflow liquor, and the assay data from these streams were used to recalculate the copper mill feed grade. This data was then compared to the rolling 20 minute average result from the belt analyser to allow for the residence time in the copper milling circuit. Encouragingly these spot samples correlated well with the on-belt analyser results Fig. 7 On-belt Analyser Performance Check.

Fig. 7 On-belt Analyser Performance Check Since commissioning of the on-belt analyser in mid 2008 the additional data has enabled necessary ROM blend changes to be made much more rapidly and with an increased level of confidence. Fig. 8 Grade Correction Using On-Belt Analyser Data illustrates a typical example where a major unscheduled change in the mill feed copper content was observed. Within twelve hours the problem was rectified without the performance of the leach circuit being affected. Prior to the on-belt analyser being available the first indication of this feed variation would have been a high leach tail assay. Copper would have already have been lost to tailings before changes to rectify the problem could be made.


Fig. 8 Grade Correction Using On-Belt Analyser Data Although the analyser was justified on copper analysis, data for multiple other elements is also provided. In particular the sulphur and iron data has been used to identify the copper associated with either the sulphidic or carbonaceous ore and to ensure that the feed contains adequate pyrite to feed the pressure oxidation circuit. Changes in the percentage carbonate ore in the feed will have implications for the circuit acid requirement and also the settling characteristics of the CCD circuit. Early identification of a variation from the scheduled blend can again enable changes to be made prior to the impact being observed in the operation of the plant. The circuit residence time results in a delay of over 24 hours between feed entering the milling circuit and reaching the flotation plant. Low levels of pyrite in the feed can be identified early and the mill feed supplemented with pyrite ore prior to the impact being observed in the flotation and pressure oxidation circuits.

The condition of the on-belt analyser is remotely monitored as part of a service agreement which includes bi-annual site visits by the vendor service engineer. The equipment has proven to be mechanically and electrically reliable and there have been very few occasions when analyser data has not been available. Calibration of the equipment has not been a major focus since installation, as the analyser is mainly used to identify feed grade variance rather than determine an absolute value.


CONCLUSIONS High feed grade variability and the requirement to mill in acidic cupric process solution has resulted in a significant challenge to the Sepon copper process. The additional data provided through on-belt analysis has enabled changes to be made before the problem impacts copper recovery. The data has also assisted the ROM management team in verifying data from their block models, identifying problems with stockpile control and rectifying them before they impact the copper process.

ACKNOWLEDGEMENTS The authors wish to thank the processing, mechanical and electrical teams and SCANTECH who all contributed to the successful installation and commissioning of the on-belt analyser. The authors also wish to thank MMG for permission to publish this paper.

REFERENCES Keokhounsy, S, Moore, T, and Liu, M, 2006. Hydromet at Sepon, ALTA Conference, Perth Western Australia, Australia 2006.


The Casposo Gold-Silver Project – Process Selection and Design

D Connelly1 and K Nilsson2

1. MAusIMM(CP), Director/Principal Consulting Engineer, Mineral Engineering Technical Services Pty Ltd (METS), PO Box 3211, Perth WA 6832. Email: [email protected]

2. Executive Director Operations, Troy Resources NL, 44 Ord Street, West Perth WA.

ABSTRACT

Silver and gold bearing ore from the Troy Resources NL Casposo Gold-Silver Project,

located within the San Juan Province of Argentina contains a complex mix of gold, silver,

electrum and silver sulphides in the ore. Casposo is a typical low sulphidation epithermal

style gold–silver deposit where mineralisation is hosted within rhyolite – andesite flows and

breccias. Veins are typically banded quartz–chalcedony colloform - crustiform banded with

quartz - carbonate infill. Mineralisation is associated with an assemblage consisting of

quartz, chalcedony, adularia, calcite, illite, sericite and trace sulphides. Gold and silver occur

as electrum, native silver, sulfosalts and silver sulphides. This paper presents a case study

that describes the process selection, development and engineering of a tailored process

plant for the recovery of gold and silver from this ore.

The original feasibility study, modelling, metallurgical testwork and engineering studies are

outlined. The final flowsheet was based on minimising capital cost, operability, recovery and

minimising technical risk.

The process flowsheet uses the traditional Merrill Crowe route because of the high silver to

gold ratio in the feed and achieves maximum gold and silver recovery while providing a

common facility for future ore body treatment.

INTRODUCTION

The Casposo Gold-Silver Project is Troy Resources NL third mine in South America (Fig 1).

The plant is designed to treat a minimum of 400,000 tpa of ore. The current mine plan

envisages production peaking at over 110,000 of gold equivalent in any single year.



Fig 1- Casposo Gold-Silver Project Location

Like many resource projects, the Casposo Gold-Silver Project was prohibited by high

capital costs. The financial model developed by the previous owners was based on

the assumption of building a completely new processing plant. A major challenge in

order to make the project a reality was to develop an innovative approach to the

project to reduce the capital expenditure (CAPEX) so as to improve the project’s

fundamental economics.

The results of the comminution testwork were used to provide an assessment of the

suitability of an existing second-hand crushing and grinding circuit for the application.

Simulation software (JKSimMet) was used to model the comminution circuit

providing a high level of confidence. Mineral Engineering Technical Services Pty Ltd

(METS) evaluated and sourced suitable second-hand equipment from multiple

locations for the rest of the plant. A three-dimensional computer model of the plant

was created using SolidWorks to allow proper layout of the plant and to provide a



basis for generating general arrangement drawings for on-site engineering personnel

(Fig 2).

Fig 2- 3-D plant model

GEOLOGY

The San Juan province straddles 3 major north-south trending ranges, the Cordillera

Principal, Cordillera Frontal and Precordillera, as well as part of the Sierras

Pampean range. The Casposo Gold-Silver Project is located on the eastern border

of the Cordillera Frontal range, separated from the Precordillera by the Rode-

Calingasta-Upsallata Valley.

The Cordillera Frontal range is underlain by marine sediments (shale, sandstone and

conglomerates) of the Carboniferous Cerro Aqua Negra Formation. These

sedimentary materials are overlain by a thick intrusive and volcanic sequence

assigned to the Permian-Triassic Choiyoi group. Basal andesitic volcanic flows, tuffs

and breccias are the main sub-surface unit in the Casposo Property and are overlain

by rhyolite breccias, rhyolite-dacite flows and dacitic ignimbrite flows.



Mineralisation at the project occurs along a 7 km long northwest- southeast regional

structural corridor, with the main Kamila vein system forming a sigmoidal set 500 m

long near the centre. The Mercado vein system is the north-westerly continuation of

Kamila, separated by an east–west fault.

The mineralisation identified at the Kamila and Mercado as well as other prospects

within the Casposo Gold-Silver Project are examples of low-sulphidation epithermal

gold and silver deposits.

Mineralisation characteristically comprises pyrite, electrum, gold, silver, and

argentite. Other minerals can include chalcopyrite, sphalerite, galena, tetrahedrite,

and silver sulphosalt and/or selenide minerals.

The gold–silver mineralisation at the Kamila deposit is structurally controlled and

occurs in crustiform- colloform quartz veins and stockworks in both andesites and

rhyolites. The Kamila Deposit consists of the following features:

Main Corridor NW-SE Structure > 200 metres long with two parallel

Veins Sets Dipping 65o-55o SW (B Vein & Inca Veins),

Sigmoidal Structures N-S – Dipping 70o-60 W (Aztec Vein),

Ore Shoots: Variable length ~50 – 200 metres – Lenticular Bodies; and

Variable width from 1 to 15 metres.

Vein alteration is characterised by strong to pervasive silicification. Banded quartz–

calcite veins with lattice bladed textures are common in the andesites. Interpretations

of the drill core show that mineralization is vertically zoned.

Mineralisation within the Mercado vein system contains moderately higher base

metal values, as well as increased amounts of iron and arsenic sulphides, in

comparison to the Kamila deposit.

FEASIBILITY STUDY

A Feasibility Study (FS) for the Casposo Gold-Silver Project was commissioned by

the previous owners in 2005 and completed by the Peruvian offices of EPCM

services consultancy in 2007. Some of the key findings are summarised in this

section.



The 2007 Feasibility Study was based upon the proposed mining of two deposits –

Kamila and Mercado. The Kamila deposit is larger and would be mined first. It was to

be mined by open pit methods and the deeper portion by underground mining

methods (2007).

Total Indicated Resources at the two deposits are as currently quoted and updated

on the Troy Resources NL website (www.troyres.com.au).

The plant and facilities were designed for an average milling rate of 1000 t/d (365000

t/a). Projected overall gold and silver recoveries were 93.7% and 80.6% respectively.

Over a mine life of five years, average annual production was estimated at 50,478 oz

of gold and 1.1Moz of silver per year, over a mine life of five years.

The process flowsheet was designed to use conventional primary jaw and secondary

cone crushing, ball milling, gravity concentration for coarse gold and silver, cyanide

leach, counter current decantation and washing and dewatering of tails by belt

filtration. Gold and silver would be recovered by standard Merrill Crowe precipitation

and smelted to produce doré bars. Tailings solids are washed and rinsed on a belt

filter to remove cyanide and then filtered tailings trucked to a lined tailings storage

facility.

Major infrastructure was required to develop the project including site roads, diesel

power generation plant, fuel storage, water supply, contractor areas and

administration facilities.

In summary CAPEX plant costs were US$44.2M as at 31 December 2010. Financial

analysis of the project suggested it was more sensitive to changes in metal price and

grade than either capital or operating costs.

METALLURGICAL TESTWORK

Several exploration and drilling campaigns have been undertaken at Casposo since

1998 in order to establish a resource estimate and map the deposit.

Thorough, accurate testwork is vital to producing a viable economic project and a

process which will not provide unwanted surprises once production is underway. A

variety of metallurgical testwork has been undertaken on Casposo ore over the last



10 or so years (METS, 2009a and METS, 2010a). A brief summary of this work is

presented here.

Communition

Typical comminution testwork was completed to establish Crushing Work Index,

Bond Work Indices, Abrasion index and semi-autogenous grinding (SAG) mill

comminution parameters on composites of typical ore. These typical figures were

used to evaluate crusher power requirements, milling power requirements, indicative

liner wear rates and SAG mill breakage parameters respectively. Generally, the

Casposo ore was very hard and very abrasive. Tests including unconfined

compressive strength (UCS) and point load strength (PLS) tests were also

undertaken to identify the strength of the ore. The PLS test is a simpler version of the

UCS but both results are used to determine the rock strength. The results of the two

tests were similar and indicate the ore is very strong.

Gravity Testwork

Several gravity tests were undertaken in various testwork programmes. The Master

Composite sample responded well to gravity separation in May 2006. A gold yield of

33% was achieved. A further study later that year showed similar responses with

other samples yielding recoveries from 26% to 34%. The latter study showed

increased recoveries were exhibited by increased gold grades. The result indicates

the ore contains free milling gold.

Flotation

Flotation test on the Master Composite concentrate was conducted at two grind

sizes of P80 168 and 106 µm. The results showed that improved recoveries were

achieved with finer grind. A recovery of 72% Au was achieved at P80 168 µm. The

highest recovery was performed at P80 106 µm, yielding 80% Au. The results

achieved are regarded as being low against industry values. Gravity / flotation

testwork conducted in September 2005 showed similar results. Improved recoveries

were achieved with finer grind of the primary feed. Overall recoveries using the

gravity / flotation regime yielded 86.2% Au at P80 106 µm and 73.3% Au at P80 150

µm.



Head Assay

A head assay performed in 2010 gave gold and silver grades of 8.9 and 157 g/t

respectively. These were in agreement with previous analyses. The high ratio of

silver to gold in the ore indicated that a Merrill Crowe process option be considered.

Leach Testwork

A range of leach testwork performed on the samples showed the leach kinetics of

the gold are significantly faster than the leaching kinetics of the silver. This is likely

due to the presence of significant quantities of silver sulphide minerals. The impact

of this is that a longer residence time is required in order to attain acceptable silver

recoveries. A longer residence time requires larger equipment volumes and therefore

higher capital cost.

Other key results included:

Good recoveries were achieved using oxygen / air.

High cyanide concentration improved the leach kinetics of silver

significantly, although it did not alter that of the gold.

No benefit in recoveries with zinc addition. Zinc addition caused

increased reagent consumption as zinc is a cyanide-consuming

species.

High levels of iron in solution occurred.

Finer grinding can improve recoveries.

The high levels of iron in solution incurred a negative impact on

reagent consumption. The addition of a gravity circuit within the

process can mitigate this problem.

Several tests were carried out to assess the effect of lead nitrate on the silver

leaching kinetics. The lead nitrate did not significantly improve the leaching kinetics

of the silver as expected.

Gravity plus cyanide leaching was chosen as the preferred route as it gave the

highest recoveries and least variable results.



Settling and Filtration

Thickening testwork on the Mercado and Kamila ore samples yielded poor solution

clarity.

Filtration tests were conducted on the leach residue using both flocculant and guar

gum (filtration aid). The aim of these tests was to determine the impact of each on

the filtration rates. Flocculant and guar added individually both reduce the cake

formation time thereby improving the filtration rate. However the rate improved

further when the flocculant and guar gum are added together.

Zinc Dust Cementation (Merrill Crowe Process)

Some operating data is available in the literature including reagent addition rates and

solution quality requirements. However some key process parameters such as the

influence of zinc dust particle size on the reaction kinetics are not well understood.

Further tests were conducted looking at different zinc dust particle sizes. Finer zinc

dust accelerates the reaction kinetics. However if too fine a product was used it may

impact on the filter press cycle times by blinding the filter press cloths.

Cyanide Destruction

The cyanide destruction process is important in maintaining a low level of cyanide in

the tailings as required by the environmental authorities.

The INCO SO2/Air cyanide destruction process was investigated. This uses a

combination of SO2 and O2, in the presence of soluble copper, to oxidise cyanide.

Both free cyanide (CN) and cyanide weakly complexed with metals such as copper,

zinc and nickel are oxidized to form cyanate ions (OCN-), which are less toxic by two

orders of magnitude. During the process, the ferro-cyanide complexes are removed

as insoluble ferro- cyanide salts. In addition, a small amount of thiocyanate (SCN) is

oxidised.

Continuous tests were conducted to determine what levels of free cyanide could be

achieved. The kinetics of the reaction depends on the oxidising conditions. In a

highly oxidising environment it is expected that the reactions would proceed quickly.

A key measure for the oxidising conditions is the ratio of the mass of SO2 to that of



the weak-acid-dissociated CN. Testwork suggested this ratio should be maintained

above 3.

MINIMISING THE CAPITAL COST- SECOND HAND EQUIPMENT

The front end of the processing plant is based on second hand equipment acquired

previously from the McKinnon’s Gold Project in Cobar New South Wales which was

already owned by Troy Resources NL. Other second hand plant was also

incorporated wherever possible in order to achieve a financially viable project. The

second hand equipment also had the advantage of being immediately available. In

addition second hand filters and thickeners were purchased. Essentially, the use of

the second hand plant reduced the capital cost of the plant.

PROCESS MODELLING

An important stage in the development of the Casposo Gold-Silver Project was the

modelling and simulation of the McKinnon’s crushing and grinding plant that was to

be incorporated into the Casposo circuit.

The ore treated by the McKinnon’s plant was a medium-hard ore at a Bond ball mill

work index of14.4 kWh/t (METS 2009b). The McKinnon’s plant did not incorporate a

pebble crusher and there were concerns of the possibility of scat formation when

treating the harder Casposo ore (METS 2009b).

To investigate the means of overcoming the possible generation of large amounts of

scats and the resultant need for a pebble crusher at the Casposo Gold-Silver

Project, computer modelling of the circuit was conducted. This provided a means to

model a number of different scenarios which could be simulated before

implementation. Since JKSimMet software is widely regarded as the industry

standard for comminution simulation, it was employed to simulate the circuit. The

simulation was first initiated by first re-creating the McKinnon’s circuit and then

incorporating a pebble crusher into the circuit (METS 2009b).

FLOWSHEET

The processing plant will handle in excess of 400,000 tpa of ore (METS 2010b). At

8000 working hours per annum, this is equivalent to 50 tph (METS 2010b). The



crushing section will operate at a higher rate than the rest of the plant to allow for

reduced operating hours and equipment maintenance. Thus the crushing plant is

designed to operate at a feedrate of up to 110 tph (METS 2010b).

Crushed ore entering the grinding circuit is fed into the SAG Mill Feed Chute

together with recycled grinding circuit water. It is added at a rate to maintain a mill

discharge density of up to 65% solids by weight. The

SAG mill operates with a ball charge up to 5% by volume. Mill scats report to the

scats crusher, and then crushed scats are fed back to the SAG mill feed belt.

Slurry from the SAG mill is fed to a cluster of six 250 mm hydrocyclones. The

majority of the cyclone underflow returns to the mill and cyclone overflow reports to

the leach feed thickener.

A split stream from the cyclone underflow, feeds the Gravity Circuit Screen. The

Gravity Circuit Screen oversize is combined with the gravity concentrator tailings and

returned to the SAG Mill Discharge Hopper.

The Gravity Concentrator processes the screen undersize. This concentrator retains

a gravity concentrate and produces a gravity tail. The gravity tail is returned to the

comminution circuit via the SAG Mill Discharge Hopper for further liberation. When

the gravity concentrate becomes sufficiently enriched, it is transferred to the

Intensive Leach Reactor (ILR) for leaching under intensive conditions.

The ILR is a specialist precious metals leach unit that uses alkaline cyanide solution

to leach gold and metallic silver from the high-grade gravity concentrates. The

concentrates are collected in the Concentrate Feed Tank. When sufficient

concentrate has been collected, the batch is transferred to a horizontal rotating drum

together with barren solution and hydrogen peroxide. Cyanide solution, caustic

solution and lead nitrate solution are added to the reactor. The pregnant solution is

then pumped to join other streams entering the Clarification area to remove any

solids and recover the gold and silver.

Thickened cyclone overflow slurry is pumped to the leach tanks. There are nine

agitated leach tanks linked in series. The total volume of the tanks is designed for a



total of up to 80 hours of leach residence time depending on the leach throughput.

Due to the slow leaching kinetics of silver, an elevated cyanide concentration needs

to be maintained in each of the leach tanks. Cyanide solution is added to Leach

Tank 1 to achieve a nominal concentration of 2200 ppm and ensure a sufficient

cyanide concentration is maintained throughout the leach train.

The leach slurry feeds the first counter current decantation (CCD) thickener where

the solids settle to the underflow for pumping to the second CCD thickener. The final

underflow from the second CCD thickener is filtered using belt filters.

Two Tailings Belt Filters operate in parallel to produce a filter cake of <14% moisture

for tailings disposal at a nominal rate of 46 dt/h and a maximum rate of 55 dt/h.

The overflow from CCD Thickener 1 is pumped from the CCD Thickener Overflow

Tank into the Pre- Clarification Hopper. The Merrill Crowe process requires low

levels of solids, <1 ppm, to be effective. The solution from the CCD and filtration area

is clarified to achieve the target solids content. The Pre-Clarification Hopper removes

fine solids carried over from the thickeners by forming a bed of flocculant through

which the dirty solution is passed upwards. Following deaeration, the clean, oxygen-

free pregnant solution is mixed with zinc dust, in the form of slurry, to recover the

gold and silver by cementation.

Lead nitrate is added to prevent passivation of the surface of the zinc powder and to

promote dendritic growth. This is important for good recovery of silver.

The gold and silver precipitate is recovered by filtration. The precipitate from the

Merrill Crowe process contains a mixture of gold, silver, un-reacted zinc and minor

amounts of mercury. The mercury removal retort initially removes surplus water and

ultimately brings the charge material up to the target treatment temperature of

750°C. At this temperature any mercury in the precipitate is volatilised.

After the mercury has been captured and the retort has cooled the remaining

precipitate is removed from the retort and further cooled. The precipitate is emptied

into the Merrill Crowe Precipitate Feeder which transfers the material into the

Precipitate Leach Tank. Dilute nitric acid is pumped into the leach tank to dissolve

the silver and zinc. The gold remains as a solid residue.



When the leaching reaction has stopped, the solution is filtered to recover the gold.

This is dried, then smelted and cast into ingots.

Sodium chloride is added to the filtrate to precipitate silver chloride and zinc chloride

remains in solution. The silver chloride precipitate is recovered by filtration. The

silver chloride is dried, smelted and cast into ingots.

A portion of the barren solution is passed through the cyanide destruction area. In

this area the levels of free and weak acid dissociable (WAD) cyanide ions are

reduced using the Inco air/SO2 process. The cyanide destruction process also

removes some of the soluble zinc with the remaining zinc removed by a sulphide

precipitation process. This is important for silver recovery.

Merrill Crowe

The key reasons for the selection of this process was that the silver to gold ratio in

the Casposo ore at about 24 to 1 meant it is not an economical process to use either

Carbon In Pulp (CIP) or Carbon In Leach (CIL) methods, due to the high amount of

carbon required for stripping (METS 2010b). The Merrill Crowe process was included

in the Casposo process flowsheet so as to not use carbon thus reducing the

stripping cost (METS 2010b).

Several other reasons are also influenced the decision to choose the Merrill Crowe

process.

Adsorption

The silver is present at a much higher grade than the gold. This silver will compete

for active sites on the carbon so may lead to gold losses. The adsorption capacity of

the total carbon inventory needs to be higher and more carbon will need to be moved

to keep the active sites available. Carbon activity and solution losses become more

critical with competition for the active sites so more carbon contact stages may be

required in order to achieve adsorption of the valuable materials.

Stripping

Silver is removed from carbon before gold and at a lower temperature. These

conditions are at odds with the gold elution so operations operate to maximise gold



recovery from the carbon. The silver cyanide complex is less stable and breaks

down at the high temperature, forming metallic silver on the carbon. This has a

negative impact on carbon activity and hence increases solution gold losses.

Electrowinning

The electro-potential for gold is -0.63 volts and for silver it is -0.45 volts. This means

the silver will preferentially plate out of solution before gold thus slowing the

electrowinning process. In addition there is about 24 times more silver metal to

electrowin so more cell capacity is required.

Further, in this stage of the process, a review was undertaken to evaluate whether to

produce a combined gold-silver doré, or to separate the two metals. It was decided

that in order to minimise the refining charges, gold / silver separation should follow

Merrill Crowe precipitation as described above.

PROCESS CONTROL- DELTA V SYSTEM

The DeltaV control package was chosen as the process control system for the

project. The DeltaV distributed control system (DCS) provides an interface to the

programmable logic controller (PLC) for control and monitoring of the plant. Due to

the remoteness of the site, lack of skilled personnel and cost constraints, a number

of criteria were identified. These criteria included, but were not limited to the

systems:

Field-proven capability,

Capability of expansion without interrupting the process,

Ability to control the Merrill Crowe precious metals recovery, and

Ease of installation and operation.

The DeltaV system ticked all of these boxes and other plant operating

requirements.

Three particular issues were resolved by this choice:

1. Hardware The modular design allowed for the purchase of the exact number of

I/O cards, carriers, workstations, controllers etc, and at the same time

gave the option to add or remove any of these components “on-the-fly”



as the system might require, with no resultant downtime,

Regardless of I/O type, the setup prevents a mix-up of cards with

carriers. That means far less mistakes and field work,

The DeltaV has proven itself to be rugged and can be to mounted

almost anywhere including hazardous environments and extremes of

temperature, and

DeltaV can be maintained and operated via remote networks including

local area networks(LAN), wide area networks (WAN), dial-up,

microwave, or satellite communications. This proves important for

remote maintenance support in areas where technical support on the

ground is limited and hard to get.

2. The operator is aided with easy to use “day-to-day” operation tools such as: DeltaV Operate Run,

Process History View, and

Diagnostics.

3. Techniques like drag-and-drop to simplify system modifications; Windows

looks alike and Plug-and- play technology for hardware configuration,

allows anyone to operate the system with only basic training.

LESSONS LEARNT

A major lesson from the successful completion of this project was the importance of

modelling to clarify and confirm the suitability of the second hand equipment. The

JKSimMet process modelling confirmed that the crushing and grinding plant would

be suitable, but that the scats crusher would be a necessary addition. Similarly, it

confirmed the importance of the metallurgical testwork in evaluating the comminution

indices and evaluating the variability of each of these.

Merrill Crowe is not as user friendly as CIP. Settling and clarity of overflow solutions

is paramount in successful solution processing and precipitating gold. The use of

tailings filtering is subject to variability in throughput due to different ore

characteristics i.e. on filtration rates, the effect of weathered ore (low) versus fresh

free quartz veins (high).

Water management is also critical particularly when one part of the process is taken

off line.



The need to consider innovative and “outside-the-box” solutions is also highlighted

by the success of this project. Being able to halve the capital cost by utilising second

hand plant was imperative in getting this project to fruition. Selecting the Merrill

Crowe route over CIP or CIL processes was another. It is unclear how many projects

consider these options, and often it may not be available, but if it presents a real

possibility, it may make all the difference.

CONCLUSIONS

The Single Stage SAG is applicable to this ore and optimisation of the cyclones, the

grate and liner/lifter configuration and speed will achieve improved throughputs over

time.

Merrill Crowe is not as straight forward as CIP and there are a number of

disadvantages with Merrill Crowe but for high silver ores there is no other choice.

Managing the water balance presents challenges as do weathered ore when it

comes to settling and filtering. The importance of clean solutions for processing and

precipitation is also a challenge if the ore is not fresh. New technology such as resins

and EMEW cells will be looked at now that the project is operational and the

technical risks of such technology can be managed.

The concept of minimal engineering, (i.e. project management of a group of sub-

consultants to execute the project) did result in substantial savings in time and

money.

The use of whole tailings filtration was dictated by the seismicity and sensitive

environmental area the project was located in making a conventional tailings dam

impractical.

The project impact has been very positive for Troy Resources NL and the local

community. Exploration in the region has discovered new mineralisation and the

operation may continue production beyond the present resources.



ACKNOWLEDGEMENTS

The authors would like to thank Troy Resources NL for permission to publish this

paper and also all colleagues and engineers at various sites, METS staff and other

consultants for their contribution and the management of METS for their permission

and constructive criticism of various drafts of this paper.

REFERENCES

2007. Casposo Project Technical Report On Feasibility Study: unpublished report to

previous owners.

METS 2009a. Casposo Gold-Silver Project: Previous Metallurgical Testwork

Summary: unpublished report to Troy Resources NL.

METS 2009b. Casposo Milling Circuit Simulation Report: internal report to Troy

Resources NL.

METS 2010a. Casposo Gold-Silver Project: Metallurgical Testwork summary report:

internal report to Troy Resources NL.

METS 2010b. Casposo Gold-Silver Project: Process description: internal report to

Troy Resources NL.

www.troyres.com.au, 2011.


8 ‐ 9 August 2011 Perth, WA

Optimising Western Australia Magnetite Circuit Design

D David1, M Larson2 and M Li3

1. FAusIMM, Process Consultant, AMEC Minproc, Level 14, 140 St Georges Terrace, Perth WA 6000. Email: [email protected]

2. Senior Process Engineer, Xstrata Technology, 5th Floor, 509 Richards Street, Vancouver BC, V6B 2Z6, Canada. Email: [email protected]

3. Senior Metallurgy Manager, Grange Resources, Level 11, 200 St Georges Terrace, Perth WA 6000. Email: [email protected]

ABSTRACT The development of Western Australian magnetite deposits has lead to the design of some of the largest grinding mills and plants in the world. One of the projects demonstrates the efficiency gains possible by developing a simple yet thorough test program for circuit design. By drawing on the experience of current magnetite operations in Australia and the Mesabi and Marquette iron ranges in the United States, a basic flowsheet was developed. Through comprehensive testwork with AG, ball and stirred milling the flowsheet was optimized to take full advantage of each grinding mill’s strengths to reach the required final grind size. Laboratory work was verified in the pilot plant to optimize the energy efficiency of each grinding step while ensuring adequate liberation at each step for sufficient gangue rejection. By using three stages of grinding, the ball mill can best be employed in ensuring all top size gangue material is liberated and removed in the second magnetic separation step. The inclusion of the IsaMill, with its inherent steep product size distribution, as the tertiary grinding stage ensured that maximum grade was achieved and simplified the downstream process while giving further improvements in total grinding capital and operating costs. In this way the combination of the two technologies downstream from the AG mill is far more efficient than either would be on its own by reducing the total installed power by 1/3 and annual grinding media cost by 2/3.

INTRODUCTION In the past decade the interest in Australia’s iron ore deposits has shifted to include the vast magnetite deposits scattered throughout Western Australia. These deposits have resulted in numerous magnetite concentrators currently in the design and construction phase. Previous magnetite concentrators in Australia have been limited to smaller operations such as Savage River and OneSteel’s Project Magnet.

With a feed tonnage of 3600 tph and a final grind P80 of 34 µm the amount of grinding required for this project will be extensive. A lower than typical final silica grade of sub 3% SiO2 is desired. Finally a small amount of pyrrhotite is present in the ore requiring a final flotation step to meet sulfur requirements. In this case the complete flowsheet is examined testing previously proven technologies from existing plants with the aim of optimizing each step of the process. Laboratory and pilot work is combined to ensure maximum economical efficiency while still maintaining a quality product.

TYPICAL MAGNETITE CONCENTRATOR DESIGN Five magnetite concentrator flowsheets are shown below from Minnesota (Figs. 1, 4 and 5), Michigan (Fig. 3) and Australia (Fig. 2). Though they have a mix of rod, ball and autogenous milling (depending on what was commonplace or available when the plants were designed) two steps in the process should be pointed out.

The use of a hydroseparator is commonplace in plants with and without reverse silica flotation. This removal of slimes is absolutely necessary prior to a silica flotation step and also aids in removing fine silica that would be more likely entrained in a magnetic concentrate.



Whether called a hydroseparator, hydro-sizer, siphon-sizer or thickener sizer the basic principal is the same. A “thickener” is employed but with an upflow of water added. The coarser, heavier magnetite settles to the bottom and the upflow of water carries the fine silica to the overflow. A magnetic flocculator can also be used on the feed to increase the settling rate and minimise chemical flocculant consumption.

The fine screens are necessary to recirculate back to grinding the +105 µm material that would negatively impact the final concentrate grade, and, in the case of Savage River, result in unaceptable concentrate pipeline wear. In the case of the east Mesabi ores such as Minntac, a concentrate ground to a P80 of 325 mesh (44 µm) will typically have a grade of 8% SiO2 with half of that silica in the 10% of the mass making up the +140 mesh (105 µm) size fraction. By using fine screens that size fraction can be removed and the final concentrate grade reduced to 4-5% silica.

Of all these operations, only the Empire Mine flowsheet does not use fine screens. The Empire Mine grind results in a final P80 of 20 µm, making finisher screens both impractical and unnecessary. The Empire Mine is also the only one of these plants shown to use reverse flotation as a final upgrade step.

Reverese silica flotation is usually seen as a last resort when processing magnetite as there is a certainty that magnetite will be lost to tails. The finisher magnetic separators will have rejected all liberated silica (except for silica slimes that follow the water in the process) ahead of flotation. Consequently, virtually all coarse floated silica will have magnetite attached to it. Amine flotation of silica also presents issues with slimes and is particularly sensitive to water chemistry. While hydroseparators should probably be installed regardless to remove the slimes ahead of finisher magnetic separation, the added hydroseparator benefit of reducing soluble salts is only of benefit when a silica flotation step is performed. According to Iwasaki (1983), the fatty acid flotation step for iron ores is quite sensitive to magnesium and calcium ions in pulp solutions and for optimal performance the level of these ions should be kept below 100 ppm.

Figs. 1 and 2: Erie magnetite and Savage River flowsheets (Devaney, 1985)

.



Figs. 3 and 4: Empire Mine and Hibbing Taconite flowsheets (McIvor and Greenwood 1996)

Fig. 5: Minntac flowsheet (Devaney)

TESTWORK PROGRAM The testwork at AMMTEC for this Western Australian magnetite deposit consisted of:

Pilot autogenous primary milling,

laboratory work (Levin test) and pilot secondary ball milling,

laboratory and limited continuous secondary IsaMill testing,

laboratory tertiary, limited continuous and pilot IsaMill testwork,

Davis tube and pilot magnetic separation tests of the different intermediate and final products,

hydroseparating tests of the final IsaMill magnetic concentrate,

sulfide flotation tests of the final magnetic concentrate, and

final concentrate filter testing by vendors.

Autogenous primary milling All of the ore was ground in the AMMTEC pilot AG mill (1.74 m diameter inside liners and 0.46 m EGL) and then run through the magnetic separator. To protect the magnetic separator the AG mill was closed with a 1 mm screen. In the full scale plant it is intended that the AG mill will be closed at a coarser size, up to a maximum of 4 mm. The full scale plant is required to process 3600 tph of of feed, ideally through two grinding lines, each with a large dual pinion or gearless drive primary



autogenous mill. With the energy available from the pilot AG mill operating at about a 25% filling the P80 of the magnetic separator feed (-1mm) was approximately 330 µm. The magnetic concentrate was significantly coarser than the non-magnetics with P80’s of 420 µm and 200 µm respectively. This first stage of magnetic separation removed 40% of the total mass as barren tails. The remaining 2200 tph is fed to the second stage of grinding. Additional analysis suggests that the P80 of the magnetics will increase from 420 µm to 770 µm when a 3 mm screen is used at full scale compared to the 1 mm pilot plant closing screen.

Ball Mill Results Prior to any pilot ball mill work, one Levin test was completed on the secondary grind feed shown in Fig. 6. The Levin test is a fine-feed substitute for a standard Bond Grinding Work Index test and it uses sequential open circuit milling passes applying a known energy at each pass. The results from this test, with the P80 size in microns plotted against the specific energy (kWh/t) is shown in Fig. 6.

Fig. 6: Levin test for the ball mill

The Levin test for fine ball mill product predicts a net specific energy requirement of 52.5 kWh/t to grind from an F80 of ~400 µm (AG mill/magnetic separator product) to a P80 of 34 µm. An additional 2.3 kWh/t would be needed to reduce the plant feed P80 of 770 µm to the test feed P80 of 420 µm, giving a net energy requirement of 54.8 kWh/t. However, as the Levin Test does not incorporate a classification step to remove fines, energy is unnecessarily used grinding material that has already achieved final target grind size or less. Consequently, it is expected that the test will over-predict the specific power requirement for most duties, especially over wide size ranges. The actual results obtained with the AMMTEC 6’ pilot ball mill with 25 mm top size media was a P80 of 37 µm using between 40 and 45 kWh/t of energy. At 2200 tph of combined feed for lines 1 and 2 and incorporating the 2.3 kWh/t to reduce from 770 µm to 420 µm, using single pass ball milling from 770 µm to 34 µm would require about 114 MW of ball mill energy, excluding that associated with cyclone feed pumps. This is equivalent to 6 of the largest ball mills in existence, 3 per line. In addition approximately $86 million would be spent on steel grinding media annually. The high media consumption is due to an unusually high average Bond abrasion index of 0.44, caused by the presence of gangue silicates and garnets. Typical magnetite ore will have an Ai of 0.25 or less. The requirement to use 25 mm steel balls to reach the required fine grind exacerbates the media wear rate. A $15 million capital cost would also be incurred in any ball milling circuit targeting a 34 µm product for finishing screens which are needed to minimise silica and to protect the concentrate pipeline. The ball mill cyclone combination cannot guarantee elimination of +100 µm particles from the pipeline feed.



IsaMill Secondary Grind The same feed as was tested in the previous Levin test was run through an M4 IsaMill to generate a signature plot. A graded 5 mm ceramic media charge was used to ensure top size breakage. The signature plot test consists of two tanks, one feeding the mill and one acting as a discharge tank. In the middle of each pass a sample is taken for a sizing and density. The flowrate, time and net energy of the pass is recorded and the size is plotted against the specific energy (kWh/t). Valves are then changed to switch the feed and discharge tanks and the process is repeated multiple times, each pass incrementing the input energy. From this a signature plot is developed. This creates a straight line on a log-log size vs energy plot that scales directly to full scale IsaMills. The result of this signature plot test is show in Fig. 7.

Fig. 7: IsaMill 400 µm signature plot

The IsaMill was able to reduce the secondary grind feed of 400 µm to 34 µm using 31.7 kWh/t. With the 2.3 kWh/t required for reduction from 770 µm to 420 µm (in the previous ball mill test) the total predicted IsaMill power is 34 kWh/t (note that this calculation is theoretical as 770 µm is an impractical IsaMill magnetite feed F80 and the inefficiency of the IsaMills on this size of material would mean that 2.3 kWh/t is optimistic). This is an improvement in specific energy of 30% over the ball mill (for this theoretical coarser feed). This would result in an installed power of 78 MW if IsaMills were used compared to the 114 MW necessary for the ball mill. In addition, the annual grinding media cost would drop from $86M for steel balls to $57M for ceramic IsaMill media. Besides the abrasiveness of the ore, the galvanic action of the pyrrhotite present in the ore may be contributing to the high steel media wear (Iwasaki, 1999). This would not be a factor with ceramic media.

Even with this drastic improvement, as with the ball mill there are still issues with using the IsaMill in a single stage for this duty. Due to the coarse, hard garnet and gangue silicates present, and the friction created in mixing these particles, media wear was higher than expected. Despite showing better wear than the ball mill steel media, the IsaMill ceramic media wear was still about double the expected rate. The energy used in grinding this garnet down to 34 µm can be considered wasted. The garnet in question is a hard alumina silicate and actually has properties close to that of typical IsaMill grinding media. It would be much more efficient to remove these waste materials as close to liberation as possible. Fig. 8 shows a close up view of the garnet crystals fully liberated in the 125-150 µm size range.



Fig. 8: Liberated garnet after 150 µm grind

By plotting the Levin test and secondary grind IsaMill test together (Fig. 9) it can be seen that the gain in efficiency experienced with the IsaMill for this particular ore starts at a product P80 of approximately 100 µm. At product sizes above that point the ball mill is more efficient. This is mainly a function of media size. A media larger than 5 mm used in the IsaMill would achieve more efficient size reduction to 100 µm but would not be as efficient to the finer 34 µm final target. With this magnitude of size reduction applied to a fairly hard ore, grinding the entire stream in one step will never achieve optimum efficiency across all size fractions with any grinding technology.

Fig. 9: IsaMill/Levin Test comparison

The IsaMill improvement in efficiency at the fine sizes can be explained by the grinding mechanism and media used. At fine product sizes below 70-100 µm, attrition becomes the main grinding mechanism. The 5 mm media provides not only more surface area compared to the 25 mm ball mill media, and a greater probability for collisions than the larger steel media, but it is also stiffer than the

Liberated Garnet



steel media. The more elastic steel media does not transfer energy to the ore particles as efficiently as the ceramic media. The Vickers Hardness of the ceramic used is about 1000, whereas that for steel media will typically range from 500-600.

The efficiency changeover is also seen in Fig. 10 (Burford and Niva, 2008), a comparison of the IsaMill with the Tower Mill grinding an Ernest Henry magnetite feed taken from the copper concentrator tails. In this case below 70 µm the smaller harder IsaMill media is more efficient than the 12 mm steel media used in the Tower Mill. This test was conducted with 3.5 mm ceramic media. By using 5 mm media this intersection point would be shifted coarser.

Signature Plot - P80: IsaMillTM vs Tower Mill

1.00

10.00

100.00

1.00 10.00 100.00 1000.00

Size (µm)

Spec

ific

En

erg

y (k

Wh

/t)

P80 IsaMill Test 1

P80 IsaMill Test 2

P80 Tower Mill Test 1

P80 Tower Mill Test 2

Fig. 10: Ernest Henry Magnetite IsaMill/Tower Mill comparison

It is clear from Figs. 9 and 10 that ball or Tower milling with typical media is most efficient when applied to generating “coarse” products in the +80 µm range. Consequently ball milling was selected to bridge the efficiency gap that exists between AG milling with this ore to about 400 µm P80 and the application of IsaMilling to feed sizes of 100 µm and finer. The Levin test suggest that about 15 kWh/t is required to grind from the 400 µm feed P80 to a 100 µm product P80. Pilot testing achieved a ball mill grind (with 32 mm media) from 420 µm to 78 µm at a specific energy of 11.6 kWh/t, again significantly more efficient than the Levin Test prediction. When translated to full scale operation grinding 2200 tph from 770 µm to 100 µm, the ball mill installed power requirement is 34 MW, achievable in two large twin pinion ball mills (one per line).

By only requiring a product of 80 to 100 µm from the ball mill it gives the added benefit of being able to switch from the 25 or 32 mm steel balls used in the tests to 40 mm balls. The 40 mm balls will ensure adequate top size reduction when treating the 770 µm F80 feed and and will significantly reduce the steel media wear.

IsaMill tertiary grind The product from the intermediate pilot ball mill was subject to a magnetic separation step which removed a further 20% of the mass, the majority of this being quartz and some of it being garnet. This resulted in a less abrasive 76 µm feed to the IsaMill stage and this material was tested in the M4 IsaMill using the standard signature plot procedure and 5mm ceramic media. The results are shown in Fig. 11 indicating that from a F80 of 76.5um, the IsaMill energy required for a P80 of 34 µm is 12.4 kWh/t.

The coarse IsaMill test (Fig. 7) indicated that 13.4 kWh/t would be necessary to grind from F80 of 76.5um to P80 of 34um. This difference of 7.5% could be the result of a softer feed due to the extra magnetic separation step. It is also very close to the 5% margin of error associated with these tests.

This signature plot compares well with other iron ore IsaMill testwork completed on similiar iron ore feed with F80’s and energy adjusted to 76.5 µm as shown in Fig. 12 (Larson, 2011).



The Levin test indicates that the ball mill would require 30 kWh/t for the same size reduction. However, the difference between the two pilot scale ball mill trials (420 µm to 34 µm vs 420 µm to 100 µm) was only 22 kWh/t, a much better estimate of the ball mill power necessary to achieve the final grind. Again the Levin test has overpredicted the ball mill power requirement. The best availabe comparison between the two tested fine milling contenders is, therefore, 22 kWh/t for ball milling vs 12 kWh/t for IsaMilling, an advantage of 45% in specific energy alone to the IsaMill for this step. Additional installed power and capital cost savings are realized because cyclones are not required in the IsaMill circuit while they are necessary for ball milling.

Fig. 11: Fine magnetite feed IsaMill signature plot

Fig. 12: Common iron ore signature plots

This finer F80 76.5 um feed was also run through the M4 IsaMill in a short 180 kg continuous test and then continuously as part of the pilot plant. An energy of about 12-13 kWh/t was targeted by adjusting the speed of the mill. The 34 µm P80 target size was maintained through these tests including the top size reduction.

There was also a benefit seen in ceramic media wear in grinding this finer, upgraded material. Media wear was reduced by 35% per kWh as compared to the secondary IsaMill feed (420 µm F80).

Given the silica contamination and iron grade implications top size material can have, it is important to consider the complete product size distribution produced in regrinding. The size distribution from the signature plot pass that was closest to the 34 µm target is shown in Fig. 13.



Fig. 13: IsaMill product size distribution curve

The IsaMill product size distribution is 98.9% passing 75 µm and 100% passing 105 µm at a P80 of 37 µm. This removes the need for finisher screens which saves up to $15 million in capital costs together with operating costs associated with a potentially maintenance intensive piece of equipment. Typically 15 inch cyclones will produce a P98/P80 ratio of about 3-4 and this would defintiely necessitate finishing screens. In this case the IsaMill ratio is under 2.

Besides a natural improvement in concentrate grade without the need for fine screens, the steeper size distribution should have an additional benefit of producing less ultrafines. The intermediate magnetic separation step after the ball mill removes about 80% of the silica present in that stream. This results in less silica that has to be ground to 34 µm. This reduction in material that could be turned into silica slimes, along with the previously mentioned sharper size distribution, may result in a smaller hydroseparator design, reducing the footprint of one of the larger pieces of equipment in the flowsheet. Benefits should also be seen in filtering the concentrate. That testwork is still ongoing as of the writing of this paper.

Final Results A combination of bench and pilot testing was used to determine the most efficient means of reducing primary crushed magnetite ore to a target grind P80 in the region of 34 µm. Efficiency was maximised by a combination of using appropriate machines for the various grinding duties and also by rejecting barren mass as early as poosible in the flowsheet.

The AG milling stage was found to generate a relatively coarse product at pilot scale and this stream will only get coarser at full scale. The coarseness of the primary magnetic concentrate makes it unsuited for feeding to a fine milling device such as an IsaMill but it is ideal ball mill feed. However, the ball mill was found to be unsuitable for taking the ore from AG discharge to the final target grind in a single step as it becomes relatively inefficient for grinding finer than 80 µm. Testwork showed it was appropriate to use the ball mill to grind to about 100 µm P80, follow this with magnetic separation and complete the grinding to 34 µm in an IsaMill. Approximately 60 MW of power is saved (40% of total power and ~50% of the power for that grinding step) over the single stage ball mill circuit and $62M anually in grinding media costs as shown in Table 1. There will be the need for an additional magnetic separator step in the flowsheet but the finisher screens can be eliminated and the hydroseparator requirement will be smaller. The IsaMill has an internal centripetal classifier so there is no need for classifying cyclones in tertiary grinding. Wherever possible, gravity will be used to avoid the installation of extra pumps. For example, the ball mill cyclone overflow will gravity flow to the intermediate magnetic separator distributor and the magnetic separators themselves will be positioned so that the magnetics gravitate directly to the IsaMill feed tanks.



Table 1. Installed power and annual media cost comparison

Section Feed Rate (t/h)

Specific Energy (kWh/t)

Installed Power (MW)

Annual Media Cost Estimate

Autogenous Mill 770 µm Product

3800 8.5 40 $0

Single Stage Ball Mill 34 µm Product

2200 47 114 $86M

Single Stage IsaMill 34 µm Product

2200 34 78 $57M

Ball Mill 100 µm Product

IsaMill 34 µm Product

2200

1720

12

13

34

24

58

$13M

$11M

$24M

While 58 MW of ball mill and IsaMill power will be installed, only about 50 MW of operating power will be necessary under average operating conditions. The installation of eight 3.0 MW M10 000 IsaMills (four per line) also leaves some room for expansion in the future. Although each mill is powered by a 3 MW motor only 2.7-2.8 MW will be necessary per mill on average. Also, in this case, when one IsaMill is shut down for maintenance, the other mills will be able to be ramped up to full power without affecting production.

The addition of the hydroseparator to the circuit is critical to achieve a silica content of less than 2%. The magnetic finishers are incapable of fully removing the slimes present and entrained at the final grind size. The hydroseparator has been shown to decrease the tested finisher magnetic separator concentrate silica content by about 1%.

After the hydroseparator and finisher magnetic separation there is a sulfide flotation step. While pyrite is removed to tails in the magnetic separation steps the pyrrhotite present in the ore is magnetic and would otherwise report to the final concentrate. The sulfur content of the final magnetic concentrate is 0.6% and this is unacceptable for pelletisation plants. Through the sulfide flotation step the sulfur content of the final product is reduced to less than 0.1% which is acceptable.

The final concentrator flowsheet developed for this project thus becomes:

Two 20 MW AG mills closed off with 3mm screens followed by magnetic separation. The concentrate of this feeds two 17 MW ball mills closed off by hydrocyclones, from which the overflow feeds a second set of magnetic separators. This magnetic concentrate feeds eight 3 MW M10 000 IsaMills. The IsaMill product is at the final grind size without further classification and feeds in succesion to the hydroseparators, followed by the final magnetic separation step and finally the sulfide float. The underflow of the sulfide float step is pumped to a filter plant on the coast. Each separation step of magnetic separation, the hydroseparators and flotation will produce a stream of final tailings, with no recirculating loads planned.

CONCLUSIONS This work can be considered a success in that significant improvements have been made over the preliminary design in both capital and operating costs. Further, the stringent requirements for silica and sulphur levels in the final concentrate have been achieved even with these cost savings. This is made all the more impressive by the short time frame in which the work and design were completed. The pilot plant campaign on the AG mill discharge lasted just over one week. This included running two different ball mills and adding the M4 IsaMill to run continuously in the pilot plant. The savings of +56 MW of grinding power and over $60M annually in grinding media speak for the effectiveness of having a simple, yet well thought out plan of testwork and adapting that plan as results became available.



ACKNOWLEDGEMENTS The authors wish to thank the staff at AMMTEC for their work in this pilot plant campaign.

REFERENCES Burford B, Niva E, 2008. Comparing Energy Efficiency in Grinding Mills, in Metallurgical Plant Design and Operating Strategies (MetPlant 2008,) pp 45-64, (AUSIMM, Melbourne).

Devaney F, 1985. Section 20, Iron Ore, in SME Mineral Processing Handbook (ed: NL Weiss), Society of Mining Engineers of AIME.

Iwasaki I, 1983. Iron Ore Flotation, Theory and Practice, Gaudin Lecture, 1982 Annual AIME Meeting, AIME Transactions Volume 274 pp 622-631.

Iwasaki I, 1999. Iron Ore Flotation-Historical Perspective and Future Prospects, Nonsulfide Minerals, Section 3: in Advances in Flotation Technology pp 231-243.

Larson M, 2011. Xstrata Technology Internal Iron Ore Report.

McIvor R, Greenwood B, 1996. Pebble Use and Treatment at Cleveland-Cliffs’ Autogenous Milling Operations, in SAG 1996, pp 1129-1141 (Mining and Mineral Processing Engineering, UBC, Vancouver).



Contract Commissioning and Operation of Western Areas’

Cosmic Boy Concentrator

C Dick1,2, D Boska3 and C Fitzmaurice4

1. MAusIMM, Senior Metallurgist, Independent Metallurgical Operations Pty Ltd, West Perth WA 6005. Email: [email protected]

2. Formerly: Metallurgist, Cosmic Boy Concentrator.

3. Project Development Manager, Independent Metallurgical Operations Pty Ltd, West Perth WA 6005. Email: [email protected]

4. MAusIMM, Metallurgy Manager, Western Areas NL, West Perth WA 6005. Email: [email protected]

ABSTRACT

Over the past three decades, mining operations have seen a significant increase in the use of

contractors for mining, camp management and catering, recruitment, maintenance, power

management, and commissioning. More recently contractors have been engaged for stores and

logistics, OH&S, environmental, exploration, training etc. In some countries, such as South Africa, the

trend is towards total outsourcing of all aspects of a mining operation which is coordinated by the

owners’ management team.

During the Global Financial Crisis (GFC) more Australian companies considered and/or took on this

approach to recruiting for both projects and operations as mining companies were forced to rethink

their corporate strategies as commodity prices plummeted to low levels, pushing operations and

projects into potentially financially unviable territory.

This paper discusses the labour recruitment, commissioning and operating strategy of Western Areas

NL (WSA), with the assistance of Independent Metallurgical Operations Pty Ltd (IMO), as it

progressed construction, commissioning and operation of the Cosmic Boy nickel concentrator at the

Forrestania Nickel Project during and post GFC. The use of IMO contractors allowed Western Areas

to mitigate labour cost risk during the uncertainty caused by the GFC and let their management team

focus on the task of bringing a new operation into viable and profitable production.

BACKGROUND OF WESTERN AREAS COSMIC BOY

CONCENTRATOR

Western Areas, a growing mid-sized nickel sulphide mining company finalised the acquisition of the

Forrestania Nickel Project in 2003. This process started in 2002 with the purchase of Viceroy’s 75%

interest in the project, and was completed with the acquisition of Outokumpu’s remaining 25%

holding.



The Forrestania Nickel Project, located 400km south-east of Perth (Fig. 1 [Western Areas 2010a]),

became Western Areas’ primary asset. Amongst the suite of deposits in the project, Flying Fox would

become one of the highest grade nickel sulfide deposits in the world

Early in the lifecycle under Western Areas ownership it was identified that mineral resources at Flying

Fox totaled 2.0 million tonnes at an average grade of 4.6% Ni (2.5 Mt @ 5.5% Ni for 125 Kilo tonnes

of contained nickel including Lounge Lizard), containing approximately 94 050 tonnes of contained

metal (Western Areas 2010b).

Western Areas commenced mining nickel sulphide ore from the Flying Fox Underground Mine in

2005, producing 8000 tonnes of nickel in its first full year of production. The ore was sold to LionOre

(taken over by Norilsk Nickel in 2007) and processed at the Lake Johnston Nickel Operation

concentrator located approximately 80 km east of Forrestania. Further high grade nickel sulphide

areas (Diggers South, Spotted Quoll, and Cosmic Boy) at Forrestania (Fig. 2 [Western Areas 2007a])

indicated potential to further increase reserves.

The extent of the favorable geology provided Western Areas with the confidence required to pursue

the development of a concentrator at Forrestania. This would provide the company with the ability to

treat its own ore and importantly establish Forrestania as a major nickel province. The benefits

foreseen by Western Areas from constructing the Cosmic Boy concentrator included:

Ability to treat ore from a number of sources in a centrally located facility;

Ownership and management of the concentrate process;

Ability to improve metallurgical recoveries;

Ability to blend high and low grade ores from a number of deposits;

Potential to significantly increase profitability;

Move from Miner to a Concentrate Producer.

These benefits would enable Western Areas to progress towards the goal of becoming the third

largest nickel producer in Australia.



Fig. 1 ‐ Location of the Forrestania Nickel Project.

Fig. 2 ‐ Location of Forrestania Nickel deposits.

In June 2007 Western Areas successfully raised funds to proceed with a two stage plant construction,

and provided approval for the construction of the Stage One (250 000 tonnes per annum capacity

concentrator) in the third quarter of 2007. The construction phase of the Cosmic Boy Concentrator

Project commenced in 2008, with commissioning scheduled for first quarter of 2009.

Stage One of the project was designed to initially treat ore from the high grade Flying Fox nickel mine.

The proposed Stage Two plant was envisaged to double throughput capacity and allow treatment of

ore from Flying Fox and Spotted Quoll deposits.

The fast tracking of the project was made possible in part as a result of Western Areas acquiring key

plant components prior to the formal engagement of an Engineering Procurement Construction (EPC)

provider. This equipment included the crushing circuit, ball mill, conveyor system, flotation cells and

dewatering circuits.

Additionally the project benefitted from pre-existing major infrastructure at the Cosmic Boy site, such

as grid supplied power, a large Tailings Storage Facility (“TSF”), bore fields, water storage dams and

an ore bin.

On 20th February 2009 the Cosmic Boy nickel concentrator was commissioned, and officially opened

on 24th March 2009.



ISSUES & RISKS FACED BY THE COSMIC BOY CONCENTRATOR

PROJECT

Western Areas entered into the engineering stage of the Cosmic Boy Concentrator project in mid-

2007 as the GFC was starting to unfold overseas, particularly in the United States of America and

Europe. The effect of the GFC would ultimately have an effect on the Australian mining industry by

2008.

This extent and impact on the local mining industry was widespread and well documented.

Companies experienced significant reductions in revenue as commodity prices plummeted. Further

complicating the falling commodity prices were the elevated operating costs at many mining

operations. These were at levels remnant from the pre-existing mining boom and particularly reflected

in labour costs.

The widespread downturns in fortunes ultimately forced mining companies to scale back existing

mining and processing operations, mothball new projects or re-evaluate well advanced projects. Many

companies struggled to overcome cost pressures, resulting in mass workforce retrenchments and in

some circumstances cessation of operations.

Base metal producers did not escape lightly. Nickel prices plummeted from highs of $24 US/lb in April

2007 to almost $4 US/lb in late 2008, leading into 2009 (Fig. 3 [Infomine, 2011]), with a corresponding

increase in LME nickel stocks (Fig. 4 [Kitco Metals 2011]). Through this period many producers

cancelled projects or put operating processing plants into care and maintenance, including projects

such as BHP Billiton’s Ravensthorpe Nickel Laterite Operation and Norilsk Nickel Australia’s’ Cawse,

Black Swan and Lake Johnston Operations.



Fig. 3 – Nickel price trend over period of January 2005 to January 2010.

Fig. 4 – Five year LME Nickel Warehouse Stocks Level.

Under this economic climate, Western Areas identified labour as a key risk to the Cosmic Boy

Concentrator project. The company was not willing to recruit additional staff during this period of



uncertainty. Central to their concern were potential liabilities and overheads that would result from

abandonment or postponement of the Cosmic Boy Concentrator project from deteriorating economic

conditions and further depressed nickel prices.

To overcome this Western Areas employed the approach of using a contract workforce at the Cosmic

Boy Concentrator. This approach lessened the typical risks associated with recruiting permanent staff

during periods of market uncertainty, with additional uncertainty and risks coming from non-market

sources during the stages of plant construction, commissioning and production ramp up.

This enabled an external organisation to manage the Human Resources (HR) administration work as

Western Areas did not have a HR department, and preferred not to develop one. In turn this afforded

Western Areas the opportunity to continue the use of a relatively lean and flat management structure

at the Forrestania Nickel Project.

WESTERN AREAS RISK MITIGATION STRATEGY

To provide assistance to Western Areas in minimising the risk to the Cosmic Boy Concentrator Project

the use of a contract workforce was considered the suitable option.

Contractors have been used in an ever-increasing proportion, as individuals and contractor groups,

throughout the mining industry over the last three decades. The areas of work in which contractors

have been employed has also increased to include:

Mining production;

Mine and concentrator maintenance;

Logistics and stores management;

Catering;

Geological exploration;

Transport of ore and process products such as concentrate;

Training;

Occupational health and safety;

Environmental management.

In some countries, such as South Africa, there is an increasing trend towards total outsourcing of all

aspects of a mining operation. This outsourcing is typically coordinated by the owners’ management

team. The extent of this approach is demonstrated in its application at over 15 mineral processing

facilities in Southern Africa, which include platinum concentrator facilities, coal processing facilities,



and diamond processing plants. The advantages commonly touted are a method for new mines to

reduce up-front capital costs on equipment and engineering facilities, and operating mines to

minimise operating costs.

The mining industry in Australia has been aware of the advantages of using contract labour to

overcome the types of concerns that Western Areas were facing in starting their operation. This is

evidenced by the rise of contract labour usage across the entire mining industry in Australia over the

past 20 years (Fig. 5 [Government of WA, 2010a]).

Percentage of Contractors of Total Mining Mining Labour Force (Excluding Exploration & Petroleum)

0%

10%

20%

30%

40%

50%

60%

70%

1987

1989

1991

1993

1995

1997

1999

2001

2003

2005

2007

2009

Year

%C

on

tra

cto

rs in

Min

ing

Ind

us

try

Fig. 5 – Proportion of contractors of the total mining labour force (excluding exploration and

petroleum) between 1987 and 2009.

Contract and permanent labour workforces in the Western Australian nickel industry were significantly

affected by the GFC (Fig. 6 [Government of WA, 2010b]). The dramatic decrease in the total labour

force (permanent employees and contractors) over the period June 2008 to January 2009 put a

substantially large number of skilled personnel with nickel industry experience onto the market looking

for employment.



Mining Employment Nickel January 2003 to December 2009

0

2,500

5,000

7,500

10,000

12,500

15,000

17,500

Dec

-02

Mar

-03

Jun-

03

Sep

-03

Dec

-03

Mar

-04

Jun-

04

Sep

-04

Dec

-04

Mar

-05

Jun-

05

Sep

-05

Dec

-05

Mar

-06

Jun-

06

Sep

-06

Dec

-06

Mar

-07

Jun-

07

Sep

-07

Dec

-07

Mar

-08

Jun-

08

Sep

-08

Dec

-08

Mar

-09

Jun-

09

Sep

-09

Dec

-09

Employees Contractors TotalSource: DMP

Fig. 6 – Graph of permanent and contractor employment in the nickel industry between December

2002 and December 2009.

Prior to the Cosmic Boy project, Western Areas had used a mining contractor for the mine production

at the Flying Fox mine. This successfully demonstrated to Western Areas the benefits of a contracted

workforce. Other contractor groups used at Forrestania included camp administration and geological

exploration, similarly providing successful outcomes for Western Areas.

Supporting the risk factors, industry conditions, and experience with contractors, the decision to

pursue a contract labour force for Cosmic Boy was aligned with the development philosophy stated

below:

Building a low throughput, high grade concentrator (250 000 tonne per annum);

Using predominantly second hand equipment;

Construction completion of the concentrator scheduled for January 2009;

Start up of Cosmic Boy Concentrator on February 20th 2009;

Fast-track approach commissioning of concentrator to reach target recoveries of greater than

90% nickel recovery within 3 months of startup;

Using contract labour to commission and operate the concentrator and associated services.



Aspects relating to design and construction of the concentrator were planned prior to the GFC and

ensured minimal capital expenditure. Although the size and design throughput were relatively small

compared to other Australian nickel sulphide concentrators, the design would incorporate allowances

for future expansions.

Defined Feasibility Study metallurgical testwork conducted on the Flying Fox ore (4-6% Ni) had shown

that high nickel recoveries of greater than 90% were achievable at marketable concentrate grades of

14% Ni (Western Areas, 2004). The recovery and grade expectations were supported by production

data from toll treatment of early parcels of Flying Fox ore through Norilsk Nickel Australia’s Lake

Johnston concentrator (Western Areas, 2007b). This gave Western Areas the confidence to

commission and quickly ramp up a low throughput concentrator, demonstrating relatively low

metallurgical risks ahead of embarking upon further expansions over a short number of years to

increase nickel production.

The expansions were envisaged to be undertaken within 1-2 years of each upgrade stage. In making

allowances for the expansions at the design stage, Western Areas were offered a cost effective

pathway to take advantage of improving nickel prices alongside increased production, whilst

preempting a return to production from operations placed under care and maintenance.

The final part the strategy focused on the selection of a specialised contractor group. Western Areas

selected several groups to provide proposals for the manning of the concentrator. The main criteria

was the selection of a mature, experienced, and technically trained management group (metallurgists

and production coordinators), followed by the ability to provide process and laboratory technicians.

Western Areas’ preference was to deal with only one contracting group for these personnel, which

would make the workforce easier to manage given the limited resources within Western Areas.

The result of this process was the selection of IMO to provide the concentrator management team

and the technicians required. IMO would supply and maintain the majority of concentrator personnel

during commissioning and ramp-up of production at the Cosmic Boy Concentrator.

IMO approached the development of a labour workforce for Cosmic Boy using a simple organisational

chart (Fig. 7 [Western Areas 2011a]) requiring a minimum number of personnel on a cost plus basis

organised as follows:

Metallurgists (x2);

Production Co-ordinators (x2);

Shift Supervisors (x3);

Process Technicians (x8) – 3 x 2 shift process technicians, 2 x dayshift process technicians;

Laboratory Supervisors (x2) and Laboratory Technicians (x2);

Metallurgical Technicians (x2).

Supply/Logistics Officer (x1)



With IMO focused on the technical metallurgy and operational processing, the decision was made to

outsource certain maintenance personnel to other contractor groups. Some supply/logistics personnel

were sourced through other contractor groups, although later they would fall under IMO’s umbrella,

following Western Areas preferences to deal with a minimum number of contractor groups.

Mill Manager (WSA)

2 x Production Co-ordinator(IMO)

2 x Plant Metallurgist

(IMO)

Maintenance Superintendent

(WSA)

3 x Shift Supervisor (IMO)

2 x Lab Supervisor (IMO)

2 x Metallurgical Technician

(IMO) 2 x Lab Technician

(IMO)

8 x Process Techs(IMO)

Maintenance Planner/Supervisor

(IMO)

2 x Electrician(Reside)

2 x Mechanical Fitter

2 x Boilermaker(Interquip)

1 x Supply / Logistics Officer

(IMO)

Fig. 7 – Organisational chart for Cosmic Boy Concentrator.

To justify the low cost approach only three teams of process technicians were used. This approach

required the use of a rotational roster of 2 weeks on and 1 week off for these contractors.

Supporting this decision was the availability of processing personnel with nickel experience as a result

of the downturn. This allowed a reversal of the trend towards shorter rotational rosters noticed during

previous years and resulted in a shorter recruitment timeframe and reduced overall cost to Western

Areas.

In addition, the longer rotations facilitated commissioning and ramp up to full production by ensuring

extended continuity of personnel during this period and promoted more rapid training and plant

familiarity.

With a good knowledge of suitable personnel in the industry, WSA personnel assisted IMO in

targeting key personnel, allowing IMO to recruit a whole metallurgical team in a short period of time of

six to eight weeks and also provide leave coverage and replacements. The ability to respond rapidly

to personnel changes and requirements allowed the operation to successfully manage the transition

that typically occurs from commissioning to production. Personnel movements typically associated

with this period are reflective of the different skill sets and experience required for each of these

stages in the life of the plant. This flexibility and ability to match skills sets to the various stages of

plant startup may have been significantly reduced if permanent staff were deployed from the onset of

the project.



COMMISSIONING & FIRST 12 MONTHS OF OPERATION

Senior IMO contract personnel (Production Coordinators and Metallurgists) started arriving on site in

late January 2009, with the remainder of the concentrator workforce on site by mid February 2009.

The construction of the concentrator was completed in January 2009, within Western Areas allocated

time frame and budget by GR Engineering Services (GRES). The design of the 250 ktpa plant (Fig. 8

[Western Areas, 2011b]) incorporated the following processing stages and related equipment

(Western Areas, 2011c):

Single stage crushing (Minspec 750x1050 Single Toggle Jaw);

SAG Mill grinding (3.8 m Diam., rubber lined, 700 kW motor);

Flash flotation (SK-80 Outotec flotation cell);

Roughing and Scavenging Flotation (12 x 5.7 m3 Agitair flotation cells);

Single Stage Scavenger Cleaner Flotation (5 x 2.8 m3 Metquip flotation cells);

Concentrate Thickening (Supaflo High Rate 10 m Diam.);

Tailings Thickening (Supaflo High Rate 7 m Diam.);

Concentrate Filtration (Lasta 1500 x 1500 40/40 plate and frame filter).

Reagents: Potassium Amyl Xanthate (Collector), IF56 (Frother), Quicklime (pH modifier),

Magnafloc E10 (Flocculant), Guar (Depressant).

Onsite laboratory services were provided which consisted of an onsite analytical laboratory (sample

preparation facilities, wet chemistry lab - AAS) and a small metallurgical laboratory (rod mill, flotation

cell, sizing equipment, etc). In addition onsite maintenance workshop, stores facilities and office

buildings were constructed.



Fig. 8 – Cosmic Boy Concentrator flowsheet (summarised).

Dry commissioning of some equipment was conducted in late 2008 by GRES. These included the

crushing circuit and mill feed system as far as the mill feed chute (apron feeder and conveyor).

Water testing of the flotation and dewatering circuits occurred during the first two weeks of February

2009. Despite a narrow timeframe between processing personnel arriving onsite testing was able to

be conducted over both day and nightshifts by the end of the second week of February.

Photos of the main concentrator circuits, infrastructure and layout are shown in Fig. 9 (Western Areas,

2011d).



Fig. 9 – Cosmic Boy Concentrator photos clockwise from top left ‐ Grinding Circuit, Crushing circuit,

Rougher and Scavenger Flotation Cells, Site layout during construction).

During the first month on site IMO personnel worked closely with GRES and Western Areas

personnel to understand the flowsheet, associated equipment, review the construction and undertake



a gap analysis. All parties worked together to address safety concerns and improvements to

operability of plant and equipment.

On the target date of February 20th 2009 the first ore was treated through the milling circuit (Fig. 10

[Western Areas 2011e]). The concentrator was commissioned on relatively low grade material, <2%

Ni head grade (Western Areas, 2011f); to ensure that nickel losses were kept to a minimum over the

first few days of commissioning. During this time commissioning continued on the following

parameters (Western Areas, 2011g):

Dry and wet plant (crushing and milling circuits) throughput;

Dry and wet plant availability;

Grind product size;

Thickener overflow clarities;

Thickener underflow densities;

Filtration rate;

Filter cake moisture content.

Final handover to Western Areas and sign off occurred in the first week of March 2009.

Fig. 10 – First ore to the Cosmic Boy milling circuit.

Throughout the remainder of February and March, IMO and Western Areas personnel worked closely

with GRES construction / commissioning teams thus ensuring a smooth handover as higher grade ore

(~3.7% Ni) from Flying Fox mine was feed into the concentrator (Western Areas, 2009a).



GRES were very supportive of contract concentrator personnel throughout the startup and

commissioning through:

Providing Training and information transfer where possible;

Experienced commissioning engineers sharing the metallurgical testwork workload with

senior IMO personnel during startup;

Providing maintenance support beyond contractual requirements during startup, through

resources and equipment;

Building a solid relationship between GRES staff and concentrator personnel through the final

stages of the construction and throughout commissioning.

Metallurgical and Production Challenges during Commissioning

During commissioning it was demonstrated that target throughput could be achieved on more

competent and abrasive lower grade ore (<2% Ni) (Western Areas, 2011h). This throughput was only

achieved for short periods due to a build up of critical sizes, hard abrasive gangue material

(sedimentary rock and granite) being held up in the mill and resulted in high mill power loads,

restricting overall plant throughput and causing significant wear on mill lifters and liners.

This build up of competent gangue material in the mill severely hampered the ability to maintain

concentrator throughput at target tonnages and achieving grinding circuit stability for extended

periods. It also limited the ability to run other equipment at design and target capacities.

The source of this problem was identified to be a result of extremely high ore dilution by competent

gangue material. The level of this dilution significantly exceeded levels used in metallurgical

comminution testwork.

To effectively manage this problem, site geologists and mining engineers worked closely with senior

concentrator personnel to manage dilution ratios of ROM stockpiles in order to anticipate grinding

media requirements and throughput capacities for planned ore blends.

The underestimation of waste rock dilution and its effect on mill lifter/liner wear was significant,

effectively doubling the frequency of mill relines. This problem was tackled on several fronts:

More frequent inspections of mill liner wear (at least every 2-3 weeks);

Close consultation with the mill liner supplier for lifter/liner change out scheduling;

Installing pebble ports into the mill discharge grates to allow critical sized gangue material to

exit the mill at an acceptable rate;

Changing from rubber to Polymet lifters in October 2009 to decrease mill reline frequency.



These measures, implemented in a staged approach, allowed mill throughput to be stabilised with

minimal mill feed variation, and gradually throughput increased over time as grinding stability

improved.

Metallurgical performance of the Cosmic Boy Concentrator was improved methodically and quickly

over the first 6 months of operation. Some key aspects and observations were:

Concentrator personnel identified problem areas and opportunities, working closely with

metallurgical and laboratory technicians to conduct metallurgical testwork (grinding and

flotation plant surveys, plant pulp chemistry surveys, diagnostic flotation tests of flotation feed

and flotation tailings) to understand the metallurgical issues and develop solutions. During this

period an additional metallurgist and metallurgical technician were recruited to help drive and

complete the required metallurgical testwork programs;

Senior concentrator personnel (WSA and IMO) worked closely with Western Areas geologists

to understand the sulphide and gangue mineralogy of the Flying Fox ore and developed

methods of improving flotation selectivity of the nickel sulphides;

A strong relationship was built between concentrator processing and maintenance personnel

to target problematic equipment whilst ensuring a well structured preventative maintenance

programs was implemented;

Senior concentrator personnel worked closely with the grinding mill liner supplier to manage

the difficulties arising from the accelerated wear on mill liners, which on occasion over the first

6 months necessitated mill relines at short notice. Feed, shell and discharge liners were also

being changed out at different frequency rates which generally did not coincide with change

outs for other liner sections of the mill, hindering overall operational performance and

availability.

One of the most significant metallurgical problems to overcome occurred in the third quarter of 2009.

In this period an increasing proportion of the ore being mined at the Flying Fox mine was coming from

the T4 section of mine relative to that from the T2 section (Fig. 11 [Western Areas, 2007c]). The T4

section of the mine had previously been removed from the mining schedule due to concerns over

ability to mine and treat the ore to make a marketable concentrate. However it was added to the

mining schedule to supply more ore to the concentrator after metallurgical testwork demonstrated it

was treatable (Western Areas, 2011i).

The flotation performance of the concentrator decreased as the proportion of the T4 ore in the mill

feed blends was increased. Consultation with site geologists found there were significant differences

between the mineralogy of the T2 ore and the T4 ore. Understanding mineralogical variation

throughout the Flying Fox mine, especially in iron sulphide alteration zones (Pyrrhotite to Pyrite ratio),



was found to be critical in the understanding of required concentrator operating strategies to ensure

iron sulphide variation (i.e. proportion of Pyrite in ore) had minimal effect on concentrate grade.

Significant amounts of testwork and research were conducted by the metallurgical team to understand

the Pyrite flotation characteristics and recovery distribution in the concentrator and development of

methods to improve flotation selectivity of nickel sulphides against the variable amounts of Pyrite in

stockpile blends.



Fig. 11 – Diagram showing Flying Fox mine sections (Interpreted long section 2007).



By early June 2009 as a result of a strong focus on metallurgical testwork and group effort the

recovery target of 90% nickel recovery at a design concentrate grade of 14%Ni was achieved (Fig. 12

[Western Areas, 2009b]). Ongoing plant optimisation work during the quarter further improved

metallurgical recoveries to 92.6% that month.

In total, by the end of June 90 300 tonnes of ore at 4.2% Ni had been milled since commissioning

resulting in the production of 22 946 tonnes of concentrate grading 14.5% Ni (Western Areas, 2009c).

YTD Concentrator Monthly Performance

0.0

0.5

1.0

1.5

2.0

2.5

3.0

3.5

4.0

4.5

5.0

Feb‐09 Mar‐09 Apr‐09 May‐09 Jun‐09

Feed Grade (%Ni)

84

85

86

87

88

89

90

91

92

93

94

Nickel R

ecovery (%)

Feed nickel grade (calc) Nickel Recovery (calc)

Fig. 12 – Metallurgical production data for first 6 months of operation.

Despite the relatively short commissioning period target throughputs and nickel recoveries were

consistently achieved within 5 months, resulting in the metallurgical performance over 2009

exceeding the Western Areas expectations (Fig. 13 [Western Areas, 2010c]).

By the end of December 2009 the concentrator had milled 215 356 tonnes at 3.8% Ni head grade and

produced 49 641 tonnes of concentrate averaging 14.5% Ni (7180 Ni tonnes) (Western Areas,

2009d).



Concentrator Monthly Performance

0.0

5.0

10.0

15.0

20.0

25.0

30.0

35.0

Feb‐09 Mar‐09 Apr‐09 May‐09 Jun‐09 Jul‐09 Aug‐09 Sep‐09 Oct‐09 Nov‐09 Dec‐09

Feed nickel grade (calc) Mill Rate (tph) Con % Ni Design Mill Rate (tph)

Fig. 13 – Cosmic Boy Concentrator metallurgical production results for February to December 2009.

During the plant commissioning, the project encountered a parallel hurdle with respect to concentrate

sales. Western Areas sole off-take agreement designated concentrate to be transported to Norilsk

Nickel Australia’s Lake Johnston Operation. The cessation of operations at Lake Johnston forced

Western Areas to find an alternative customer and negotiations were initiated with BHP Billiton (Nickel

West).

During this period nickel concentrate had to be stockpiled outside Cosmic Boys concentrate shed

which had rapidly reached its capacity of approximately 6000 tonnes. It was only in May 2009, three

months after the commencement of commissioning, that the first truckload of Cosmic Boy nickel

concentrate was transported to BHP Billiton’s Kambalda Nickel Concentrator and Kalgoorlie Nickel

Smelter.

The management of the concentrate stockpiling and truck loading was challenging due to the limited

ground area available and narrow traffic corridors for the road trains. Successful management of this

situation was a credit to the coordinated efforts of the multiple contractor groups involved.

Fig. 14 shows a road train being loaded from the concentrate stockpiled outside of the Cosmic Boy

Concentrate shed (Western Areas 2011j).



Fig. 14 – Loading outside stockpiled nickel concentrate at Cosmic Boy in 2009.

Interestingly and advantageously for Western Areas, the closing of the Lake Johnston operation made

available processing personnel who had experience in treating Flying Fox ore, albeit via a different

flowsheet. The flexibility provided by using a contract workforce allowed targeting and recruitment of

some of these personnel, effectively allowing the Cosmic Boy workforce to rapidly adapt and instantly

up skill to meet evolving operational requirements borne through Flying Fox ore. In addition to the

technical skills gained the project was afforded additional skills in the areas of maintenance, supply

and logistics and important local awareness. To achieve a similar outcome with permanent staff

already in place would have been much more difficult.

Safety Performance during Commissioning

With only two MTI's and no LTI's over the first 12 month period of operation (Figs. 15 [Western Areas,

2011k] and 16 [Government of WA, 2011c]), the safety performance of Cosmic Boy concentrator

personnel was extremely good and well below the overall statistics for the industry in general.



Fig. 15 – Graph of major safety statistics from Cosmic Boy Concentrator for first 12 months of

operation.

Fig. 16 – Safety statistics (LTIs) for the nickel sector for period 2004‐05 to 2008‐09.



Evaluation of Contractor Personnel during Commissioning

During commissioning in the majority of concentrators it is likely, and accepted, there will be a degree

of employee turnover. In addition, there is a widely held view within the mining industry that employee

turnover is higher with contractors than for permanent employees. Despite this perception the

turnover of IMO contract personnel at Cosmic Boy during the period discussed was typical of the

mining industry permanent employees and in the range of 15 to 30%.

Western Areas initiated recruitment through IMO quite late in project lifecycle, starting in December

2009, with the aim of minimising labour costs during the last stages of construction. Despite keeping

labour cost low, the limited recruiting time had some negative consequences. One such consequence

was that some contract personnel were employed despite not having sufficient relevant experience.

This was demonstrable whereby some process technicians did not have suitable base metal grinding

and flotation experience, which was later found to be critical, especially in making quick adjustments

to variations in ore mineralogy in new mill feed blends.

The limited recruiting time also allowed very little time for processing personnel to understand the

concentrator flowsheet, learn the metallurgical process from the GRES commissioning metallurgists,

and develop safe working procedures (SWPs) and job safety analysis (JSAs) before startup. In some

cases, insufficient training and problems may have been prevented had sufficient time been allowed

for these items to be completed. This also applied to the laboratory and maintenance personnel.

After the first few months of commissioning it was realised that there were some areas of work, which

could be conducted by the personnel on site and additional personnel needed to be sourced to

conduct these areas of work. The roles were:

1 x Process Technician on each crew - the 2 Dayshift process technicians were converted to

shift process technicians and 1 more process technician was recruited;

1 x Project Metallurgist - to conduct required metallurgical testwork to improve concentrate

grade and nickel recovery, especially as mill feed transitioned T2 ore to T4 ore from the Flying

Fox mine;

1 x Project Metallurgical Technician – assist the Project Metallurgist conducting the additional

metallurgical testwork during the commissioning period and beyond;

Contract Training Coordinator – to assist in writing SWPs and JSAs for all departments and

areas of the concentrator and help train process technicians;

1 x Maintenance Supervisor – to assist the Maintenance Superintendent in managing

maintenance teams, shutdown contractors, and help improve maintenance systems and

planning.

1 x Supply/Logistics Officer – to ensure minimum spare parts were maintained, provide

regular contact with suppliers, and expedite urgent spares when required.



SUMMARY AND CONCLUSIONS

Western Areas approached the Cosmic Boy Concentrator during the boom period prior to the onset of

the GFC with a low cost view point, in respect of capital and operating expenditure. This placed the

company in a position where it could continue to progress the project once the full effect of the GFC

was realised with the nickel price plummeting.

Western Areas used a project management strategy that would provide flexibility and minimize risk, in

addition to reducing (and in some cases transferring) liability, should the economic climate worsen. A

key aspect of this strategy was the use of a contract labour force to commission and operate the

concentrator.

There were numerous challenges experienced during the startup and commissioning of the Cosmic

Boy Concentrator, which arose over the first twelve months of operation. These challenges were

overcome by the combined efforts of all site-based groups, most of which were contractor groups.

Overall the experience in commissioning the Cosmic Boy Concentrator in 2009 showed that it is

possible to progress capital intensive projects through market down turns, even ones as significant as

the GFC, with the use of sound project planning and evaluation to minimise exposure to high

operational costs, especially labour costs.

ACKNOWLEDGEMENTS

The authors would like to acknowledge the management of Western Areas and IMO for allowing this

paper to be written and providing production and information from the first year of operation of the

Cosmic Boy Concentrator.

The authors would also like to acknowledge the efforts of all Western Areas staff, and IMO and

Interquip contractors who contributed to the successful commissioning of the Cosmic Boy

Concentrator in 2009, especially that of John Bower, Devon McNeill, Liz Brown, and Pauline Vass. If

not for their experience, knowledge and efforts this paper may not have been written.

REFERENCES

Government of WA, 2010a. Department of Mines & Petroleum (DMP). East Perth, Western Australia.

DMP AXTAT reporting system. Available from: <http://www.dmp.wa.gov.au/8481_1476.aspx>

[Accessed 1st March 2011].

Government of WA, 2010b. Department of Mines & Petroleum (DMP). East Perth, Western Australia.

AXTAT reporting system. Available from: <http://www.dmp.wa.gov.au/8481_1476.aspx> [Accessed

1st March 2011].



Government of WA, 2010c. Department of Mines & Petroleum (DMP) (Resources Safety). East Perth,

Western Australia. Safety Performance in the Western Australian Mineral Industry – Accident and

Injury Statistics 2008-2009. Figures 22 & 23, page 26.

Infomine Inc, 2011. Infomine, Mining Intelligence & Technology. Vancouver, British Columbia,

Canada. Available from:

<http://www.infomine.com/chartsanddata/chartbuilder.aspx?z=f&g=127670&dr=5y>. [Accessed 10th

February 2011].

Kitco Metals, 2011. Montreal, Quebec, Canada. Available from:

<http://www.kitcometals.com/charts/nickel_historical_large.html>. [Accessed 28th February 2011].

Western Areas 2004. Flying Fox Metallurgical Testwork Summary for Western Areas NL. 2004.

Internal company document.

Western Areas 2007a. Western Areas 2007 AGM Meeting Presentation, 30th November.

Western Areas 2007b. First Production 31 December 2006 Half Year Report” Western Areas News

Release, 26th February 2007.

Western Areas 2007c. Western Areas 2007 March Quarterly Report, Figure 1. Long Projection of

Flying Fox mine. (Drill hole intersections are down hole widths).

Western Areas 2009a. Western Areas March 2009 Quarterly Report.

Western Areas 2009b. Western Areas June 2009 Quarterly Report. Figure 1: Monthly metallurgical

recoveries from commencement of production at Cosmic Boy Concentrator. September 2009.

Western Areas 2009c. Western Areas March 2009 & June 2009 Quarterly Reports.

Western Areas 2009d. Western Areas March 2009, June 2009, September 2009 and December 2009

Quarterly Reports.

Western Areas 2010a. Western Areas 2010 AGM Meeting Presentation, 8th October.

Western Areas 2010b. Western Areas NL website. Available from:

<http://www.westernareas.com.au/phoenix.zhtml?c=169423&p=irol-index> [Accessed 30th January

2011].

Western Areas 2010c. Internal Western Areas document. Personal communication. February 2011.

Western Areas 2011a. Internal Western Areas document. Personal communication. February 2011.

Western Areas 2011b. Internal Western Areas document. Personal communication. February 2011.

Western Areas 2011c. Internal Western Areas document. Personal communication. February 2011.

Western Areas 2011d. Internal Western Areas document. Personal communication. February 2011.

Western Areas 2011e. Internal Western Areas document. Personal communication. February 2011.

Western Areas 2011f. Internal Western Areas document. Personal communication. February 2011.



Western Areas 2011g. Internal Western Areas document. Personal communication. February 2011.

Western Areas 2011h. Internal Western Areas document. Personal communication. February 2011.

Western Areas 2011i. Internal Western Areas document. Personal communication. February 2011.

Western Areas 2011j. Internal Western Areas document. Personal communication. February 2011.

Western Areas 2011k. Internal Western Areas document. Personal communication. February 2011.



Float it, Clean it, Depress it – Consolidating the Separation Stages at Clarabelle Mill

V Lawson1 and M Xu2

1. MAusIMM, Manager, Mineral Separation Technology Canada, Vale, North Atlantic Base Metals, 18 Rink Street, Copper Cliff P0M 1N0, Canada. Email: [email protected]

2. Section Head Mineral Processing, Vale, Base Metals Technology Development, 2060 Flavelle Boulevard, Mississauga, ON, L5K 1Z9, Canada. Email: [email protected]

ABSTRACT

Vale’s Sudbury Clarabelle Mill processes Cu/Ni ores from a number of mines in the Sudbury basin in Ontario, Canada. Bulk sulphide flotation produces a copper-nickel concentrate with pyrrhotite as the major diluent. After successful commissioning of the Voisey’s Bay Cu/Ni Mill (VBN) in Labrador, a miniplant campaign on Clarabelle ore with the VBN flow sheet produced a significant recovery improvement over current performance.

In 2001, the Mill Redesign Project moved grinding energy into rock/sulphide and pentlandite/pyrrhotite separation and in 2006 Cu/Ni separation was re-introduced to Clarabelle Mill to produce a saleable Cu concentrate. The projects were a great success. What remained was to simplify the Clarabelle flow sheet by introducing the missing components found in the VBN flow sheet. This project known as CORe has evolved through many stages including the mining boom and surviving the 2008 economic downturn and it is scheduled for completion in mid 2013. This paper outlines some of the challenges faced in this major brownfields project – organizational, technical, financial and operational.

INTRODUCTION

Vale’s Clarabelle Mill processes copper-nickel ores from mines in the Sudbury basin in Ontario, Canada. The ore contains economically significant quantities of pentlandite (Pn – (NiFe)9S8), chalcopyrite (Cp – CuFeS2) and pyrrhotite (Po – Fe8S9). The ore also contains quantities of precious metals notably Pt, Pd and Au. A typical Clarabelle Mill ore feed grade is given in Table 1. Clarabelle Mill produces a copper concentrate with <0.4% Ni for sale and a Ni concentrate at a grade of between 10 and 12% Ni for the smelter.

Table 1: Typical Clarabelle Mill Feed

Base Metal Assay (%) Precious metal (ppm) Mineral Assay (%)

Element Cu Ni Fe S Pt Pd Au Cp Pn Po Rock Assay 1.4 1.2 19.4 9.1 0.6 0.6 0.2 4.0 3.0 16.9 76.0

In Sudbury, the pentlandite occurs in association with and is less abundant than the pyrrhotite. The pyrrhotite in the ore occurs in two phases namely monoclinic and hexagonal. The monoclinic pyrrhotite in the ore is ferromagnetic whilst the hexagonal pyrrhotite is paramagnetic. Rejection of pyrrhotite while minimising pentlandite, chalcopyrite and precious metal losses has been the subject of development work since the 1970’s (Marticorena et al., 1994, Kerr et al., 2003, Lawson et al., 2005). The abundance of monoclinic pyrrhotite was notable in the development of the current Clarabelle Mill flow sheet. Investigations of ore characteristics for the five year mine plan identified that the proportion of hexagonal pyrrhotite in the mill feed is expected to increase by as much as



30% and the Cu:Ni ratio was expected to increase by 30%. Changes in the flow sheet would be required to address some of the long range mineralogical challenges.

CLARABELLE FLOWSHEET

The first stage of separation is targeting sulphide/rock separation. Twenty percent of the total ore mass is recovered to a magnetic concentrate followed by rougher scavenger flotation from which three concentrates, A, B and CD, are produced. The A concentrate is sent uncleaned to the copper nickel separation circuit and the B and CD concentrates are cleaned to reject rock and pyrrhotite in separate cleaning stages. Pyrrhotite is intentionally recovered to the CD concentrate to produce a rock tails low in sulphur. Regrinding currently only occurs in the magnetic circuit where the liberation of pentlandite is increased from 65% to 85% to reject pyrrhotite. The Clarabelle Mill current circuit flow sheet is given in Fig. 1.

Fig. 1: Clarabelle Mill current flowsheet

The flowsheet at Clarabelle Mill takes account of the predominant mineralogy as shown in Table 2. At a separation circuit feed size of 150 micron P80 all minerals have adequate liberation for primary separation. The other interesting feature is that the separation of pentlandite from pyrrhotite and the separation of rock from chalcopyrite based on the liberation data support the need for regrinding. However, the separation of pentlandite from chalcopyrite should be achieved without regrinding as the associations are not abundant.



Table 2: Liberation data for Clarabelle Mill primary grinding circuit product (as feed to separation circuit at P80 of 150 μm)

Size Range Mineral Liberation-2 Dimensions

Mineral Status Pn Cp Po Rk

Liberated 69 75 85 96

Binary - Pn 2 5 0

Binary - Cp 2 1 2

Binary - Po 22 3 1

Binary - Rk 1 13 5

Multiphase 6 6 5 1

Total 100 100 100 100

Flowsheet Development 1998 - 2002 (The road to change…)

Up until 1998, the Clarabelle Mill produced three concentrates and three tailings as seen in Fig. 2. The rougher concentrates were sent uncleaned to the bulk concentrate in addiiton to two minor contributions from both a magnetic and non-magnetic pyrrhotite rejection circuit. A major advance in pyrrhotite rejection was the introduction of diethylenetriamine (DETA) used in conjunction with sodium sulphite (SS). The DETA/SS combination clearly improved pyrrhotite rejection and its effect has been discussed in detail by Marticorena et al., (1994) and Xu et al., (1997) and will not be discussed here as its efficiacy is not in doubt.

Fig. 2: Clarabelle Mill flowsheet 1994 - 1998

The first significant flowsheet change became necessary when a new orebody was brought on-line and there was deterioration in metallurgical performance. The new orebody was characterised by high copper and nickel grades but the pyrrhotite was almost entirely hexagonal and thus not recovered to the magnetic concentrate. The abundance of pyrrhotite reporting to the non-magnetic circuit resulted in lower concentrate grades. To improve the rejection of hexagonal pyrrhotite, the non-magnetic pyrrhotite circuit was reconfigured to a two stage circuit with regrind as shown in Fig. 3 (1999-2001). Laboratory



and Pilot plant testwork demonstrated the advantages of a regrind and a two stage cleaning circuit with DETA/SS resulted in rock rejection followed by pyrrhotite rejection. The testwork clearly showed that DETA/SS on its own actually increased rock recovery and it was postulated that the increased froth stability incurred with DETA increased entrained rock and the extra flotation stage effectively cleaned the rock from the concentrate.

Fig. 3: Clarabelle Mill flowsheet after scavenger recleaning project

The two stage scavenger cleaner circuit was a success and gave the metallurgical team at Clarabelle and at the research facility in Mississauga the confidence of the management team for the next project improvement. This next opportunity was identified and progressed to a new capital project known as the Mill Redesign Project (Kerr et al., 2003). The Mill Redesign team carried on with the successful work and reconfigured mill grinding power from primary grinding to regrinding as the ore demand reduced. The resultant changes improved pyrrhotite rejection by increasing the regrinding power from 2250 kW to 3000 kW in the magnetic circuit and by applying 750 kW of regrinding power to the B rougher concentrate with a new cleaning stage. These changes addressed chalcopyrite rock composites in the B rougher stage and helped depress hexagonal pyrrhotite that was increasing in predominance as the Vale mines were deepening. The mill flowsheet is shown in Fig. 4 (2001-2006) for this period of time.



Fig. 4: Clarabelle Mill flowsheet post Mill Redesign project

Mill Redesign was a very successful project and the learning’s from research and pilot plant operation were being directly applied in successful projects addressing current metallurgical issues. In 2002, attention shifted to the study of Voisey’s Bay Nickel (VBN) to ensure the successful design and construction of mine and concentrator in Newfoundland and Labrador.

The impact on bulk concentrate quality is shown in the Fig. 5 below indicating that even a small proportion of hexagonal pyrrhotite in feed results in significant hexagonal pyrrhotite in concentrate. Isolating the monoclinic pyrrhotite in its own circuit using magnetic separation becomes less relevant as the proportion of hexagonal pyrrhotite in the feed increases. The use of DETA/SS in three independent pyrrhotite circuits reduces the cost effectiveness. Simplifying the pyrrhotite circuit was a high priority and the strategy was used effectively in the VBN flowsheet development.

VOISEY’S BAY PROJECT DEVELOPMENT

Pilot plant and laboratory studies were made through the 1990’s and continuing through to 2003 to develop the Voisey’s Bay deposit. A final pilot plant operation at SGS Lakefield in Ontario, Canada in collaboration with Vale Base Metals Technology Development personnel

Fig. 5: Comparison of Feed and Bulk concentrate for Clarabelle Mill



was performed in 2002 and 2003 that resulted in the final concentrator flowsheet. Significant learnings from operations at Clarabelle Mill were investigated as the VBN flowsheet evolved into the constructed flowsheet shown below in Fig. 6.

The main differences in feed properties and flow sheets between VBN and Clarabelle can be summarized as follows:

Feed grades are significantly higher (years 1-5 2.2% Cu and 3.2% Ni and 20% S).

Significantly less rock as the deposit is generally massive sulphide (>50% sulphides).

More association of Cp and Po and existence of pentlandite flames in chalcopyrite.

Abundance of troilite (stoichiometric non-magnetic FeS) occurs in association with pyrrhotite.

There is an abundance of hexagonal pyrrhotite.

The liberation of the valuable minerals of the VBN ore at a grinding P80 of 100 µm was generally higher than Clarabelle although also more complex intergrowths of Cp/Pn and Pn/Po exist. These intergrowths supported the inclusion of regrind mills in the copper-nickel separation circuit as well as in the scavenger cleaning circuit. The liberation characteristics are shown in Table 3.

Fig. 6: VBN constructed flowsheet



Table 3: Liberation characteristics of Voisey's Bay Ore at P80 of 100 μm

Size Range

Mineral Liberation-2

Dimensions

Mineral Status Pn Cp Po

Liberated 85 78 93

Binary - Pn 3 2

Binary - Cp 2 2

Binary - Po 8 12

Binary - Gn 3 4 3

Multiphase 3 4 1

Total 100 100 100

The VBN commissioning in 2005 and operation since has been very successful. The metallurgy has met or exceeded expectations. The regrind mills have not been required to date resulting in a significant power savings.

Copper Nickel Separation at Clarabelle (2005 – 2006)

Until construction of the hydrometallurgical facility at Long Harbour in Newfoundland is complete, the concentrates from VBN are transported to Thompson and Sudbury. As the mill and smelter in Sudbury were matched in terms of production of concentrate, a strategy was required to make room for VBN concentrate in the Sudbury Smelter. This strategy came as the reintroduction of copper-nickel separation into the milling flowsheet at Clarabelle Mill. Copper-nickel separation is still beneficial on the matte from the smelter in the matte processing plant but the removal of some copper at the mill enables the matte processing plant to operate at a lower copper:nickel ratio (Cu:Ni). A plan was developed to remove 150,000 tonnes of copper concentrate to make room in the smelter for the VBN nickel concentrate.

Copper/nickel separation had been conducted at Copper Cliff Mill on Clarabelle’s Bulk concentrate from 1972 to 1991 before bulk smelting was introduced. As a great deal of work had been conducted on copper-nickel separation in the milling flow sheet so the technical risks were considered low but for equipment sizing and orientation a miniplant operation was required. Miniplant investigations conducted at Clarabelle Mill from May to November 2005 confirmed a simple flow sheet which by using some of the key elements of the VBN and historical Copper Cliff flowsheets, could produce a copper concentrate grading 31.5% copper with less than 0.4% nickel at a copper recovery of 50% from ore (Xu et al., 2004 and 2005). As the Cu:Ni ratio is a key parameter in effective copper nickel separation in matte processing complete copper removal at Clarabelle Mill was not the objective.

The copper separation circuit was commissioned in October 2006 (Fig. 7, 2006-2008). Success was immediate and now in 2011 the circuit is operating at a rate of 225,000 tonnes of copper concentrate or 150% of original design as the copper head grades have increased. Plans are currently underway to remove more concentrate by open circuiting the copper scavengers and conversion of an existing Jameson cell to copper duty. Currently column flotation is the circuit bottleneck.



The Clarabelle flowsheet as shown in Fig. 7 (post copper separation) had now evolved into a complex circuit with many stages each serving a critical purpose but becoming more complex to operate. The five key separation steps were;

Magnetic circuit – for pentlandite/magnetic pyrrhotite separation,

Rougher flotation – for selective sulphide/rock separation,

Scavenger flotation – for sulphide/rock separation,

B rougher cleaning and scavenger cleaning - for chalcopyrite/rock and pentlandite/non magnetic pyrrhotite separation, and

Copper nickel separation.

CLARABELLE FLOWSHEET DEVELOPMENT VIA MINIPLANT

In 2007 the VITSL miniplant was again brought to Clarabelle Mill to investigate the impact of a simplified flowsheet on nickel metallurgy. The flowsheet was essentially the VBN flowsheet that had been operating successfully for two years. Previous work by Xu, et al. (2000 and 2003), using the miniplant at Vale Base Metals Technology Development had shown that the flowsheet could produce a higher concentrate grade recovery curve. The confidence in this result was buffered by the knowledge that the miniplant had not been able to reproduce Clarabelle Mill metallurgy using the Clarabelle flowsheet. It was expected that some of this concern was from trying to reproduce the pulp chemistry produced in the magnetic and non-magnetic pyrrhotite circuits in Mississauga. The intention was to move the miniplant to Clarabelle Mill so that the Mill process water and chemistry could be matched more closely and a better understanding of mill feed variability could be undertaken to give the project team the confidence that was required to get project approval. Previous work on the differences between Clarabelle Mill and the miniplant has been published and will not be discussed further (Lawson et al 2005).

Fig. 7: Clarabelle Mill flowsheet post copper separation 2006



The miniplant campaigns lasted from January to June 2007 and totalled in excess of 85 tests including a campaign to demonstrate that the miniplant could match Clarabelle Mill performance with the same flowsheet something that had been missing in campaigns in Mississauga. In Fig. 8, the miniplant performance is compared to the mill model and it shows a consistent improvement in recovery across a large range of operating conditions and plant feed variability. The average improvement of the miniplant over the mill model is a 4.2% recovery increase. Many parameters were tested including regrind size and chemistry, pyrrhotite circuit configuration and a number of reagent schemes.

PROJECT EVOLUTION

The flotation circuit reconfiguration engineering project completed FEL (front end loading) 1 and FEL 2 Stages as part of a larger project known as Clarabelle Mill Expansion and Recovery Project (CMERP). CMERP was well underway when INCO was acquired by Vale in 2006. A change in company ownership can be unsettling for major projects but the new management were brought quickly on board with the teams demonstrated successes and gave the blessing to continue in 2007. As these project study phases were completed during the mining boom, mill expansion was a key driver of the project economics. The project objectives were to increase mill production by over 20% and to improve the recovery of nickel by in excess of 3%. As the resources boom deflated in 2008 so did many projects and CMERP like many other expansions in the minerals industry was shelved.

Research work continued over the period of the FEL stages of CMERP on a large component of the five year mining plan to identify geometallurgical issues looking into the future. Using the Mineral Liberation Analyser (MLA) and the model developed by Ford et al., (2009) a significant component of the five year plan had deleterious components and below target metallurgical performance identified. Challenging deleterious minerals that were identified included:

Increased proportion of hexagonal (non-magnetic) pyrrhotite,

Occurrences of floatable gangue species such as talc and orthopyroxene,

Zones of high pyrite,

Fig. 8: Miniplant Recovery compared to Mill Model



Minor areas of complex textures, and

Millerite associated with high copper zones.

One of the ways of combating the deleterious components of the future ores was with the new flowsheet as it was able to better tolerate the feed changes. One of the most significant deleterious components of the future mine plan was a significant increase in hexagonal pyrrhotite. Hexagonal pyrrhotite historically comprised less than 15% of the pyrrhotite in the feed but its high floatability resulted in its dominance in the concentrate as shown in Fig. 5, causing significant dilution into the concentrate. The new flowsheet had two approaches to combat the hexagonal pyrrhotite, namely;

Do not attempt to segregate the two pyrrhotite polymorphs as this leads the circuit to a susceptibility to their variation in the feed.

To clean the rougher concentrate before sending it to copper nickel separation and thus allow an exit stream for entrained rock gangue and depressed hexagonal pyrrhotite.

The metallurgical team presented a scaled down version of the CMERP project to move forward with the flotation circuit modifications that represented the recovery component of the project. This flotation recovery component became referred to as the Challenging Ore Recovery Project (CORe) and was given approval to proceed to FEL 3 in early 2010 and commenced FEL 4 or project execution in late 2010. Project completion of the main phase of work is scheduled for July 2012 with the post commissioning components to be complete by July 2013.

The components of the CORe flow sheet that enable improved separation of pentlandite from pyrrhotite and rock from chalcopyrite are:

No magnetic separation ahead of the roughers,

Rougher concentrate cleaning with DETA/SS,

Pyrrhotite flotation with acid to produce a low sulphur tail,

Fine regrind in the pyrrhotite circuit,

Column flotation to reduce entrained pyrrhotite recovery, and

Rock/Sulphide separation.



During the engineering phases, the operating circuit at Clarabelle was required to increase primary ball mill power to combat the increasing copper head grades with higher proportion of rock in the feed increasing the grinding circuit work index. An additional 1500 kW was required in the primary ball mill circuit and this came by converting the non-magnetic regrind mill back to primary grinding duty. In addition to the loss of the 1500 kW, the small 750 kW regrind mill on the secondary rougher concentrate regrind duty was inspected and found to have suffered major washout damage to the discharge head. The CORe flowsheet at 9.5 Mtpa required 4500 kW of regrinding power to achieve the P80 of 40 micron in the pyrrhotite regrind circuit and only 3000 kW of existing installed capacity was available so a decision was made to include an M3000 Isamill in the flowsheet. The Isamill testwork had already been completed as a part of CMERP and an M4 Isamill had been used in the miniplant testwork so these data were available to make the change late in the engineering. The real estate released by the removal of the 750 kW mill would be used for the Isamill ensuring that very little engineering rework was required.

Equipment Selection

The prevalence of sulphides in the Clarabelle ore and in particular the coarse pyrrhotite had caused historical issues with sanding of solids in the flotation cells overcoming this problem was a slow learning process. In particular the cell scaleup from 100 ft3 cells to 38 m3 cells had resulted in significant operational issues during commissioning in 1991 and the prospect of scaling up to 100 m3 presented further operational concerns. These concerns were mitigated by site visits from the project team to discuss with end users and testing new design flotation mechanisms in the existing 38 m3 cells to demonstrate improved mixing capabilities. A three day plant trial in the existing circuit while operating the plant with all the pyrrhotite reporting to flotation by bypassing the magnetic separators also mitigated these concerns. This trial enabled plant measurements of tph/(m lip length) and tph/(m2 cell surface area) as this could not be scaled from miniplant testwork. Both the flotation mechanism and flotation plant trials were successful and the team developed a greater confidence to successfully scale up to the 100 m3 cells. There was no incentive to go larger than the 100 m3 cells as the large mass of concentrate to be removed resulted in a compromise between cell surface area and cell volume. The trade-off resulted in residence times of 35 minutes in roughing and 50 minutes in scavenging with sufficient surface area to remove the mass.

Fig. 9: The proposed CORe flowsheet



As mentioned, the Isamill testwork had been completed at the miniplant level and the Isamill chemistry had been thoroughly examined as a part of the P260E AMIRA research project. In particular the effectiveness of DETA/SS was evaluated in the flotation stage following grinding in an Isamill with ceramic media. The research defined the pulp chemistry required for effective pentlandite/pyrrhotite separation. It is proposed that by operating an Isamill in parallel with the existing regrind ball mills on the same feed stream, the demonstrated benefits will be the basis for the next step change in pyrrhotite rejection.

Other scale-up concerns came from the miniplant column as it was only 50 mm in diameter and it could not be used for scale-up factors. The use of a column in the pyrrhotite rejection circuit is a deviation from the VBN flowsheet that uses conventional U shaped cells. This will be the first application of a column in a pyrrhotite rejection circuit for Vale and so additional work was warranted. The feed stream chosen for the work was the pyrrhotite circuit operated at Clarabelle Mill during the work stoppage of 2009/10. In this circuit the magnetic separators were bypassed and all the pyrrhotite floated in the scavenger stage. The work was performed in laboratory flotation tests so that rate data could be collected for simulation studies and also by operating a 250 mm pilot venture aerated column and a 300 mm pilot Jameson cell on pyrrhotite circuit feed. The simulation and pilot plant data were then used to correctly size the flotation columns for number, lip length and surface area.

While Clarabelle operated with the magnetic separators bypassed during the work stoppage, the sulphur level in the rock tail was closely monitored. The CORe flowsheet uses acid in the scavenger stage to activate pyrrhotite and it was clear from laboratory studies of the rock tail that it did indeed further reduce the sulphur content of the low sulphur rock tails and in particular appeared to successfully improve the flotation of coarse pyrrhotite that was a key to the recovery improvement resulting from this flowsheet.

CONCLUSION

By understanding the mineralogy of the feed a flowsheet has been developed that addresses the key deficiencies in the current Clarabelle flowsheet and the challenging components of the more complex ores in the future mine plan. The key steps including regrind, cleaning with DETA/SS deal with both magnetic and non-magnetic pyrrhotite and the removal of magnetic separators from the head of the circuit reduces the rock that unnecessarily consumes regrinding power. Copper can be removed effectively early in the flowsheet and the rock entrained in the roughers has an exit stream with the sulphide/rock composites ending up in the regrind circuit.

The evolution of the flowsheet by successful step change projects and the testing in miniplant and plant and using features of the VBN flowsheet has allowed the CORe project to meet the challenges faced by the mining boom and bust cycle and survived a change of company ownership. Planned commissioning of the new roughers, cleaners and pyrrhotite rejection circuit is expected in July 2012. Challenges faced by brownfield construction await the team.

ACKNOWLEDGEMENTS

The authors would like to acknowledge the many Mineral Processors who were involved in both plant, pilot plant studies and the engineering phases of the project as well as Vale for permission to publish the paper.



REFERENCES

Ford, F D, Lee, A W, Davis, C, Xu, M, and Lawson, V, 2009. Predicting Clarabelle Mill Recoveries Using Mineral Liberation Analyzer (MLA) Grade-Recovery Curves, in Proceedings of the 48th Conference of Metallurgists 2009, pp 31-41 (Hamilton, Hart and Whittaker: Sudbury).

Kerr, A, Barrette, J, Bouchard, A, Labonte, G and Truskoski, J, 2003. The Mill Redesign Project at INCO’s Clarabelle Mill, in Proceedings of 35th Annual Meeting of the Canadian Mineral Processors 2003, pp 29-49 Ottawa.

Lawson, V, Kerr, A N, Shields, Y, Wells, P W, Xu, M and Dai, Z, 2005. Improving Pentlandite Pyrrhotite Separation at INCO’s Clarabelle Mill in Proceedings Centenary of Flotation Symposium 2005 pp 875-885 (The Australian Institute of Mining and Metallurgy: Melbourne).

Marticorena, M A, Hill, G, Kerr, A N, Liechti, D and Pelland, D A, 1994. Inco Develops New Pyrrhotite Depressant in Proceedings Innovations in Mineral Processing 1994, pp 15-33 (Editor, T. Yalcin).

Xu, Z, Rao, S R, Finch, J A, Kelebek, S and Wells, P, 1997. Role of diethyl triamine (DETA) in pentlandite-pyrrhotite separation – Part 1: Complexation of metals with DETA, Trans. IMM Section C, 106, pp C15-C20.

Xu, M, Quinn, P, Robertson, G, and Wilson, S, 2000. Development of A Two-Stage Middlings Circuit at Inco’s Clarabelle Mill, Part 1: Laboratory Studies, in Proceedings of 32nd annual meeting of the Canadian Mineral Processors 2000, pp 349-361 (The Canadian Mineral Processors: Ottawa).

Xu, M, Quinn, P and Wells, PF, 2003. A Moveable Mineral Processing Miniplant in Proceedings of 35th Annual Meeting of the Canadian Mineral Processors 2003, pp 317-329 (The Canadian Mineral Processors: Ottawa).

Xu, M, and Wells, P, 2004. Development of Copper/Nickel Separation at INCO in Proceedings of 36th Annual Meeting of the Canadian Mineral Processors 2004, pp 15-28 (The Canadian Mineral Processors: Ottawa).

Xu, M, and Wells, P, 2005. Laboratory and Miniplant Studies on Cu/Ni Separation. In Proceedings Nickel and Cobalt 2005 pp 347-358 (eds. J. Donald and R. Schonewille).



Flotation Process Control Optimisation at Prominent Hill

J Lombardi1, N Muhamad2 and M Weidenbach2

1. MAusIMM, Project Metallurgist, OZ Minerals Ltd – Prominent Hill, Ground Floor, 170 Greenhill Road, Parkside SA 5063. Email: [email protected]

2. MAusIMM, Senior Plant Metallurgist, OZ Minerals Ltd – Prominent Hill, Ground Floor, 170 Greenhill Road, Parkside SA 5063.

ABSTRACT

OZ Minerals' Prominent Hill copper-gold concentrator is located 130 kilometres south east of the town of

Coober Pedy in the Gawler Craton of South Australia. The concentrator was built in 2008 and commenced

commercial production in early 2009. The Prominent Hill concentrator is comprised of a conventional

grinding and flotation processing plant with a 9.6Mtpa ore throughput capacity. The flotation circuit includes

six rougher cells, an IsaMill for regrinding the rougher concentrate and a Jameson cell heading up the three

stage conventional cell cleaner circuit. In total there are four level controllers in the rougher train and ten

level controllers in the cleaning circuit for eighteen cells.

Generic proportional – integral and derivative (PID) control used on the level controllers alone propagated

any disturbances downstream in the circuit that were generated from the grinding circuit, hoppers, between

cells and interconnected banks of cells, having a negative impact on plant performance. To better control

such disturbances, FloatStar level stabiliser was selected for installation on the flotation circuit to account for

the interaction between the cells. Multivariable control was also installed on the five concentrate hoppers to

maintain consistent feed to the cells and to the IsaMill.

An additional area identified for optimisation in the flotation circuit was the mass pull rate from the rougher

cells. FloatStar flow optimiser was selected to be installed subsequent to the FloatStar level stabiliser. This

allowed for a unified, consistent and optimal approach to running the rougher circuit. This paper describes

the improvement in the stabilisation of the circuit achieved by the FloatStar level stabiliser by using the

interaction matrix between cell level controllers and the results and benefits of implementing the FloatStar

flow optimiser on the rougher train.

INTRODUCTION

OZ Minerals’ Prominent Hill operation is located 650 km northwest of Adelaide in South Australia and

approximately half way between the town of Coober Pedy and the BHP Billiton Olympic Dam operation. The

project area landscape is flat gibber terrain with few notable features. The climate is arid with high daytime

temperatures, high evaporation and low annual rainfall.

The Prominent Hill deposit was discovered in 2001 by Minotaur Resources with OZ Minerals securing 100

per cent ownership in 2005. Mining commenced in 2006 with plant construction in 2008 and first production

and sales in February 2009.



Prominent Hill is an iron oxide hosted copper-gold (IOGC) deposit with geological characteristics similar to

Olympic Dam (Reeve et al., 1990). Mineralisation consists of copper-gold breccia (80 per cent of the known

mineralization) broken down into four main types (Colbert et al., 2009):

chalcocite-bornite,

bornite-chalcopyrite,

chalcopyrite-bornite and

chalcopyrite-pyrite.

In addition, a considerable amount of “gold only” low sulphide content ore with a host rock of iron oxide was

identified in the upper area of the deposit. This material forms up to twenty per cent of the feed with the gold

being recovered as part of the flotation concentrate.

Mine reserve is currently 74.5 million tonnes (Mt) at a grade of 1.21 per cent copper and 0.67 gram per tonne

gold with an estimated mine life currently at eight years. Significant exploration activities occurring within the

lease boundaries have potential to add to the mine life.

By the end of 2009 the plant achieved the initial design target in throughput, copper recovery and produced

copper concentrate consistently to the grade specification. In 2010, the metallurgical teams’ focus shifted

from delivering successful commissioning and ramp-up of the plant to implementing continuous development

to improve the stability and metallurgical performance of the concentrator. An analysis performed during the

second quarter of 2010 showed there was an opportunity to improve the recovery of copper by stabilising the

entire flotation circuit and optimising the roughing circuit (Weidenbach and Rajiwate, 2010).

Standard PID control used on the level controllers in the flotation circuit propagated any disturbances

downstream from the hoppers, between cells and interconnected banks of cells. This was particularly evident

in the start-up of the plant after a planned or unplanned shutdown. As a result an improved means of

process control was sought.

Mintek had implemented its multi-variable control expert system and optimisation methods on over 40 plants

including Century (Muller et al., 2004), Sally Malay (Bennet et al., 2006) and Cadia Valley Operations from

which all showed improved control after the commissioning of an expert control system. As a result Mintek

were contracted by Prominent Hill to implement multi-variable control on the flotation circuit. Level

stabilisation control was applied to each of the cells level controllers, including the Jameson cell and to five

key hopper controllers. The second stage was to implement flow optimisation on the rougher train to improve

the mass pull consistency with the aim of improving recovery.

MINERAL SEPARATION PROCESS DESCRIPTION

The optimisation of the metallurgical performance at the Prominent Hill plant begins in the mine production

plan and ends with the finished concentrate product. The copper sulphide ore types are blended with the

“gold only” ore into the primary crusher located on the edge of the run-of-mine-ore pad. The proportions of

each ore are calculated based on availability and grade required to achieve a consistent and stable feed to

the process plant to meet production targets. The Prominent Hill flow sheet is shown in Fig. 1.



The primary crusher is a Fuller-Taylor type NT gyratory crusher. The primary crushing is open-circuit without

any screening and recycling of oversize particles. The nominal throughput is up to 3,600 tonnes per hour,

approximately three times the throughput of the grinding circuit.

The grinding circuit consists of a semi-autogenous grinding (SAG) mill followed by a ball mill. The SAG mill is

10.36 meter (m) diameter inside the shell with an effective length of 5.18m powered by twin hyper

synchronous wound motors to deliver the required 12 mega watts (MW) of energy. The SAG mill operates

with a variable speed drive operating between 8.6 and 10.6rpm. A single 12MW ball mill which is 7.3m in

diameter with a 10.4m effective grind length operates in closed circuit with hydro cyclones to produce a grind

size of 130 micron (µm).

The flotation circuit is the primary method of separation of the copper sulphide minerals. The flotation

equipment is listed in Table 1. Flotation cells are Outotec cells, other major equipment is from Xstrata

Technology.

Table 1 – Floatation equipment

Number of items Function Model Capacity in cubic meters

6 Rougher OK-TC-150 150

1 Jameson cell J5400/18 5.4 m diameter and 18 down comers

8 Cleaner 1 OK-TC-50 50



An Outotec Courier 5 SL online stream analyzer (OSA) is used to provide online information to the operators

regarding copper and iron in the feed, concentrate and tails from various points of interest around the plant.

This is coupled with an Outotec Courier PSI-500, which gives online sizing data of both the flotation feed and

reground rougher concentrate. This gives operators and metallurgists continual information about changes

happening in the plant so action can be taken to optimize recovery and concentrate grade.

The six rougher cells consistently recover between 87 per cent and 92 per cent of the total copper sulphide

minerals. The greatest loss of copper sulphide is from the >106µm size fraction. The typical upgrade ratio for

the rougher concentrate is between 8 to 10:1 and follows flotation feed grades.

The rougher concentrate is classified by a cluster of regrind hydro cyclones with the oversize sent to an

Xstrata IsaMill for regrinding. The IsaMill uses ceramic beads as the grinding media and typically reduces the

grind size of the rougher concentrate from 50-60µm to approximately 15µm. Regrinding significantly

increases the copper sulphide liberation which in turn helps in the rejection of diluents such as fluorine and

uranium.



The IsaMill discharge combines with the overflow from regrind hydro cyclones to feed the Jameson cell. The

Jameson cell acts as a “cleaner scalper” at the beginning of the cleaning circuit. It provides additional

cleaning capacity and depending on the blend of ore types produces between 30 to 50 per cent of the total

mass recovery of the concentrate. The Jameson cell typically produces high grade copper concentrate with

low fluorine and uranium levels with the wash water reducing the recovery of fine non-sulphide gangue.

The Jameson cell tailing is fed to the three stage conventional cleaning circuit. The re-circulating loads

include:

Cleaner 3 tailings to cleaner 2 cell 1,

Cleaner 2 tailings to the cleaner 1 feed hopper and

Cleaner 2 concentrate to the cleaner feed 3 feed hopper.

The cleaner 3 feed hopper is referred to as the cleaner 2 concentrate hopper. The cleaner 3 concentrate and

Jameson cell concentrate are the two streams which make up the final concentrate. While the rougher tail

and cleaner 1 tail both report to final tails. The cleaner block recovers 98 to 99 per cent of the copper feed to

the cleaners and based on the overall mass balance and its tailings contributes between 8 and 13 per cent of

copper sulphide losses to the final tailings stream.



Fig. 1 – Prominent Hill flow sheet



EXISTING FLOTATION PROCESS CONTROL

The existing process control prior to the implementation of the flotation expert system included PID control

on each of the cells. The processing technicians would adjust the levels of the cells and aeration rates by

changing the set point on the Citect interface. The level sensor measures the actual cell level (usually one

level controller per two cells) and converts this measurement to a signal sent to the distributive control

system (DCS) where it compares the valve to the set point. The level of the cells is the process variable (PV)

and the controller is the position of the dart valve or current value (CV). Based on the controller tuning

parameters in the DCS, the controller calculates the control output and moves the position of the dart valve

accordingly. The issue with this type of control in the flotation circuit is that the single PID controller acts

independently of the other cells and can often not react to any disturbance without further propagating the

disturbance downstream.

Each of the transfers of slurry between cells that feed a part of the flotation circuit delivers the flow utilising a

single duty pump with a variable speed drive. The tuning of the pump variable speed controller is itself tightly

tuned to the level of the hopper, with a typical relationship seen below in Fig. 2. The pump speed would react

to any variation in the hopper level and in turn increase the flow rate to the subsequent flotation train. The

increase or decrease in flow rate to a bank of cells would show the same reaction to the level set points in

the following train and propagate this disturbance around the circuit. As a result plant performance was sub

optimal.

Fig. 2 – Tightly controlled pump variable speed based on hopper level



PROJECT COMMISSIONING

Once the feasibility and proposed design for the expert control system had been completed the

commissioning of the project was broken into four different phases:

Infrastructure checks,

Level stabilization,

Rougher mass pull optimization and

Training.

This is inline with the Mintek approach to process control hierarchy (Singh et al., 2003) to ensure the best

possible control.

Infrastructure

The infrastructure used to physically measure and control the variables was checked and found to be in

good working order. Each of the Outotec level sensors and both dart valves are regularly inspected. Under

normal flow rates the dart valves are sized to use one of the dart valves however if the first valve is 100 per

cent open but still not reaching a set point level, the second valve will also open and vice versa in the

opposite direction.

The layout of the plant had been well thought out to ensure that there is good accessibility and visibility in

regards to the froth of each cell. A walkway next to each train which follows the height difference in cells was

installed to maintain visibility of the froth as a process technician or metallurgist is walking along each train.

Each cell has also a walkway on top of the cells to allow access to the level sensors and positions for dart

valve maintenance. The accessibility in the float cells on the shutdown is made easier by an additional

walkway below, adjacent to the hatches of the flotation cells which allows for easy access for inspections

and maintenance.

The flotation area is also protected from the cross wind by metal sheeting on both sides of the building and

made a more comfortable area to work in due to the roof. These attributes make spending time visually

inspecting any difference in the equipment more conducive to the process technicians and metallurgists.

Level Stabilisation

Prior to the implementation of the level stabilisation the communications between the DCS and the expert

system computer were established, followed by testing of the DCS programming. The level stabiliser was

applied to the four rougher controllers and the ten cleaner controllers. The level stabilisers were applied to

five key hoppers by manipulating the variable speed pump except for the Jameson cell where the recycle

valve was used. The following hoppers which interact with the flotation trains included:

IsaMill feed hopper,

Jameson feed hopper,

Cleaner 1 Concentrate Hopper,

Cleaner 2 Tails Hopper and

Cleaner 2 Cons Hopper.



The level control on the pump would utilise the hopper level, therefore allowing the level to fluctuate and the

pump level to move as little as possible maintaining a constant flow to the rest of the circuit. The level

stabilisation was left working for a month prior to commencing the next stage of the project.

Rougher Mass Pull Optimisation

The flotation process technicians were spending the largest amount of time on one single task continuously

adjusting the level set points of the four rougher controllers to maintain mass flows. Each shift was running

the rougher set points in a different way. The mass pull optimisation system was implemented on each of the

four controllers in the rougher circuit. The expert system would now control the set points of the rougher

levels based on different parameters set by the metallurgists, which could be optimised further in the future

pending information regarding the rougher performance. The mass pull system would pull the floats based

on a set input flow rate and the parameters for relative pull rates and base levels within restricted minimum

and maximum levels for each cell.

Training

As with any new process or equipment, training of the users of the system is paramount to the long term

successful operation. The training process can often be challenging in fly-in, fly-out sites. The process

technicians work on an eight days on and six days off rotation. Several training sessions were delivered to

cater for the four different shifts.

Training was delivered not only regarding the changes made on the Citect system but also covered the

principles of control used by the flotation expert system. The two stage implementation made the training

easier as the idea of FloatStar was well entrenched by the time the rougher mass pull optimisation was

implemented. This mass flow controller was more of a change on the way to operate the flotation circuit than

the level stabilisation which was mainly in the background.

PROJECT OUTCOMES

Improvement in Stabilisation on Start-up

A comparison was performed based on the amount of time taken to reach stabilisation of the circuit after

start up with the expert system on and off. Faster stabilisation was recorded for the start-ups with the expert

system on. Fig. 3 and 4 show the difference between the start-up with the expert system on and off (the time

scales have been adjusted for comparison on the two figures). The OSA online system recorded a higher

recovery when the expert system was on. This is also denoted in the rougher concentrate hopper level

stability (represents the pull rates from the primary flotation circuit) and more stable recovery output.



FLOTATION BEHAVIOUR ON START UPBEFORE THE EXPERT SYSTEM INSTALLATION

0

20

40

60

80

100

18:00 19:12 20:24 21:36 22:48 0:00 1:12 2:24 3:36 4:48 6:00

0

1000

2000

3000

4000

5000

6000

Rougher Conc Tank Level PV Flotation Recovery Rougher Conc Tank Level SP

Flot Feed Mass Flow Mill Feed

G 1

Fig. 3 – Flotation on the start on the mill without expert system

FLOTATION BEHAVIOUR ON START UPAFTER THE EXPERT SYSTEM INSTALLATION

0

20

40

60

80

100

18:00 19:12 20:24 21:36 22:48 0:00 1:12 2:24 3:36 4:48 6:00

0

500

1000

1500

2000

2500

3000

3500

4000

4500

Rougher Conc Tank Level PV Flotation Recovery Rougher Conc Tank Level SP

Flot Feed Mass Flow Mill Feed

FIG 2

Fig. 4 – Flotation on the start on the mill with expert system

Analysis on plant start up data post the installation of the expert system demonstrated that the circuit

reached stabilisation with a significant improvement equating to more than a 50 per cent reduction in start up

time over the existing PID control system.



Historically a 6.65 per cent copper recovery loss (Table 2) has been seen during the first 12 hours (shift

performance) after start up before the installation of the expert system. After the installation the recovery loss

reduced to 1.08 per cent.

Table 2 – Average recovery losses due to start - up

Before Installation of Expert System After Installation of Expert System

Normal shift recovery

Start up shift recovery

Difference Normal shift

recovery Start up shift

recovery Difference

91.7 88.5 3.2 92.1 91.3 0.8

91.7 85.5 6.2 93.6 92.8 0.8

87.2 80.2 7 92.3 90.5 1.8

91.5 81.3 10.2 87.1 86.2 0.9

Annualised benefit based on an average of 40 start ups for planned and unplanned shutdowns was a 0.5 per

cent copper recovery increase. There is a planned shutdown every six weeks and a major shutdown for a

reline of the SAG mill every 13 weeks with a re-torque of the new bolts. Additional shutdowns are breakdown

related.

Level Stabilisation in day to day Operations

With the expert system in operation most of the level control across the flotation circuit has been improved to

be between 15 to 25 per cent standard deviation from the set point as identified via comparison of three

months of data collected since FloatStar had been in operation. Fig. 5 shows the performance of the cleaner

train before and after while Fig. 6 shows the lower average standard deviation of error between the set point

and PV with the expert system on, providing a better control response to maintain stability of the circuit.

C L E A N E R 1 F L O T A T IO N L E V E L C O N T R O L P E R F O R M A N C E

-

1 0

2 0

3 0

4 0

5 0

6 0

7 0

8 0

9 0

1 0 0

0 6 :0 0 0 8 : 2 4 1 0 : 4 8 1 3 : 1 2 1 5 :3 6 1 8 :0 0 2 0 : 2 4 2 2 : 4 8 0 1 :1 2 0 3 :3 6 0 6 : 0 0

C le a n e r 1 C e l l 7 /8 L vl S P C le a n e r 1 C e ll 5 / 6 L vl S P

C le a n e r 1 C e l l 3 /4 L vl S P C le a n e r 1 C e ll 1 / 2 L vl S P

E X P E R T S Y S T E M O F F E X P E R T S Y S T E M O N

Fig. 5 – Cleaner 1 level control performance without and then with Expert System



Rougher Flotation Level Control Performance During Expert System On and Off

-

0.10

0.20

0.30

0.40

0.50

0.60

0.70

0.00 0.50 1.00 1.50 2.00 2.50 3.00 3.50 4.00 4.50

Standard Deviation of Error

Pro

bab

ilit

y

Expert System OFF Expert System ON

Expert System ONSTD of error between SP - PV

Average 1.99

Expert System OFFSTD of error between SP - PV

Average 2.50

21 % Level Control Improvement

Fig. 6 – Rougher level control performance with expert system on and off

Hopper Stabilisation in day to day Operations

Fig. 2 demonstrated that the hopper pump reacts very quickly to any changes in the hopper level (pre expert

system), propagating the mass flow increase/decrease downstream in the circuit. After the implementation of

the hopper level stabilisers, the pump variable speed makes minimal movements and allows the fluctuation

of the hopper as highlighted in Fig. 7 below. There was a reduction of the pump speed/recycle movements of

the five flow rates from the hoppers.

Fig. 7 – Cleaner 2 Hopper level



Rougher Mass Pull Optimisation

The rougher mass pull optimiser smoothed out the level set point changes in the rougher flotation circuit, as

seen in Fig. 8 and gave greater consistency in terms of mass pull allowing recovery to be maximised. In

addition the mass pull optimiser allowed the flotation process technician to spend more time working on

other tasks around the flotation circuit which in turn created positive feedback by the process technicians for

simplifying the process control.

Fig. 8 – Rougher Flotation Level Control

Summary of Benefits

The implementation of the expert system provided the following benefits:

Reaches stability faster after start-ups resulting in overall increased recovery,

Overall stabilisation of the flotation circuit which improved by 15 to 25 per cent,

Reduced movement of pumps and propagated disturbances,

Simplified operation of the rougher circuit and mass pull rates and

Unified approach by all crews to the operation of the level set points of the flotation circuit.

It would be inferred that higher recovery is being achieved under normal day to day operation with improved

control attributed to the expert system but without conducting an on/off trial it is difficult to attribute further

recovery gain to the 0.5 per cent proven from start up data analysis.



CONCLUSIONS

The FloatStar level stabiliser and optimiser for the roughers were selected for installation on the flotation

circuit to recitify for the previous poor interaction between the cells. Multivariable control was also installed on

the five hoppers to maintain consistent feed to the cells and to the IsaMill.

The project was fully implemented by November 2010 and benefits have been measurable in terms of

recovery for the start-up of the plant. The rougher mass pull optimiser allows the metallurgists to further

optimise by changing the parameters in the Floatstar controller.

The system has had positive feedback by the process technicians and has been in use over 99 per cent of

the time when the plant is operational. The rougher mass pull optimiser has simplified the process of

maintaining constant mass pull to the rest of the downstream circuit.

The flotation expert system delivered significant value to the bottom line of Prominent Hill Operations, with

near seamless installation and extremely quick benefit delivery.

ACKNOWLEDGEMENTS

The authors’ would like to thank OZ Minerals Prominent Hill and Mintek for the permission to publish the

paper and plant data.

REFERENCES

Bennet C, Knobauch J, Foggo C, Strobos P J, Van der Spuy, D, Dorfling C, Smit H S, 2006 Automation of

the Kimberly Nickel Mine Floatation Operation Using an Advanced Control System, in Proceedings

Metallurgical Plant Design and Operating Strategies, pp 66-78 (The Australian Institute of Metallurgy

and Mining: Melbourne).

Colbert, P J, Munro P D, Yeowart G, 2009. Prominent Hill Concentrator- Designed for Operators and

Maintainers, in Proceedings Tenth Mill Operators’ Conference, pp 23-31(The Australian Institute of

Metallurgy and Mining: Melbourne).

Muller B, Smith G C, Smit S, Singh A, Strobos P J, Reemeyer L, 2004 Enhancing Floatation Performance

with Process Control at Century Mine, in Metallurgical Design and Plant Operation Strategies (The

Australian Institute of Metallurgy and Mining: Melbourne).

Reeve, J S, Cross K C, Smith, R N and Oreskes, N, 1990 Olympic Dam copper-uranium-gold-silver deposit,

in Geology of the Mineral Deposits of Australia and Papua New Guinea, vol 2, pp 1009-1035, (The

Australasian Institute of Metallurgy and Mining: Melbourne).

Weidenbach, M and Rajiwate, F, 2010, Prominent Hill, Feasibility study for installing Floatstar level stabilizer

on the flotation circuit, OZ Minerals internal technical note PH/TN/3400/10-133.

Singh, A, Louw, J J, Hulbert, D G, 2003 Floatation stabilization and optimization in Journal of The South

African Institute of Mining and Metallurgy pp 581-588.


Ore Ageing Test Work for the Ok Tedi Skarns

M Morey1 and R Cantrell2

1. MAusIMM, Ok Tedi Mining Limited, PO Box 1, Tabubil, WP 332, Papua New Guinea. Email: [email protected]

2. FAusIMM(CP), Mineral Technology and Management Services, 35 Yuulong Street, North Tamborine Qld 4272. Email: [email protected]

ABSTRACT

Ok Tedi Mine is located eighteen kilometers from the Irian Jaya border in the Star Mountains, Western Province, Papua New Guinea. Sequential flotation concentrators produce a copper / gold concentrate for market, and a pyrite concentrate for environmental containment, with low sulphur tailings from the pyrite flotation sequence being discharged to river. Strict sulphur assay targets are set for the riverine discharge.

Current porphyry and skarn ore blends from open pit operations at Mt. Fubilan will cease by 2017. However, drill cores indicate an economic ore body in the skarn ores surrounding and below the current pit limits. With proposed mining from a western pit wall ‘cut back’ and an eastern pit wall underground operation, potential for ore ageing in sub level cave underground stockpiles is apparent.

With little information in the literature, Ore Ageing test work at prefeasibility level consisted of multiple one tonne stockpiles of -50 mm crushed rock. For the bankable feasibility study, a successful test work program was implemented utilizing 40 tonne ROM ore stockpiles composed of exposed pit extensions of future ores. Realistic ageing profiles of Skarn ores were obtained from one tonne monthly samples subjected to laboratory testing. Test work outcomes included:

• Similar ore ageing profiles from both the pre feasibility study (PFS) and feasibility study (FS) programs, suggesting ore susceptibility to oxidation has greater effect in ore ageing than rock size for Ok Tedi skarns.

• Laboratory rougher flotation of pyrite showed significant variation in sulphur recovery rate, suggesting three classes of pyrite floatability.

• Laboratory rougher flotation of both copper and pyrite for up to 19 months ore ageing from 11 stockpiles, suggested three stages in the ore ageing process.

INTRODUCTION

From 2017 to 2021, Ok Tedi Mining intends to process a mine life extension (MLE) 100% skarn orebody surrounding the current pit. The MLE skarn reserve is 80% sulphide skarn, which is predominantly magnetite but which also contains significant pyrite. The 20% endoskarn component has a mineral matrix of ferro-magnesium silicates. Endoskarn is the


transition ore zone between the volcanic intrusive (porphyry) and the skarn mineralization generated by a mixing of the intrusive with host limestone rock. A plan view of the skarn resource is shown in Fig. 1.

424200 N

424000 N

423800 N

423600 N

423400 N

423000 N

423200 N

422800 N

Gold Coast

315000 E 315500 E

Berlin

Paris

N

Fig. 1: Plan view of MLE ore reserves in relation to the OTML final pit design. The location of the proposed Gold Coast underground mine is circled.

Ore ageing is a common term for the oxidation of sulphide minerals. Oxidation is enhanced by:

• Exposure of broken rock to moisture and oxygen,

• Increased exposure of surfaces through blasting,

• Electrochemical interaction between different minerals (referred to as galvanic coupling),

• Impurities in the mineral sulphide crystal structure, and

• Predisposition of minerals to oxidation, or mineral semi-conductor type (Harris, P, et al, 1985).

The effects of oxidation of sulphide minerals in froth flotation include:

• Formation of particle surface poly-sulphide with subsequent enhanced flotation (in early stages),

• Formation of particle surface hydroxide and gypsum precipitates (flotation depressants), resulting in reduced flotation recovery,

• Reduced adsorption of copper sulphide collector on value minerals, due to less availability of particle surface sulphide, resulting in reduced flotation recovery, and

• Formation of clay agglomerates in flotation for heavily weathered ores.


With planned extraction methods to include sub level caving in the Gold Coast region, a significant potential exists for delays in processing and ageing of blasted rock. Furthermore, ores adjacent to fault zones and underground aquifers may also be subject to ageing over geological time, showing oxidation effects even if processed immediately after blasting.

An assessment of the effects of ore ageing on Ok Tedi skarns has enabled an estimation of the potential effects on processing delays.

METHOD

The primary method of assessment was via an ore ageing stockpile sampling program. Supportive evidence was obtained from mineralogical analysis and flotation tests of ‘un-aged’ skarns.

The PFS ore ageing program used one tonne stockpiles of rock crushed to -50mm. The small stockpiles and finer size distribution as compared to that of run of mine (ROM) ore were believed to potentially enhance mineral oxidation rates due to the increased particle surface area at fine sizes. The MLE FS ore ageing program was designed to be as realistic as possible. The sampling process is shown pictorially in Table 1 and included:

• 40 tonne stockpiles of ROM ore extracted from exposed pit extensions of MLE skarns,

• No size reduction of ore prior to sampling,

• Collection of statistically representative one tonne samples from ageing stockpiles on a monthly basis,

• Quartering of the one tonne product after crushing, and

• Laboratory testing via a standard rougher flotation test for copper, followed by a standard rougher flotation test for pyrite.

Table 1: Monthly ore ageing stockpile sampling summary procedure

40 tonne stockpiles at Geneva Ridge Quarry

Monthly 1 tonne sampling of stockpiles

Breaking of oversize rocks

Transfer to Geneva 2ndry crusher

Collection of crushed rockon canvas

Cone quartering abagging


Ore ageing sampling points were dispersed across the MLE ore zones. However, due to the high variability of Ok Tedi skarns, results may be considered as ‘snapshots’ pertaining to the collected samples rather than a full representation of skarn behavior for the total orebody.

Aged ore samples were assessed via standard rougher flotation tests for copper and pyrite metallurgical response. Due to the recovery versus grade trade-off effect, copper rougher flotation tests were assessed at a constant rougher concentrate grade of 15% Cu.

Pyrite rougher flotation tests were compared at ultimate sulphur recovery, due to the highly variable concentrate grades evident in this work.

No adjustment was made for head grade effects, although the 40 tonne ROM stockpiles were not homogenous. Skarn head grades were variable.

RESULTS

Mineralogy

Ok Tedi skarn mineralogy is often described as ‘altered’, indicating that changes have occurred within the mineral suite over geological time. This classification pertains to:

• The alteration of magnetite to siderite, particularly when surrounding chalcopyrite grains, and chalcopyrite rimming with chalcocite, covellite or bornite, and

• Alteration of copper minerals is generally greater in endoskarns, with significant chalcocite / bornite content, and potential natural floatability. Copper oxide mineralisation may also occur as cuprite and malachite in endoskarns.

Copper mineral deportment in mineralogical samples examined within the PFS and FS is shown in Table 2.

Table 2: Copper deportment in mineralogical samples of the PFS and FS

Ore Zone Ore Type DDH/BP ID Chalcopyrite Bornite Chalcocite Covellite Other Comment

Berlin Magnetite Skarn DDH963-4 64 7 6 23 0 MODA - OpticalBerlin Magnetite Skarn DDH968-4 86 1 7 6 0 MODABerlin Endoskarn DDH932-1 95 1 3 1 0 MODABerlin Massive Mag. Skarn DDH1009-3 89 4 0 7 0 MODABerlin Magnetite Skarn BP AH153 98 0 1 1 0 MODABerlin Magnetite Skarn BP AI133 94 1 1 4 0 Plant Feed - MODABerlin Magnetite Skarn DDH1015 77 1 22 0 0 MODAGold Coast Magnetite Skarn BP AJ76 73 12 14 0 0 MODA

MLE zones 70:10:10:10 BlendBP AI133, BP AJ122, BP AH161, 89 3 6 3 0

Pilot plant ore blend - MODA

Gold Coast Magnetite Skarn DDH925, DDH927 99.7 0.3 0 0 0 MLABerlin Pyrite Skarn BP AG88 93.8 3.5 0.9 0 1.7 QEM SCANBerlin Magnetite Skarn DDH934 93.7 3.6 0.3 0 2.3 QEM SCAN

Berlin Endoskarn BP AE140 94.9 2.2 2.9 QEM SCAN

Gold Coast Magnetite Skarn n/a 91.6 MCI Optical8.4

Percentage DistributionCOPPER DEPORTMENT IN FEED SAMPLES

Chalcopyrite associations include pyrite, iron oxides and iron carbonates in sulphide skarns and ferro-magnesium silicates in endoskarns.


Gold mineralization is strongly associated with sulphides, particularly of copper, with the strongest gold to copper correlation being between copper and gold grades in feed and concentrate.

Iron sulphide deportment in Ok Tedi skarns is variable. From limited assessments, pyrite and marcasite are present in significant quantities, with only traces of melnikovite. Pyrite and marcasite are associated minerals in Ok Tedi skarns.

Flotation of Fresh Ores

Variability tests for the MLE skarn ore zones were performed via standard rougher flotation tests for copper followed by pyrite using fresh ore samples. DDH (Diamond Drill Hole) core samples obtained were distributed from across the ore zones. Fig. 2 illustrates a suite of pyrite rougher flotation test results as sulphur recovery and tailings sulphur grade as a function of laboratory flotation time.

0

10

20

30

40

50

60

70

80

90

100

0 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15

Cu

mu

lati

ve S

Re

co

very

(%

)

Rougher Flotation Time (Min)

0

1

2

3

4

5

6

7

8

9

10

11

12

13

0 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15

Cu

mu

lati

ve S

Tai

lin

g A

ssa

y (%

)

Rougher Flotation Time (Min)

Fig. 2: Variability pyrite rougher flotation tests showing sulphur recovery and sulphur tailings grade versus time.

Significant variability in flotation performance is shown in Fig. 2. From the second plot at zero time, significant variation in the sulphur head grades is also apparent. Observations from Fig. 2 indicated that there were three stages of flotation, suggesting three classes of pyrite with different flotation rates. The three stages are approximately delineated by the dashed lines. The three classes of pyrite may be identified as:

1. A fast floating fraction, recovered in up to approximately 45 seconds of laboratory flotation,

2. A fraction with intermediate flotation rate, recovered with up to approximately 3.5 minutes of laboratory flotation, and

3. A slow floating (or non floating) fraction after 3.5 minutes flotation time.

It is apparent from Fig. 2 that the proportions of each of the three classes of pyrite varied between samples. This is particularly evident in regards to the slow floating or non floating fraction. As liberation is not believed to be an issue for pyrite flotation in these tests, the


explanation for the slow floating component lies in oxidation or ore ageing over geological time. This effect was confirmed for some samples via sulphidization prior to flotation (Morey, 2011).

Flotation of Aged Ores

Flotation results from the 11 FS ore ageing stockpiles were variable, and are considered to be examples of potential ore ageing outcomes. However, they are not proportionately representative of the ore zone’s metallurgical performance.

Copper flotation from Aged Ores

Monthly samples of ore ageing stockpiles were subject to standard rougher flotation tests. For comparisons, copper recovery results were normalized to 15% Cu grade via a grade versus recovery trade-off. Head grades did vary from the ROM stockpile sampling. However, no adjustment was made for head grade variation. Results for copper flotation from Berlin ore zone are presented graphically as recovery versus ageing time in Fig. 3. Results for all stockpiles from both the PFS and FS programs are shown in Table 3.

0

10

20

30

40

50

60

70

80

90

100

0 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16

Time (Months)

Berlin Skarns - Cu Recovery @ 15% Cu Grade

Cu

Re

co

very

(%)

Fig. 3: Ore ageing effects shown as copper recovery at normalized grade for Berlin skarns, with dashed lines indicating most common stages in ore ageing.

From Fig. 3 it is apparent that:

• All skarn samples aged differently.

• Some skarns appeared to be aged prior to blasting, although later sampling showed improved performance. This outcome suggests pockets of aged ore were present in the midst of less aged or pristine skarns. Pockets of aged ore prior to blasting may be associated with rock fissures and fault lines.

• The interpretation of a month of poor copper recovery followed by a return to high flotation recovery levels is that a pocket of aged ore was assessed in the month of poor flotation, and that a significant proportion of that skarn remained ‘less aged’.


• Despite the variability in flotation responses, at least three separate stages of ageing appeared to be evident (as estimated by the dashed lines in Fig. 3). This finding concurred with observations noted in the PFS ore ageing evaluation.

• For each skarn assessed, the stages of ageing differed.

An approximate summary of the three stages of ageing for each PFS and DFS ore ageing stockpile is shown in Table 3. The summary includes the approximate copper recovery value and duration in weeks for each stage of ore ageing.

Table 3: Copper flotation recovery and the stages of ore ageing

Study Ore Zone Ore TypeSample

ID

Fresh Ore Cu

Rec. (%)

Sustained Until

(Week)

Stage 2 Cu

Rec. (%)

Sustained Until

(Week)

Stage 3 Cu

Rec. (%)

Sustained Until

(Week)

PFS Berlin Endoskarn AG88 91-92 22 84-85 52 N/APFS Berlin Mag Skarn AG87 N/A <10 81-86 19 65-75 52PFS Berlin Pyrite Skarn AG88 N/A <10 88-90 22 75PFS Gold Coast Endoskarn AE140 94 <4 87 4 75-82 31PFS Gold Coast Mag Skarn AF135 N/A <3 89 3 77-84 23PFS Gold Coast Pyrite Skarn AF155 97.5 16 92 N/AFS Berlin Pyrite Skarn AJ119 N/A <4 90-91 22 VariableFS Berlin Pyrite Skarn AH153 98 <4 95 16 91FS Berlin Mag Skarn AJ130 95 <4 85 22 75FS Berlin Endoskarn AJ135 N/A <4 82 12 50 27FS Berlin Mag Skarn AI133 N/A N/A 85 27 Variable 40FS Gold Coast Mag Skarn AJ77 93 22 91 44 80FS Gold Coast Mag Skarn AK68 93-97 42 N/AFS Gold Coast Endoskarn AJ122 91 <4 87 8 80 44FS Gold Coast Endoskarn AJ49 94-97 22 80 48 VariableFS Paris Pyrite Skarn AH161 92-94 <8 85 22 0FS Paris Endoskarn AJ11b 91 22 85 36 40-60

From the Table 3 (and also the Fig. 3 above), it is apparent that:

• The initial period of pristine flotation prior to a reduced flotation recovery due to ageing effects, varied between skarns. Some samples were not pre-disposed to oxidation, and maintained pristine recovery for many months. This effect was not related to ore type and included 3 endoskarns, 2 magnetite skarns and 1 pyrite skarn.

• More than half of the skarns showed pristine flotation immediately after blasting, but indicated somewhat reduced recovery within a few weeks of ore exposure and ageing. These samples were predisposed to oxidation. The average decrease in copper recovery for the transition from Stage 1 to Stage 2 of ore ageing was 7%.

• This second stage of somewhat reduced copper recovery was generally sustained for a longer period than the first stage. However, results varied significantly. The shortest duration of Stage 2 ore ageing was for Berlin endoskarn from BP (Blast Pattern) AJ135. This skarn gave a low initial copper recovery indicating this was a Class 3 skarn, pre-oxidised in the pit over geological time. The sample was also predisposed


to further relatively rapid oxidation after blasting. It showed un-economic copper flotation recoveries after an ageing period of between 3 and 4 months. The most common duration of ore ageing ‘Stage 2’ was 22 weeks, or between 5 and 6 months.

• A number of skarns showed pockets of significantly aged ore that generated periodic decreases in copper flotation recovery for tests performed during ore ageing Stage 2. However, a significant proportion of these ores remained relatively unoxidized, as indicated by the higher copper recoveries also generated.

• A transition to Stage 3 of the ore ageing process was similar to the transition to Stage 2 for some skarns. The median decrease in copper recovery was 11%. However, copper recovery results were variable in this stage of ore ageing, with a suggestion of a potentially un-economic metallurgical performance for a number of areas containing this type of skarn.

Pyrite Flotation from Aged Ores

Pyrite rougher flotation was performed on the tailing from selected copper rougher flotation tests. Not all ore ageing stockpiles contained pyrite at recoverable levels. Six stockpiles provided meaningful results in relation to pyrite flotation which are summarised graphically as ultimate sulphur recoveries in Fig. 4. Sulphur head grades, concentrate grades and recoveries showed significant variation. Head grade variation was due to:

• Significant variation in sulphur grades in Ok Tedi skarns,

• Non-homogenous stockpiles, and

• Varying sulphur recoveries within the copper rougher flotation stage prior to pyrite flotation (including bulk sulphide flotation in copper roughing after five months ageing for Berlin pyrite skarn from BP AH153).

0

10

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30

40

50

60

70

80

90

100

0 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16

Time (Months)

Skarns - S Ultimate Recovery (%)

Su

lph

ur

Re

co

very

(%

)

Fig. 4: Suphur (pyrite) flotation from ore ageing stockpiles


From Fig. 4, the following observations have been made in relation to pyrite ageing:

• Ore ageing stages in pyrite flotation correlate somewhat to those determined from copper flotation performance.

• More variation in pyrite flotation performance is observed as compared to that in copper flotation, particularly in the early stages of ageing.

• Two results show initial pristine flotation quickly followed by early oxidation and reduced sulphur recovery. This outcome is an identical trend to that shown for copper in Fig. 3.

• As was observed in the copper assessment, some skarns showed no predisposition to oxidation, and maintained excellent pyrite flotation recovery for many months. This response was noted in particular with the Berlin magnetite skarn from blast pattern AI133 which showed pristine flotation with up to nine months of ageing. However, the second sample that maintained high recovery in this assessment contained only two percent sulphur and therefore was subject to excess flotation capacity under the standard laboratory flotation test work regime, resulting in the high recoveries into a low grade concentrate.

• The decrease in flotation recovery due to oxidation effects did not correlate between copper and pyrite flotation. For some ores, acceptable copper flotation was followed by poor pyrite flotation.

Fresh ore DDH1092 Composite 1 endoskarn, from a fault region in the Berlin ore zone, indicated different ageing effects in copper and pyrite rougher flotation, as shown in Table 4. Pyrite flotation was carried out on the tailings product from the preceding copper rougher flotation sequence.

Table 4: Mineral recovery in rougher flotation from DDH1092 composite 1, endoskarn

Cu Ro

DDH 1092 Cu

Sulp

hide

Wei

ght

Pyr

ite

Mar

casi

te

Mel

niko

vite

Cha

lcop

yrit

e

Mag

neti

te

Gan

gue

Conc 14.9% 7% 11.5% 14.4% 0% 0.2% 3.3% 70.6%

Tail 0.12% 93% 1.2% 2.9% 0.3% 0.1% 3.7% 91.8%

Feed (Calc) 1.7% 100% 1.9% 3.7% 0.3% 0.1% 3.7% 90.4%

Recovery 93.4% 40.8% 26.3% 0.0% 12.6% 6.0% 5.2%

Pyrite Rougher Flotation


Observations from Table 4 include:

• Copper rougher flotation performance was good.

• Pyrite rougher flotation performance by mineral type was determined by a mineralogical assessment. Approximately two thirds of the FeS2 in the pyrite float was marcasite with only one third present as pyrite.

• Pyrite and marcasite rougher flotation performance was poor, suggesting these minerals were significantly more oxidized than chalcopyrite.

• Marcasite flotation performance was significantly worse than that for pyrite, suggesting that marcasite was more oxidized than pyrite.

One observation was found in the literature supporting a tendency for marcasite to oxidise more rapidly than both pyrite and chalcopyrite. (Stockwell et al., 2003). However, according to the thermodynamic hierarchy of sulphide mineral rest potentials, the order of oxidation of associated minerals in Ok Tedi skarns, via galvanic interaction should be:

• Bornite

• Covellite

• Chalcopyrite

• Marcasite

• Pyrite

The hierarchy of oxidation rest potentials is shown in Table 5 (Majima, 1969).

Table 5: Rest potentials for sulphide minerals at pH 4.

MineralRest Potential

(V vs SHE)

Pyrite 0.66Marcasite 0.63Chalcopyrite 0.56Sphalerite 0.46Covellite 0.45Bornite 0.42Galena 0.4Argentite 0.28Stibnite 0.12Molybdenite 0.11


DISCUSSION

Points for discussion and clarification include:

1. The correlation between three flotation classes apparent in pyrite rougher flotation, and the three stages of oxidation evidenced in ore ageing test work, and

2. The apparent discrepancy between sulphide mineral rest potentials and the oxidation rate of chalcopyrite, marcasite and pyrite in Ok Tedi skarns.

Three Flotation Classes of Aged Ores

In relation to Point 1 above, it is logical to conclude that the three classes of mineral flotation in pyrite rougher tests indicated in Fig. 2 correlate with the three oxidation states of minerals, as described in Fig. 3, Fig. 4 and Table 5. Reasons include:

• The three classes of pyrite show three flotation rates, indicating three different states of oxidation (as FeS2 is considered to be well liberated at the 106 micron grind used).

• The three different states of oxidation are expressed in the three stages of ore ageing, apparent from monthly sampling and flotation tests.

These observations suggest that although test work identified three stages of ore ageing, it may be common for all three stages of ageing to coexist within one blended ore sample.

The oxidation of sulphide mineral particle surfaces follows a sequence of reaction steps, the rate of each step being determined by factors such as mineral semiconductor type and purity amongst others. Pure sulphide mineral is very slow to oxidise. Induction times of up to forty hours prior to a detectable uptake in oxygen have been recorded (Eadington and Prosser, 1969).

The electrochemical oxidation reaction of sulphide minerals occurs at a mixed potential, at which the rate of cathodic reduction of oxygen (one reaction half cell) equals the rate of oxidative dissolution of the mineral. This reaction rate is also dependent on the mineral type, and appears to correlate to mineral conductivity (Nowak and Chmielewski, 1994).

It has been suggested that n-type semiconductor mineral sulphides (metal rich) are slower to oxidize (Richardson and Maust, 1976). Metal vacancies in the mineral lattice tend to act as electron acceptors (Richardson and O’Dell, 1985) and consequently, metal deficient (p-type) sulphide minerals oxidize faster. This process concurs with the accepted mechanism of sulphide mineral oxidation (Ahlberg and Elfstrom Broo, 1991) viz:

• Loss of a metal cation, and

• Increase in the oxidation state of the sulphide atoms via the formation of polysulphide on the mineral surface.

The presence of impurities in the sulphide mineral is equivalent to 'doping' the semiconductor, and again increases electrical conductivity (Harris and Richter, 1985).


It may be noted that altered copper sulphides in Ok Tedi skarns are p-type semi conductors. Their increased conductivity comes from a partial covalent linking of sulphur atoms with an ability to accept electrons (Folmer et al, 1988).

The description of n-type and p-type semiconductor sulphide minerals is also a likely explanation for the three mineral oxidation classes and three stages of oxidation of sulphide minerals. These include:

• n-type semiconductor, being a relatively pure sulphide mineral, with a long induction time prior to any appreciable oxidation rate. This criterion applies to the Class 1 classification of sulphides described in test work above.

• p- type semiconductor, being a metal deficient sulphide, or with metal impurities in the sulphide lattice. This sulphide is ready to oxidize relatively rapidly when exposed to oxygen and moisture. If this mineral type is protected from the elements prior to blasting, then it constitutes a Class 2 skarn, initially showing pristine flotation performance, but prone to oxidize rapidly on exposure to the atmosphere, particularly in a moist environment.

• p-type semiconductor sulphide mineral that has been exposed to moisture and oxygen through underground aquifers, cracks and fissures. This Class 3 ore will have already oxidized over geological time and exhibits an inferior flotation performance, even if processed immediately after blasting.

Sulphide Mineral Rest Potentials and Oxidation Rates

For a large mineral stockpile, Stockwell et al., (2003) recorded greater oxidation of marcasite relative to pyrite and chalcopyrite. In some flotation tests performed on Ok Tedi skarns, chalcopyrite flotation has shown no indication of particle surface oxidation, an observation based on a very good flotation recovery. However, pyrite and marcasite flotation was poor, suggesting oxidation of these latter minerals.

These results appear to contradict the hierarchy of oxidation rest potentials found in the literature (Majima, 1969). However, as the laws of thermodynamics are confirmed, explanation is sought in mineral associations.

Sulfide minerals oxidise via an electrochemical mechanism, with electrons passing between the two half-cell reactions of oxidation and reduction (redox). The relative electro-activity of minerals may be compared on the basis of their 'rest potentials', a mineral of higher oxidation rest potential being harder to oxidise (more 'noble') than one of lower rest potential.

Galvanic interactions occur when two (or more) electro-active species (e.g. sulfide mineral particles) come into electrical contact. This situation could occur via collision within the mineral pulp or by intimate contact along the grain boundaries of locked grains in an ore. The galvanic 'couple' will have a potential somewhere between the rest potentials of the two species, with reduction occurring on the surface of the more noble cathode, and oxidation on the more active anode. Cathodic reduction is usually a function of adsorbed oxygen.


Therefore, oxygen concentration is a contributing factor to the potential of the redox system. Anodic oxidation produces dissolution of the anodic sulfide mineral.

From Table 5, it is apparent that:

• When pyrite is in intimate contact with copper sulphides, it will drive the oxidation of the copper sulphide via a galvanic interaction mechanism.

• When pyrite is in intimate contact with marcasite, it will drive the oxidation of marcasite via galvanic interaction.

From the DDH1092 example, it is concluded that pyrite has generated a more rapid oxidation of marcasite through galvanic interaction. However, the associated pristine flotation performance of chalcopyrite observed for this case suggests that:

• The low proportion of chalcopyrite relative to pyrite in this ore sample has potentially meant that chalcopyrite was protected within the host rock from exposure to oxygen and moisture.

• There is no intimate contact between the pyrite or marcasite and the chalcopyrite in this skarn. A possible explanation for this second logical deduction comes from the altered nature of Ok Tedi skarns. Mineralogical evidence suggests mineral grains in Ok Tedi skarns commonly have an altered outer layer of rimming. A common example of rimming in sulphide skarns is the encapsulation of chalcopyrite in siderite (iron carbonate). This rimming of siderite may effectively insulate chalcopyrite from galvanic coupling with pyrite or marcasite.

CONCLUSIONS

Spatially distributed 40 tonne ROM ore stockpiles from across the Ok Tedi MLE skarn resource were obtained from pit extensions of MLE ores. Stockpiles were sampled monthly for up to 19 months to assess ore ageing effects in the flotation response of copper and pyrite.

Results were highly variable, suggesting that the 11 FS stockpiles provided individual examples of ore ageing phenomena, rather than being genuinely representative of proportionate effects.

Because there was little difference in outcomes between the finer stockpiles of the PFS and the coarse ROM stockpiles of the FS, rock size may be less important than ore predisposition to oxidation in relation to the ageing characteristics of Ok Tedi skarns.

MLE skarns exhibited three classes of aged ores in flotation, and three stages in ore ageing. It is concluded that the three classes and stages of ore ageing correlate with the semi conductor nature of sulphide minerals, being:

1. Stage 1 of ore ageing and Class 1 pristine flotation performance. This category is represented by the n-type pure sulphide mineral that requires a considerable induction time to adsorb appreciable oxygen and be transformed into the p-type semiconductor.


2. Stage 2 of ore ageing and Class 2 flotation performance. This category is represented by the p-type metal deficient sulphide mineral that is ready to oxidize rapidly but which has been protected from contact with oxygen and moisture within the host rock.

3. Stage 3 of ore ageing and Class 3 flotation performance. This category is represented by the p-type metal deficient sulphide mineral that has been exposed to oxygen and moisture within the host rock over geological time and which is subject to surface tarnishing and deposition of oxidation products on the surface of value mineral grains.

Experimental evidence indicating more rapid oxidation of marcasite and pyrite relative to chalcopyrite in some skarns is explained via:

1. The intimate association which occurs between pyrite and marcasite in Ok Tedi skarns

2. The apparent electrical dissociation between FeS2 and chalcopyrite which occurs in some skarns, possibly due to siderite rimming of chalcopyrite.

The implications of ore ageing on the flotation recovery of Ok Tedi skarns are:

• To be guaranteed of optimum flotation performance, ore processing must proceed within one month of blasting.

• If this processing time frame is not possible, then processing within six months of blasting will generally ensure protection against uneconomic flotation performance.

• Oxidation of pyrite and marcasite over geological time indicates a need for remediation of pyrite and marcasite particle surfaces via sulphidization or other means, for the effective flotation of some skarns in the pyrite flotation circuit.

REFERENCES

Ahlberg, E, and Elfstrom Broo, A, 1991. Anodic polarization of galena in relation to flotation, International Journal of Mineral Processing, 33, p135.

Eadington, P, and Prosser, A P, 1969. Oxidation of lead sulfide in aqueous solution, Transactions of the Institute of Mining and Metallurgy, 78, C75.

Folmer, J C W, Jellinek, F, and Calis, G H M, 1988. The electronic structure of pyrites, particularly CuS2 and Fe1−xCuxSe2: An XPS and Mössbauer study, Journal of Solid State Chemistry, 72(1), p137.

Harris, P J, and Richter, K, 1985. The influence of surface defect properties on the activation and natural floatability of sphalerite, in Flotation of Sulphide Minerals (ed: K S E Forssberg), Amsterdam: Elsevier.

Majima, H, 1969. How oxidation affects selective flotation of complex sulphide ores, Canadian

Metallurgical Quarterly, 8(3), p269.


Morey, M S, 2011. Mine life extension feasibility study metallurgy. Internal report, Ok Tedi Mining,

PNG.

Nowak, P, and Chmielewski, T, 1994. Surface reactivity and collectorless flotation of galena,

Fizykochemiczne Problemy Mineralurgii, 28, p21.

Richardson, P E, and Maust Jr, E E, 1976. Surface Stoichiometry of galena in aqueous electrolytes and its effect on xanthate interactions, in Flotation, A.M. Gaudin Memorial Volume (ed: M C Fuerstenau), p365 (SME-American Institute of Mining Engineers).

Richardson P E, and O'Dell, C S, 1985. Semiconducting Characteristics of Galena Electrodes: Relationship to Mineral Flotation, Journal of the Electrochemical Society: Electrochemical Science and Technology, 132(6), p1350.

Stockwell, J, Beckie, R, and Smith, L, 2003. The hydrogeochemical characterization of an unsaturated waste rock pile, Key Lake, Saskatchewan, Canada in 6th International Conference on Acid Rock Drainage, Cairns Qld.

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AUTHOR INDEX

Index Terms Links

A

Akroyd, T J 430

Allen, M 213

Arena, T 527

Aydogan, N 86

B

Bartsch, P 468

Bearman, R 66

Benzer, H 86

Bickert, G 228

Binks, M 480

Bird, A 115

Bird, M 193

Boska, D 563

Bradshaw, D 39

Briggs, M 115

Broekman, K T 100

C

Cann, N 27

Canterford, J 12

Cantrell, R 615

Card, P 1

Casey, S 27

Coleman, R 405

Collins, S 442

Connelly, D 250 536

Corder, G 264

Cordingley, G 339

Creedy, S 460

D

Dakin, P 364

David, D 552

de Waal, H 176

qIndex Terms Links


Delemontex, G 138

Dick, C 563

Drinkwater, D 39

Dundar, H 86

Dunlop, I 280

E

Ellis, S 280

Elwin, D 364

Euston, J 292

Evertsson, C M 66

F

Farrow, J 515

Fitzmaurice, C 563

Fountain, C 49

Fox, J 506

Franke, J 193

G

Grattan, L J 430

Gray, A H 138

Green, S 264

Greenhill, P G 312

Greet, C J 419

Griffiths, M 328

Grigg, N 138

H

Hall, S 468

Harbort, G 339

Hollonds, A 26

I

Innes, B 213

K

Kinal, J 419

Klepper, R 488

Koenig, R L 100

qIndex Terms Links


Kyle, G 391

L

Lane, G 49 364

Länger, B 228

Larson, M 552

Lawson, V 589

Li, M 552

Libânio, P 49

Lombardi, J 602

Lynch, A J 86

M

Matusewicz, R 460

McCarthy, P 19

McCurdie, P 488

McLean, E 374

McTiernan, J 527

Mills, R 27

Morey, M 615

Mudd, G M 391

Muhamad, N 602

Munro, P 39

Munro, S 66

Musa, F 154

N

Nairn, D 515

Nguyen, B 515

Nilsson, K 536

P

Packer, B 107

Pax, R A 163

Perkins, T 193

Phillips, M 339

Powell, M 193

R

Rantucci, D G 430

Readett, D 506

qIndex Terms Links


Reuter, M 460

Reynolds, B 442

Rinne, A 405

Rosaguti, N 213

Ross, J 27

Rule, C 176

S

Smith, H D 391

Stegink, D 515

Steward, N 328

Stewart, M 154 213

T

Thompson, A 391

Tilyard, P 39

Toor, P 193

W

Weidenbach, M 602

Weiss, G 154

Wemyss, P 480

Wheeler, J 107

Wood, J 460

Wu, J 515

X

Xu, M 589

Y

Yeomans, T 138

Young , R 26

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