Team1 Sp2010 TexLig2Methanol Anonymous-no Exec Summary

8/2/2019 Team1 Sp2010 TexLig2Methanol Anonymous-no Exec Summary

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Chemical Engineering XXX Design Project:

Texas Lignite Coal Gasification and

Methanol Production

Final Report

Submitted: XX/XX/XXXX

Team X


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Table of ContentsExecutive Summary ......................................................................................................................... 1

I. Introduction .................................................................................................................................. 2

II. Process Description ..................................................................................................................... 2

III. Process Flow Description ............................................................................................................ 2

i. Gasifier (Gasification Unit) ....................................................................................................... 2

ii. Gas Treatment Units (Pre-Rectisol, Rectisol and Claus Processes) ......................................... 3

iii. Water Gas Shift (WGS System) ............................................................................................... 5

iv. Methanol Production (MeOH Production Loop) .................................................................... 5

IV. Safety and Controls .................................................................................................................... 6

V. Sizing Calculations ....................................................................................................................... 7

VI. Utility Requirements .................................................................................................................. 8VII. Economic Analysis ..................................................................................................................... 9

i. Fixed Capital Investment and Working Capital ........................................................................ 9

ii. Annual Operating Expenses ................................................................................................... 11

iii. Revenues ............................................................................................................................... 11

iv. Discounted Cash Flow ........................................................................................................... 11

v. Profibility Analysis .................................................................................................................. 12

vi. Sensitivity Analysis ................................................................................................................ 12

VIII. Conclusions ............................................................................................................................ 13


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Executive Summary


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I. Introduction

This paper outlines a basic comprehensive design for a methanol synthesis plant using syngas

produced from Texas Lignite, a low quality, but highly reactive coal. In order to produce

commercial grade methanol, the coal must pass through several process stages to produce

syngas, remove undesirables from the process stream, optimize syngas composition, and

synthesize methanol. These processes were modeled in Aspen Plus, and the proposed design

has been subsequently characterized with respect to technical and economic considerations as

well as safety concerns.

II. Process Description

Traditional methanol synthesis involves the use of: X) a gasification unit, X) a series of gas

treatment units, X) a water-gas shift reactor, and X) a methanol production loop, which

optimizes the production of pure methanol and includes a recycle for unreacted gases.

In the gasifier, a fraction of the coal is combusted to provide the energy required for the

concurrent endothermic gasification reaction, C(s) + HXO --> CO + HX. The combustion,

gasification, and accompanying reactions in the gasifier generates raw syngas from the coal.

Ash, ammonia, and water are removed from the raw syngas prior to the Rectisol process unit

(which removes HXS and COX via physical absorption) followed by a Claus process (which

converts HXS to solid sulfur). After passing through the gas treatment units, the clean syngas

consists primarily of HX, HXO, CO, and COX. In preparation for the methanol production

reaction, a water gas shift (WGS) reactor is used to generate a X.X:X mol ratio of HX:CO. Thisslightly-above stoichiometric ratio of the reactants ensures an excess of HX, preventing the

precipitation of solid carbon on the catalyst. The process stream then enters the methanol

production loop, where the syngas is used to produce commercial grade methanol by the

reaction XHX + CO --> CHXOH. The methanol is distilled from the rest of the product stream,

and the gas stream is recycled back into the clean syngas stream to increase reaction conversion.

III. Process Flow Description

(See Appendix A.X for the Process Flow Diagram)

i. Gasifier (Gasification Unit)The gasifier is modeled in Aspen by four reactors in series. The reactions that occur in the

gasifier, as well as all other reactions, are tabulated in Appendix A.X. The reactors are modeled

at XX atm under adiabatic conditions, since running the reactors at an elevated pressure


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decreases the size of the gasifier, the number of filtration units needed to remove impurities in

the syngas, and the need for compressors later on in the production process .

In the first stoichiometric reactor (INTRXNS) the hydrogen, nitrogen and sulfur compounds

present in coal are transformed into HXO, HXS and NHX. A heat correction of XXX kcal per kg ofcoal is applied to the unit in order to account for the difference in the enthalpies between a

mixture of pure elements present in coal and the actual compounds present in coals complex

chemical structure. The second stoichiometric reactor (COMBUST) models the complete

combustion of hydrogen into HXO and the fractional combustion, approximately XX%, of carbon

into COX. This design currently operates under slagging conditions, and uses a feed of pure OX.

The target syngas temperature of XXXXXC is achieved using a design specification varying the

feed of OX. Compression of OX for the gasifier occurs in multiple stages (with cooling steps in

between) in order to maintain reasonable pressure ratios. The third stoichiometric reactor

(GASIF) uses a stoichiometric reactor to model gasification of the remaining carbon. XXXXX

kmol/hr of steam is fed into the gasifier as a reactant, and the carbon is assumed to react

completely. The final reactor in the series, an equilibrium reactor (WGSGASIF) represents the

WGS reaction that naturally occurs within the gasification unit.

ii. Gas Treatment Units (Pre-Rectisol, Rectisol and Claus Processes)At this stage in the process, the exit stream contains HX, CO, COX, HXO, and small amounts of HXS,

OX, and NHX. The process stream must be cooled to a low temperature in order to remove the

undesirable acid gases present. A combination of two heat exchangers (one producing steam,

one using cooling water) and a chiller is used to decrease the syngas temperature to XoC before

entering the acid gas removal process.


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Pre-Rectisol Scrubber

It is imperative to remove the ammonia and water from the syngas prior to treatment with

Rectisol, because ammonia builds up as ammonium carbamate in methanol, while water is a

solid at the low temperatures required for the Rectisol process. To separate the ammonia andwater from the process stream, the syngas is sent through a water scrubber (AMMSCRUBB),

modeled as a RadFrac column with XX stages.

Rectisol Process

Prior to entering the Rectisol process, the dry syngas is cooled to -XX oC by means of a multi-

stream heat exchanger, which transfers heat between the room temperature makeup methanol

and dry syngas streams and the cold acid gas and clean syngas streams. This is done in order to

obtain the low-temperature, high-pressure operating conditions necessary for optimal

absorption of HXS and COX in methanol in the absorption column (RECTSOLX), modeled as a

RadFrac column with XX stages. Recycle and makeup streams of methanol are fed into

RECTSOLX as the liquid solvent stream, countercurrent to the dry syngas. HXS and COX are

physically absorbed into the liquid methanol stream in RECTSOLX at lower temperature and

higher pressure (-XXoC and XX atm), and are then removed in the stripping column (RECTSOLX),

also modeled as a RadFrac column (but with XX stages), at higher temperature and lower

pressure (XXoC and X atm). Through the use of a reboiler and partial condenser in RECTSOLX,

methanol functions as both the liquid and vapor streams, so that no separate inert carrier gas is

needed to strip HXS and COX from the liquid methanol. Because there is no accumulation of

inert species in the methanol, no purge is used in this design.

EPA regulations cap SOX emissions, and assuming a X:X conversion ratio of HXS to SOX, the two

major limitations are: X) less than X.X% sulfur in the fuel by weight, and X) less than X.X lbs SOX

per MMBtu of energy generated from the fuel. Therefore, the clean syngas leaving the Rectisol

process must fulfill these two requirements. The first requirement is easily fulfilled, since there

is only approximately X.XXXX% sulfur present in the clean syngas stream. The second

requirement is also fulfilled, since modeling a simple turbine that generates power from the

syngas shows that only X.XXX lbs SOX would be generated per MMBtu of theoretical energy

produced (using a XX:X pressure ratio in the turbine).


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Claus Process

The Claus process is modeled using a stoichiometric reactor (CLBURNER) and a Gibbs reactor

(CLSRCT) in series. Much of the modeling is accomplished with rough approximations, since in

reality the entire Claus process unit will be purchased externally according to the amount ofsulfur it is expected to produce. In CLBURNER, one third of the HXS from the Rectisol process is

combusted with a stoichiometric amount of OX to form HXO and SOX. HX, CO, and methanol

present in the syngas will also combust to form HXO and COX. After removing excess water by

means of a flash tank (CLFLASH) to increase conversion, the remaining HXS is reacted in CLSRCT

with the SOX generated in CLBURNER to form HXO and sulfur. To be more accurate, as a Gibbs

reactor, CLSRCT determines the most energetically favorable compositions of all species present

at equilibrium under the operating conditions, which are XXXoC and X atm. Sulfur, which is

primarily present in the liquid phase as SX, can then be separated (modeled as an ideal separator,

SULFSEP) from the tailgas. The expected sulfur production rate of this design is X.X tonnes/day.

iii. Water Gas Shift (WGS System)A splitter sends a fraction of the clean syngas exiting the Rectisol process into the jacketed plug

flow water gas shift reactor, WGSINDEP, along with XXXX kmol/hr of input steam. The split ratio

was adjusted using a design specification that held the composition of the exiting syngas stream

to be the X.X:X HX:CO ratio desired for methanol production. The simulated reaction kinetics

can be found in Appendix A.X. The WGS reaction is run at high pressure, so in order to obtain a

reasonable reactor volume (see Appendix A.X.X), the catalyst should be diluted by a factor of

XXXX times to keep superficial gas velocity manageable. This was modeled in Aspen by

proportionally decreasing the pre-exponential factor.

Because the WGS reaction is exothermic, this design uses a heat exchange jacket with co-

current coolant. Most of the conversion occurs near the beginning of the reactor, so a co-

current heat exchange causes the lowest temperature coolant to contact the hottest part of the

reactor, thus maintaining a steadier temperature profile within the reactor.

iv. Methanol Production (MeOH Production Loop)The reaction kinetics used in the methanol synthesis reactor are shown in Appendix A.X.

Methanol production occurs within a recycle loop to increase conversion. The methanol

synthesis heat-jacketed plug flow reactor, METHRXN, is operated at XX atm (a small pressure


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drop in the process has been accounted for) and an inlet temperature of XXX oC, which is higher

than the rest of the recycle loop. In order to minimize the outside heating and cooling

necessary, the reactor inlet and outlet streams exchange heat, before the reaction product

stream is further cooled with cooling water, and chilled to XoC. The synthesized methanol

product is then separated through a series of two flash drums with a final yield of XXXX

tonne/day of methanol predicted.

IV. Safety and Controls

Safety precautions and protocols are critical to a comprehensive plant design. Plant safety

should be integrated at all levels within a plant, from the engineering controls to the

administrative procedures. All potential hazards should be carefully evaluated and, if possible,

preemptively addressed at the design and engineering stage. However, it is still imperative that

the plant workers are aware of the risks around them, and follow all safety procedures and

protocols set in place.

Process units should be interlocked so that upon detection of the failure of one unit, other

processes will automatically be paused or shut off. Hazards should be isolated as much as

possible, so that in the case of a fire or explosion in one part of the plant, the collateral damage

can be minimized. One example of this is the installation of firebreaks, especially in the

gasification and methanol production units.

One simple precaution that should be taken is establishing adequate checkpoints throughout

the entire process for temperature, pressure, flow rate, etc. The more information that is

monitored, or at least available, the more likely it is to locate a failure when it occurs. For

example, in this design, the WGS catalyst is diluted by XXXX times, which constitutes a potential

hot spots risk in the reactor because the dilute catalyst may become unevenly distributed along

the reactor length. Closely monitoring the temperature profile along the reactor would

decrease the severity of this risk, since quick action can be taken if the temperatures fluctuate

too severely. As another example, because the coal gasification is run under slagging conditions,

it is important to maintain a consistently high reactor temperature to ensure full slagging.

Incomplete slagging may end up damaging downstream equipment and cause unexpected

equipment failure.


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This particular process runs at high pressures throughout, and thus runs the risk of pressure

buildup. Safety measures to minimize this risk include constant monitoring of pressure at all

points in the system, and the installation of safety valves, pressure relief valves, and explosion

panels. Danger to personnel can be reduced by overseeing the workfloor with cameras in place

of direct observation, although safety glass can be used when remote observation is infeasible.

Regular plant maintenance and housekeeping is also very important in reducing safety risks.

Through regular equipment checks and replacement of old seals and gaskets, the potential of

flammable syngas or methanol leaking can be minimized. Processes involving volatile hazardous

compounds such as methanol should be carried out in areas with good ventilation, in case of

leaks, and in a cool location. Easily ignitable materials such as coal and methanol should be

properly stored away from any potential sources of ignition. Long-term or large-quantity

storage of hazardous materials should be kept to a minimum.

Solid inerts accumulation is also a consideration. Deionized water can be used in heat

exchanger cooling water pipes to prevent calcification buildup, which decreases heat transfer

between the water and the hot stream while increasing the pressure in the pipe, potentially

resulting in pipe rupture. Additionally, ammonia reacts with methanol in the Rectisol process to

produce ammonium carbamates, and though the ammonia is removed beforehand, there may

still be non-negligible accumulation of inerts; the same concept applies to ice buildup in Rectisol.

V. Sizing Calculations

(All equipment dimensions and sample calculations can be found in Appendix A.X.)

An initial estimate of gasifier size can be performed based on previous empirical values for a

similar unit [X], with the approximation that the space velocity remains the same. Given the

coal feed rate and gasifier dimensions of the empirical case, and the coal feed rate of this design,

the gasifier volume for this design was determined to be XX mX.

The water gas shift reactor size scale is approximated by assuming a gas hourly space velocity

(GHSV) found in a previous study [X]. With a reasonable GHSV, and the volumetric flow rate

entering the WGS reactor, a preliminary reactor volume estimate was determined to be X.X mX.

Because the dimensions of the WGS reactor strongly affect the reactor temperature and


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conversion, the modeled reactor dimensions were then tuned to obtain optimal profiles,

resulting in a final reactor volume of X mX.

The methanol reactor dimensions were chosen in Aspen to obtain a desirable conversion profile.

The number of tubes in the reactor was selected in order to obtain a desirable temperatureprofile. The reactor is composed of XX tubes, each with a diameter of XX cm and a length of

XXXX m. Absorption, stripping and distillation columns are sized using a correlation relating the

flooding limit to tray spacing, shown in Appendix A.X.X. [X] Since the flooding limit is the

maximum superficial gas velocity allowable in the column, the volumetric flow rate of gas

determines the diameter of the column. Column height is then calculated using tray spacing

and number of stages. Size dimensions of all columns can be found in Appendix A.X.X.

Heat transfer area is calculated by setting the heat gained by the cold stream (or the heat lost bythe hot stream) equivalent to the heat transferred across the hot and cold streams. The

temperature gradient in each of the single streams is the difference between inlet and outlet

temperatures, while the temperature gradient across the streams was taken to be a log mean

average of the single stream temperature gradients. The same calculation is used for multi-heat

exchangers, for which heat from each stream may be transferred across multiple streams. Area

requirements for all heat exchangers can be found in Appendix A.X.X.

VI. Utility Requirements

To deal with most of the plants heating and cooling requirements, steam and cooling water

heat exchangers are used, respectively. In this way, the need for refrigeration systems can be

greatly reduced, and the need for electric heaters is fully eliminated. Most of the high-pressure

steam in this design is generated at the gasifier outlet, where pressurized water at XX atm and

XX C exchanges heat with the hot process stream at XXXX C, generating over XXXX tonnes/day

of steam at XX atm. This steam is then used to feed the gasifier and WGS reactions, as well as to

heat the Rectisol methanol stream (RICHMETH) and the pre-WGS process stream (WGSX).

Additionally, low pressure steam at X atm is also produced from the reactor cooling-jackets,

some of which is used to heat the methanol leaving the Methanol Production Loop

(METHCOND). For cooling purposes, water at X atm and XX C exchanges heat with hot process


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streams. As a result, the water is heated to approximately XX C and then cycled to a cooling

tower, where it is cooled back to room temperature via evaporation at a rate of XXXX tonne/day.

VII. Economic Analysis

The fixed capital investment and working capital are calculated; the sum of these expenses

represent the costs needed to build the plant and bring it to an operational state. Yearly

expenses and revenues were estimated be extrapolating current market values for utilities, coal,

and sulfur. The potential profitability of this design is then considered in the discounted cash

flow analysis. Note that costs are estimates and the overall analysis accuracy is within XX%. The

majority of costing is based on Towler, a XXXX edition with a comprehensive approach to overall

plant costing. [XX]

i. Fixed Capital Investment and Working Capital

Fixed Capital Investment

Fixed capital investment represents the total cost of designing, constructing, purchasing all

materials, and installing a plant. The economic analysis of this project predicts a $X,XXX million

($XXXXM) fixed capital investment, which is obtained by adding the costs of its four components:

ISBL investment, OSBL investment, engineering and construction costs, and contingency charges.

Inside Battery Limits (ISBL) Investment

ISBL investment represents the cost of the plant itself, and includes (i) the major process

equipment, (ii) bulk items such as piping, valves, wiring, instruments, insulation, catalysts, etc,

(iii) civil works such as roads, foundations, buildings, sewers, etc, and (iv) installation labor and

supervision. It also includes indirect costs such as construction costs, field expenses,

construction insurance, labor benefits and burdens, and other miscellaneous costs. The

individual equipment costs can be estimated by correlations in Towler. Hand factors are then be

used to approximate the total cost including costs associated with installation. The total ISBL

costs sum to $XXXM (the depreciable portion of which is $XXXM) and the details of the

calculations appear in Appendix B.X.X.


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Outside Battery Limits (OSBL) Investment

OSBL costs include all offsite expenditures such as electrical support, cooling towers, circulation

pumps, water treatment, sewers, feed and product pipelines, loading facilities, warehouse,

railroads, laboratories and analytical equipment, offices, workshops and maintenance facilities,site security, fencing, landscaping, etc. These costs are estimated as a XX% of ISBL costs for this

initial economic analysis, as is consistent in Towler. The resultant estimate is $XXXM, as detailed

in Appendix B.X.X.

Engineering and Construction Costs

Engineering and construction costs include design of process equipment, piping systems, control

systems, plant layout, civil engineering, procurement of plant items, construction and

supervision services, administration, contractors profit, etc. Engineering costs are estimated as

XX% for small projects and XX% of ISBL plus OSBL costs in accordance of Towlers estimate for

larger projects; we chose XX% due to the large scale design of this plant. This works out to

$XXM as detailed in Appendix B.X.X.

Contingency Changes

Contingency charges account for extra costs added to the project budget to allow for variation

from the cost estimate. They account for changes in prices, project scope, currency fluctuations,

and other unexpected events. This analysis estimates the contingency charges to be XX% of ISBL

plus OSBL costs - a value reasonable for a process using established technologies. Appendix

B.X.X shows that the total for this category is $XXXM.

Working Capital

In addition to the fixed capital investment charges above, working capital- additional money

needed to start up the plant and run it until it begins earning income -- is also considered. This

includes the following:


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Item Costing Criteria Value

Cash on hand estimated as X weeks worth of

production

$XX.XM

Two weeks worth of rawmaterials

- $X.XM

spare parts inventory estimated as X% of ISBL + OSBL costs $X.XM

TOTAL $XX.XM

It can be seen that the total total, the fixed capital and working capital sum to $X,XXXM.

ii. Annual Operating Expenses

In addition to the initial money needed to start a plant, there are two categories of annual

operating expenses: (i) variable costs of production and (ii) fixed production costs.

Variable costs of production scale with plant output. They include raw materials, utilities,

consumables, and waste disposal. As Appendix B.X shows, the costs total to $XX.XM/year.

Fixed production costs do not scale with plant output. These costs include operating labor,

supervision, direct salary overhead, maintenance, property taxes, land rent, general plant

overhead, etc. These costs are detailed in appendix B.X, and sum to $XX.XM.

iii. Revenues

Sales of methanol and sulfur provide revenues for the plant. The Aspen simulation modeling the

plant predicts plant outputs of XXXX tonnes/day of methanol and X.X tonnes/day of solid sulfur.

As Appendix B.X outlines, the selling price of these commodities are $XXX/ton and $XXX/ton,

respectively. Methanol and sulfur annual profits total to $XXXM/year and $XXX,XXX/year,

respectively, resulting in total revenues of $XXXM/year.

iv. Discounted Cash FlowDiscounted cash flow diagrams are calculated to illustrate the rate of return on an initial

investment. The relevant equations used are:


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Most years, this sums to $XXX.XM, however, every Xth year (starting year X) catalyst is purchased

and an additional $XXXM is added to the operating costs. Gross Profit compares these

operating costs to the income and depreciation:

Depreciation is calculated with double declining balance until the linear depreciation values are

greater. The tax rate is then applied and the net product is calculated:

Where the overall tax rate is estimated to be XX% -- this includes federal, state, and local taxes.

The addition to cash flow (the desired annual number) is simply the sum of the depreciation and

the net profit. There is the option of multiplying this number by a discounting factor, where the discount rate is fractional (eg: XX% discount rate = X.XX) and n isthe number of years passed since year X.

For a X% discount rate, the profit over XX years is $X,XXXM. Note that this is not corrected for

the time value of money. By solving for the discount rate that gives X net cash flow at year XX,

the annual rate of return is identified. Applying this method, the rate of return is determined to

be XX.X%. The cash flow diagrams and plots are shown on the proceeding pages.

v. Profibility Analysis

A large initial capital investment ($X,XXXM) is required to cover the costs of building the plant

outlined in the design. The economic analysis, although only accurate within XX% suggests that

over a period of thirty years, this investment will generate substantial returns.

vi. Sensitivity Analysis

As electricity is a major cost in operating this plant, one of the primary considerations in

optimizing the economic viability of this design is the number and type of electricity-consuming

equipment utilized. The three operations that require electricity are: heating and cooling,

pressurization, and distillation. Through the strategic implementation of heat exchangers, the

number of electrically-powered heaters and coolers needed was reduced. Additionally, the


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process flow was modified (when possible) to pressurize liquids instead of gases, since liquids

are relatively incompressible, and thus require less energy. Finally, addition of a flash tank

before the Methanol Production Loop eliminated the need for an energy-intensive distillation

column at the end of the process. By making these changes, the total consumption of electric

power was reduced by more than XX%, from over XXX MW to fewer than XX MW.

Another aspect of the process design that is extremely susceptible to changes is the

temperature profile along the reactor in the methanol production loop, as shown in Appendix

A.X. With minor changes in the heat jacket coolant flow rate (around X.X%), the equilibrium

conversion temperature and length shift much more significantly (around XX%). Changes in the

inlet temperature rate (around X.X%, or XXoC) lead to a surprisingly large shift in equilibrium

conversion and length (around XX-XX%). Because the methanol reactor conversion equilibrium

has a direct effect on revenue, it is important to note that minor changes in operating conditions

may have a significant effect on profitability.

VIII. Conclusions

This project is an introductory survey into the feasibility of constructing a proposed methanol

plant. This plant design provides an estimated XX% return on investment for the production of

methanol from Texas Lignite coal. The process makes use of coal gasification, rectisol, water-gas

shift, and methanol synthesis technology to yield a reasonable conversion of coal to methanol.

Many factors have been taken into account and assessed in this report, including technical

considerations, science and engineering fundamentals, safety and controls, and economic

principles. This introductory investigation provides the groundwork for future analyses.


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References:

[1] Bockelie, M.J. et al. CFD Modeling For Entrained Flow Gasfiers. Neville Holt, Electric Power

Research Institute, Palo Alto, CA.

[2] Catalyst information. UC Berkeley Chemical Engineering XXX Resource, Spring XXXX,

Microsoft Word Document (accessed Apr. XX, XXXX).

[3] Choi, Y.; Stenger, H.G. Water gas shift reaction kinetics and reactor modeling for fuel cell

grade hydrogen.Journal of Power Sources. [Online] XXXX, XXX, XXX-XXX.

http://www.sciencedirect.com. (accessed Apr. XX, XXXX).

[4] Crawford, J.; Ellifritz, B.; Root, B. Process Design of an Anhydrous Ammonia Production

Facility for Dyno Nobel. UC Berkeley Chemical Engineering XXX Resource, Spring XXXX,

PDF Document (accessed Apr. XX, XXXX).

[5] Economics cheat sheet. UC Berkeley Chemical Engineering XXX Resource, Spring XXXX,

Microsoft Word Document (accessed Apr. XX, XXXX).

[6] Gasification Plant Cost and Performance Optimization; Vol. X.; DE-ACXX-XXFTXXXXX; Bechtel,

Global Energy, and Nexant: XXXX.

[7] GKT Project. Methanol Synthesis Reactor/Kinetic Information. XXXX. UC Berkeley Chemical

Engineering XXX Resource, Spring XXXX, PDF Document (accessed Apr. XX, XXXX).

[8] Methanex. Methanol Price. http://www.methanex.com/products/methanolprice.html

(accessed Apr. XX, XXXX).


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[9] Perry, R.H.; Green, D.W. Process Safety. In Perrys Chemical Engineers Handbook; Xth Ed.

McGraw-Hill: XXXX.

[10] Peters, M.S.; Timmerhaus, K.D.; West, R.E. Plant Design and Economics for Chemical

Engineers, Xth Ed.; McGraw-Hill: New York, XXXX.

[11] Towler, G.; Sinnott, R. Chemical Engineering Design. Elsevier, Inc.: XXXX.

[12] U.S. Energy Information Administration. Coal Explained: Coal Prices and Outlook.

http://tonto.eia.doe.gov/energyexplained/index.cfm?page=coal_prices (accessed Apr.

XX, XXXX).

[13] U.S. Energy Information Administration. Electricity: U.S. Data. http://www.eia.doe.gov/

fuelelectric.html (accessed Apr. XX, XXXX).

[14] Usual Weekly Earnings of Wage and Salary Workers First Quarter XXXX. Bureau of Labor

Statistics, U.S. Department of Labor; USDL-XX-XXXX; April XX, XXXX (accessed Apr. XX,

XXXX).


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Appendix A: Table of ContentsA.X Process Flow Diagram 17

A.X Materials Flow Information 17

A.X.X Overall Balance Key Summary 17

A.X.X and A.X.X Mass and Molar Flow Sheets 17

A.X Reactor Specifications 17

A.X Reaction Kinetics 18

A.X Equipment Sizing Summaries and Calculations 18

A.X.X Reactor Sizing 18

A.X.X.X Gasifier (INTRXNS, COMBUST, GASIF, WGSGASIF) 18

A.X.X.X WGS Reactor (WGSINDEP) 19

A.X.X.X Methanol Reactor (METHRXNX) 19A.X.X Column Sizing 19

A.X.X Heat Exchanger Sizing 20

A.X Power Consumption 21

A.X.X Compressors, Pumps, Columns 21

A.X.X Coolers 22

A.X Water and Steam Balance 22

A.X Power Consumption Sensitivity Analysis 22


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A.X Process Flow Diagram(see attached pages)

A.X Materials Flow Information

A.X.X

Overall Balance Key Summary

Key Overall Materials

Input

(tonne/day) Output (tonne/day)

Coal XXXX -

Steam XXXX XXXX

OX XXXX -

CHXOH X.XX XXXX

Sulfur - X.X

A.X.X and A.X.X Mass and Molar Flow Sheets(see attached pages)

A.X Reactor Specifications

Reactor Reactions specified Extent of Reaction

Property

Method

INTRXNS .X OX + HX HXO XXX% Frac. Conv. of OX SRK

S(Cisolid) + HX HXS XXX% Frac. Conv. of S

NX + X HX X NHX XXX% Frac. Conv. of NX

COMBUST HX + .X OX HXO XXX% Frac. Conv. of HX SRK

C(Cisolid) + OX COX XXX% Frac. Conv. of OX

GASIFER C(Cisolid) + HXO CO + HX XXX% Frac. Conv. of C SRK

WGSGASIF CO + HXO HX + COX Equilibrium conditions SRK

CLBURNER X HXS + X OX X HXO + X SOX XX% Frac. Conv. of HXS SRK

X HX + OX X HXO XXX% Frac. Conv. of HX

X CO + OX X COX XXX% Frac. Conv. of CO

X CHXOH + X OX X COX + X HXO XXX% Frac. Conv. of CHXOH

CLSRCT

(specified by thermodynamics) Operating conditions

@ XXX

o

C, X atm

SR-POLAR

WGSINDEP CO + HXO HX + COX Kinetics-based SRK

HX + COX CO + HXO Kinetics-based

METHRXNXCHXOH CO + X HX Kinetics-based SRK

CO + X HX CHXOH Kinetics-based


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A.X Reaction Kinetics

WGS (WGSINDEP) [X,X]

Forward: CO + HXO HX + COX

Af1 = X.Xe-XX

Ef= XXXXX Reverse: HX + COX CO + HXO

AbX = X.XXe-XX

Eb = XXXXX

Methanol (METHRXNX) [X]

Forward: CO + X HX CHXOH

Af= X.XXXE-XX

Ef= XX,XXX Reverse: CHXOH --> X HX + CO

Ab= X.XXXXXEX

Eb= XXX,XXX

A.X Equipment Sizing Summaries and Calculations

A.X.X Reactor Sizing

Reactors Diameter (m) Length (m) # of tubes Sizing Method

GASIFIER2 X.X X.X N/A Holt Neville

WGSHIFT X.XX XX XX Choi & Stenger [X]

METHRXNX XX XXXX XX -

A.X.X.X Gasifier (INTRXNS, COMBUST, GASIF, WGSGASIF) Since and are constant for reactors of a similar type (slagging entrained bed) andscale (a few thousand tonnes of coal per day), it is evident that reactor volume scaleswith feed rate of coal. Also, the height-to-diameter ratio is assumed to be constant.

() ()

Thus, given certain values for reactor dimensions and flow rate in the Holt Neville study

[X], the following values are obtained:

1Pre-exponential factor scaled down by XXXXx because of catalyst dilution

2 encompasses INTRXNS, COMBUST, GASIFIER, WGS


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Holt Neville This design

Diameter (m) X.X X.X

Length (m) X.X XX.X

Volume (mX) XX.X XX.X

Flow (tonne/day) XXXX XXXX

A.X.X.X WGS Reactor (WGSINDEP)Reasonable WGS reactor gas hourly space velocities (GHSV) noted in a previous study were in

the range of XXXX to XXXXX hr-X. A GHSV of XXXX hr-X was chosen for the sizing purposes of this

process.

Exact diameter and length values were obtained by optimizing the temperature and conversion

profiles in Aspen.

A.X.X.X Methanol Reactor (METHRXNX)

The methanol reactor was sized solely in Aspen in order to provide the optimal temperature and

conversion profile.

A.X.X

Column SizingLength (m) Diameter (m) # of trays Tray spacing (in)

AMMSCRUB XXX.X X.X XX XX.X

RECTSOLX XX.X X.X XX XX

RECTSOLX XXX.X X.X XX XX

The columns in this design are sized by correlating tray spacing, number of trays, and flooding

velocity with column diameter:

( )


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where is the gas velocity (in ) at the flooding limit, is a capacity parameter (in )correcting for surface tension (value obtained from a chart, and is dependent on flow parameterand tray spacing), is the surface tension (in ) of the liquid phase, and , are densities (in) of the liquid and gas phases.A numerical example of sizing is shown below for RECTSOLX:

( ) ( )

() 3

( )

A.X.X Heat Exchanger Sizing

HTX Unit Heat Duty (MW) Area (mX

)HTXCMPX X.X XX.X

HTXCMPX X.X XX.X

HTXSYNGX XXX.X XXXX.X

HTXSYNGX XX.X XXX.X

HTXR (MHTX) X.X -

HTXR - X X.X X.X

HTXR - X X.X X.X

HTXR - X X.X XXXX.X

HTXRECT XX.X XX.X

HTXWGS X.X XX.X

HTXWGSX XX.X XXXX.X

HTXM XXX.X XXXX.X

3From chart, Perrys Handbook X

thed. Fig. XX-XX


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HTXMETHX XX.X XXXX.X

HTXMETHX X.X XX.X

Sizing of a heat exchanger involves determining the area of heat transfer. The equations

required are shown below:

Sizing of a heat exchanger, HTXCMPX, is shown below as a numerical example:

A.X

Power Consumption

A.X.X Compressors, Pumps, Columns

Compressors, Pumps and Columns

Type Unit

Work Required

(MW)

Compressor

OXCOMPX X.X

OXCOMPX X.X

OXCOMPX X.X

Pump

WATPUMPX X.XX

AMMPUMP X.XXMEPUMP Xe-X

RECPUMP X.XX

Distillation RECTSOLX XX.X

Net Work XX.X


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A.X.X Coolers

Coolers

Unit

Heat Duty

(MW)

Work Required

(MW)

GASCHILL -XX.X X.XX

AMMCOOL -X.XX X.XX

RECTCOOL -XX.X XX.X

METHREFR -XX.X XX.X

Net Work XX.X

A.X Water and Steam Balance(see attached pages)

A.X Power Consumption Sensitivity Analysis (see attached pages)

HIGH PRESSURE STEAM (XXX C, XX ATM)

PRODUCED USED

Stream HPSTMX HPSTMX HPSTMINX HPSTMINX HPSTMINX HPSTMINX

Phase VAPOR VAPOR VAPOR VAPOR VAPOR VAPOR

Substream: MIXED

Mass Flow tonne/day

HXO XXXX.X XXX.X XXX.X XXX.X XXXX.X XXX.X

Temperature C XXX.X XXX.X XXX.X XXX.X XXX.X XXX.X

Pressure atm XX.X XX.X XX.X XX.X XX.X XX.X

Vapor Frac X.X X.X X.X X.X X.X X.X

Liquid Frac X.X X.X X.X X.X X.X X.X

Total prod. @ XXX C (tonne/day) XXXX.X

Total used @ XXX C XXXX.X

Total prod. @ XXX C (tonne/day) XXX.X

Total used @ XXX C X.X

LOW PRESSURE STEAM (X ATM)

PRODUCED USED

Stream LPSTMX LPSTMX LPSTMINX

Phase VAPOR VAPOR VAPOR

Substream: MIXED

Mass Flow tonne/day

HXO XXX.X XXXX.X XXXX.X

Temperature C XXX.X XXX.X XXX.X

Pressure atm X.X X.X X.X

Vapor Frac X.X X.X X.X

Liquid Frac X.X X.X X.X

Total prod. @ XXX C (tonne/day) XXX.X

Total used @ XXX C X.X

Total prod. @ XXX C XXXX.X

Total used @ XXX XXXX.X

HOT WATER (XX C, X ATM)


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PRODUCED

Stream HOTWATX HOTWATX HOTWATX HOTWATX HOTWATX

Phase LIQUID LIQUID LIQUID LIQUID LIQUID

Substream: MIXED

Mass Flow tonne/day

HXO XXXX.X XXXX.X XXXXX.X XXXXX.X XXXXX.X

Temperature C XX.X XX.X XX.X XX.X XX.X

Pressure atm X.X X.X X.X X.X X.XVapor Frac X.X X.X X.X X.X X.X

Liquid Frac X.X X.X X.X X.X X.X

Total prod. @ XX C (tonne/day) XXXXXX.X

Total used @ XX C X.X

COOLING WATER

USED

Stream COOLWATX COOLWATX COOLWATX COOLWATX COOLWATX COOLWATX COOLWATX COOLWATX

Phase LIQUID LIQUID LIQUID LIQUID LIQUID LIQUID LIQUID LIQUID

Substream: MIXED

Mass Flow tonne/day

HXO XXXX.X XXXX.X XXXX.X XXXXX.X XXX.X XXXXX.X XXXX.X XXXXX.X

Temperature C XX.X XX.X XX.X XX.X XX.X XX.X XX.X XX.XPressure atm X.X X.X X.X X.X X.X X.X X.X X.X

Vapor Frac X.X X.X X.X X.X X.X X.X X.X X.X

Liquid Frac X.X X.X X.X X.X X.X X.X X.X X.X

Total used (tonne/day) XXXXXX.X


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Appendix B: Table of ContentsB.X General Notes 25

B.X Fixed Capital Costs 25

B.X.X ISBL Costs 25

B.X.X OSBL Costs 28

B.X.X Engineering Costs 29

B.X.X Contingencies 29

B.X Working Capital 29

B.X Variable Costs of Production 30

B.X.X Utilities 30

B.X.X Consumables 31

B.X Fixed Costs of Production 32B.X Revenues 33


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B.X General NotesCosting is based on XXX days/year of production, where the operational days run XX/X. Costing

estimates have an overall accuracy within XX%.

B.X

Fixed Capital Costs

B.X.X ISBL Costs

Equipment Costs4

To calculate the gasifier cost, the hypothetical power output of this Texas Lignite coal at the

specified flow rate is calculated. This power is multiplied by XX%. This coal contains X.XX Btu/lb;

by multiplying by the mass flow rate and X.X, the energy output is determined to be XXXMW.

This is very close to the XXX MW gasifier listed in Bechtel [X]; hence, we approximate the design

gasifier to have the cost cited in the paper, namely $XXXM.

Other major pieces of equipment were sized according to the correlations in Towler[XX]. This

method uses a sizing parameter S and the constants a, b, and n in the following function:

Item a b n Sizing

Parameter, S

Acceptable Range

for S

heat exchanger

(U-tube shell and tube)

XX,XXX XX X area, mX XX-XX,XXX

jacketed reactor, agitated XX,XXX XX,XXX X.X volume, mX X.X-XXX

compressor (centrifugal) X,XXX XXXX X.X driver power,

kW

XXX-XX,XXX

pump

(single stage centrifugal)

X,XXX XX X.X flow, Liters/s X.X-XXX

pump driver

(electric motor, explosion

XXX XXX X.X power, kW X-X,XXX

4Because the book used was a XXXX edition, adjusting the cost to XXXX values using cost indices was

deemed unnecessary


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proof)

distillation column (stainless

steel, XXXss)

-

XX,XXX

XXX X.X shell mass, kg XX-XXX,XXX

sieve trays XXX XXX X diameter, m X.X-X

cooling tower XX,XXX XXX X.X flow, Liters/s XXX-XX,XXX

ISBL COSTS

The total ISBL cost is estimated by the Hand method, which proposes that the total ISBL cost is

given by where Fi and Ci are the Hand factor and delivered cost of the givenpiece of equipment. The Hand factor used for each type of equipment is tabulated below and a

sample calculation is provided.

Equipment Type Hand

Factor

compressors X.X

distillation columns X

fired heaters X

heat exchangers X.X

miscellaneous equipment X.X

pressure vessels X

The following are the equipment costs and installation costs:

ISBL COSTS:GASIFIER

$XXX,XXX,XXX

$ installation costs

JACKETED REACTOR

$XX,XXX WGS

$X,XXX,XXX Methanol

$XX,XXX,XXX


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Columns & Trays

$XXX,XXX RECTSOLX column

$XX,XXX all trays

$XX,XXXRECTSOLX

column

$X,XXX all trays$XXX,XXX

AMMSCRUBcolumn

$XX,XXX all trays

$X,XXX,XXX installation costs (for all columns and trays)

COMPRESSORS

$XX,XXX OXCOMPX

$XX,XXX OXCOMPX

$XX,XXX OXCOMPX

$XX,XXX installation costs (for all columns and trays)

HEAT EXCHANGERS

Cost: Heat Exchanger

$XX,XXX.XX HTXCMPX

$XX,XXX.XX HTXCMPX

$XXX,XXX HTXSYNGX

$XX,XXX HTXSYNGX

$XX,XXX HTXR - X

$XX,XXX HTXR - X

$XXX,XXX HTXR - X

$XX,XXX.XX HTXRECT

$XX,XXX.XX HTXWGS

$XXX,XXX.XX HTXWGSX

$XXX,XXX.XX HTXM

$XXX,XXX.XX HTXMETHX

$XX,XXX.XX HTXMETHX

X.XXE+XX installation costs (for all heat exchangers)

PUMPS

Cost: Pump Name

$XXX,XXX WATPUMPX

$XX,XXX AMMPUMP

$X,XXX MEPUMP

$XXX,XXX RECPUMP

$X,XXX,XXX installation costs (for all)

Cooling Tower

$XXX,XXX

XXXXXX.XXXX installation costs


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PUMP DRIVERS (electric motors)

$XX,XXX.XX WATPUMPX

$X,XXX.XX AMMPUMP

$XXX.XX MEPUMP

$XX,XXX.XX RECPUMP

$XXX,XXX installation costs (for all)

CLAUS PROCESS

$XXX,XXX

Land

$XX,XXX,XXX

Subtotal:

ISBL COSTS:

includes: equipment erection, piping, instrumentation, control, electrical, civil, structures &

buildings, lagging and paint

$XXX,XXX,XXX

depreciation amount: $XXX,XXX,XXX doesnt include installations

In addition to the costs listed above, land costs and Claus System costs were added.

Investigations of land cost suggested a reasonable purchase price of $XXM5. The Claus system

cost was based on [X]. This reference states that a XXX tons/day sulfur treatment plant costs X

million Euros. Using a linear extrapolation with an intercept of zero, and a conversion rate of

$X.X per Euro, the cost is determined to be $XXX,XXX. The adjustment from the XXXX value

listed in the paper to the XXXX value desired was adjusted by the Chemical EngineeringMagazine cost indices; the ratio of XXXX to XXXX was X.XX resulting in the adjusted Claus unit

cost of:

The total ISBL costs are used in subsequent calculations, so we note here that it sums to

$X,XXXM

B.X.X OSBL CostsTowler suggests the OSBL costs generally range from XX to XX% of the ISBL costs, with XX%

being a good estimate for processes where nothing is known. Because our primary OSBL cost isthe cooling tower, which is accounted for in ISBL in this case, we selected a relatively low value

of XX%. Thus, the OSBL costs are estimated at $XXXM.

5For approximately XX acres outside of Austin, Texas


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B.X.X Engineering CostsTowler suggests that engineering costs can be estimated as XX% of ISBL costs for larger products.

We accept this value, and estimate engineering costs to total $XXM

B.X.X ContingenciesThe minimum estimation for contingencies proposed by Towler is XX% of ISBL + OSBL costs.

Thus, we calculate:

B.X Working Capital

Working capital costs are estimated by combining cash on hand, two weeks worth of rawmaterials, and the spare parts inventory. Again, we used the values suggested by Towler. The

values for each are tabulated below along with the method used to calculate them.


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Item Costing Criteria Value

Cash on hand estimated as X weeks worth of production $XX.XM

Two weeks worth of raw materials - $X.XM

spare parts inventory estimated as X% of ISBL + OSBL costs $X.XM

TOTAL $XX.XM

B.X Variable Costs of Production

B.X.X UtilitiesCoal costs are estimated based on the US Energy Information Administration, which notes that

the cost of lignite coal is $XXXX per ton [XX]. Using our simulated coal input, we find that our

coal costs are:

( ) ( )

The average wholesale electricity cost in Texas is $XX.XX/MW*hr ($X.XXX/kW) as of XXXX [XX].

Our design uses XX.XX MW of electricity

( ) ( )

Oxygens price was estimated using Towler, which suggests $X.X/lb of OX ($XX.X/tonne). This

process uses XXXX tonnes/day of OX for combustion, and the cost is given by:

( ) Water is used to replenish water that evaporates in the cooling tower and water that is

discarded from the process as low pressure steam. The price of water is suggested by Towler to

be X.X/Mlb ($X.XXXE-XX/tonne) and this design uses XXXX tonnes/day resulting in a cost of:


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( ) B.X.X ConsumablesThe lifetime of the catalyst is finite; this analysis predicts the need to replace the catalyst (and

inert catalyst support/diluents) every five years. The value of each replacement is a product ofthe catalyst bulk volume, bulk density, and cost per weight. In the water gas shift reactor, the

catalyst is diluted to X/XXXX to keep the reaction rate under control. Catalyst cost =

$X.XX/lb=$XX.XX/kg [X] and catalyst density is X.X X.X kg/L so we use X.XX kg/L.

Ex: water gas shift, active catalyst only (diluents calculation cost is separate),

( )

( )

( ) ( )

Itemdiameter

cm

length

cm

# of

tubes

volume

(cmX)

volume

(L)

active catalyst

fraction

methanol reactor X XXXX XX XXXX X.XX X/XXXX

methanol reactor: XX X*XXX XX X.XX*XXX X.XX*XXX X

ItemVolume

(L)

catalyst bulk

density

(kg/L)

mass

(kg)

Catalyst

Cost

($/kg)

Total

Cost

methanol reactor:

active catalystX.XX*XXX X.XX X.XX*XXX $XX.XX $XXXM


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WGS active

catalystX.XX X.XX X.XX $XX.XX $XX.XX

WGS inert

supportXXXX X.XX XXXX $X.XX $X,XXX

B.X Fixed Costs of ProductionOperator Salary is calculated by (X) estimating operator hours/week (X) multiplying by operator

hourly salary. For (X), figure X-X of Peters [XX] was used to estimate the need of XX operator

hours per day per processing step. The number of processing steps was determined to be X,

which gives XXX operator hours/day. The salary of such operators was estimated at $XXX; this

value is given by the US Bureau of Labor Statistics [XX] by the median national weekly salary of

individuals with some college or associates degree. In X day:

( ) There are X days in a week, but an operator only does X shifts a week so we multiply by X/X to

get the number of operators we need on weekly salary: . Thuswe must pay about XX operators per week.

( )

Cost Item Method of Costing SourceCost

($/year)

Operator

Salary-

Textbook and Bureau of Labor

Statistics$X.XXM

Supervision XX% of operator labor PDF [XX] $XXX,XXX

Maintenance X% of ISBL investment PDF [XX]$X.XXM


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Taxes and

Insurance

X% of ISBL fixed capital PDF [XX]

$X.XXM

Plant

Overhead

XX% of total labor (including

supervision)

PDF [XX] $X.XM

Total -- -- $XX.XM

B.X RevenuesThe selling price of methanol was taken from methane, the largest methanol producer in the

world. The most recent price was methanol price from Methanex [X], whose XXXX prices were

approximately $XXX.X/tonne. This design produces XXXX tonnes/day of methanol, yielding

$XXXM/year.

Sulfur currently sells at $XXX/tonne. With the current plant output of X.XX tonne/day, the plant

is predicted to produce $X,XXX/day or $X,XXX,XXX per year.

Documents

Team1 Sp2010 TexLig2Methanol Anonymous-no Exec Summary