llydrometal/urgy, 30 ( 1992) 127-162 Elsevier Science Publishers B. Y., Amsterdam

Hydrometallurgy of precious metals recovery

ABSTRACT

C.A. Fleming Lakefield Research. Lakefield. Onto KOL 2110. Canada

(Revised version accepted January 10, 1992)

127

Fleming, C.A., 1992. Hydrometallurgy of precious metals recovery. In: W.e. Cooper and D.B. Dreis­inger (Editors). Hydrometallurgy, Theory and Practice. Proceedings of the Ernest Peters Interna­tional Symposium.llydrometal/urgy, 30: 127-162.

The hydro metallurgical process for the treatment of gold and silver ores remained unchanged for the first 70 years of this century, and consisted essentially ofJeaching in cyanide solution followed by solid-liquid separation, with the solid residues being washed as efficiently as possible, and the leach liquor being treated by zinc cementation to recover the precious metals. While this process is generally extremely efficient and fairly cheap, it does have limitations in the treatment of low-grade ores and certain complex ore types. For example, ores with a high content of clay or other soft, fine minerals are usually difficult to filter. and losses of soluble gold or silver in the residues can be unacceptably high. In other situations, where the precious metal host rock contains high concentrations of sulphides such as pyrite or arsenopyrite, for example. or base-metal oxides or carbonates, the traditional process often suffers from poor gold recovery (due to encapsulation of the precious metals in the sulphides) or high cyanide consumption, or both of these. Whereas these occurrences were fairly rare (or were avoided!) in the first half of this century, they are now assuming great importance, and each year a higher percentage of world gold production derives from sources such as these.

A number of new hydrometallurgical processes have been developed and implemented in the gold industry in the last 20 years. and these have transformed gold processing into a chemical "high tech" industry. and have allowed increasingly complex ore types and progressively lower grades of ore to be treated economically. As a result, in a period when gold production might have been expected to decline, world-wide production has almost doubled over the last two decades.

This paper describes the traditional cyanidation and zinc cementation processes, but focuses on the new developments in the industry. In particular, new leaching technologies such as heap leaching for low-grade ores and pressure leaching for refractory sulphide ores are discussed, as well as the carbon­in-pulp and carbon-in-Ieach processes that have effectively replaced filtration and countercurrent de­cantation on almost every gold plant built since 1980. Some emerging technologies such as bacterial leaching and resin-in-pulp are also discussed briefly.

INTRODUCTION

Even since the development of the cyanide/zinc cementation process in the late 19th century, hydro metallurgy has been a dominant force in the process­ing of gold and silver ores. Indeed, the cyanidation process undoubtedly qual­ifies even today as the most important and widespread of the various hydro-

Correspondence to: C.A. Fleming, Lakefield Research, Hydrometallurgy Lakefield, Onto KOL 2HO, Canada.

0304-386X/92/$05.00 © 1992 Elsevier Science Publishers B. Y. All rights reserved.

128 CA. FLEMING

metallurgical technologies used in the treatment of primary ores and concentrates. The process is so simple, efficient and inexpensive, in fact, that very few technological changes or innovations took place in the industry dur­ing the first 60-70 years of the 20th century. Most of the ores treated in this period, particularly those from the Witwatersrand in South Africa (where al­most half the gold ever mined has been recovered), were of the free-milling, quartzitic type, and it was usually possible to achieve good gold liberation merely by crushing and milling the host rock adequately. Processing of the milled ore by a combination of gravity separation (to recover coarse gold particles) and cyanidation generally yielded overall gold recoveries well in excess of 90%. The only inherent inefficiency in this process was associated with the separation of the gold-bearing cyanide solution from the solid leach residue. Consequently, this was the only area that benefitted significantly from new technology, with the replacement of cumbersome decantation/washing systems by more efficient, continuous processes, such as vacuum filtration and countercurrent thickening/washing.

Several factors combined in the late 1960s and early 1970s to trigger the technological revolution that has transformed the gold industry over the last 20 years. Firstly, and perhaps most importantly, the price of gold was released in 1968 from the rigid constraints of the $35/ oz gold standard, and was al­lowed for the first time in modern history to find its own value in response to free-market supply and demand forces. This resulted in a ten-fold increase in the price of gold in a remarkably short period of time and this, in turn, had several important spin-offs. Firstly, the profits realised by established gold­mining companies increased greatly and, consequently, there was more to spend on research and development and exploration for new deposits. Sec­ondly, it became economically viable to treat lower-grade deposits, and eco­nomically expedient to maximize extraction efficiencies, since even minor improvements in recovery now had major economic significance. Thirdly, many new gold mining companies were born in the United States, Canada, Australia, the Pacific Rim, South America and Africa, and they brought with them a refreshingly innovative and entrepreneurial spirit, into an industry that had become so conservative and resistant to change as to be almost moribund.

The second force that has played a major role in the technological revolu­tion has been the natural process of depletion of free-milling, easily extracta­ble, gold resources around the world. In most of the important new deposits that are being found today, the gold is very finely disseminated and encapsu­lated in a host matrix that is inert and impermeable to the cyanide leach solution. In these cases, it is necessary to first break down the host matrix (usually by hydrometallurgical processing) and only then is it possible to leach the liberated gold. There have been many technical innovations in this area, and the issue facing those involved in the processing of these so-called refrac-

HYDROMETALLURGY OF PRECIOUS METALS RECOVERY 129

tory gold deposits is not so much whether or not the gold can be recovered efficiently, but rather, which one of several processing options should be adopted.

A third force has entered the primary precious metal mining industry and, depending on the perspective, this force either hangs as a sword of Damocles ( over the developers of new properties) or opens up as a window of oppor­tunity (for those involved in the development of new technology). This force is the environmental legislation that is increasingly regulating and controlling the precious metals industry, particularly in the First World. Environmental issues are already having a significant influence on process flowsheet selec­tion and treatment costs, and will probably be the major stimulus for the de­velopment of new technology in the gold and silver mining industries over the next 10-20 years.

THE CONVENTIONAL CYAN IDA TION /CEMENTA TION PROCESS

Practical aspects

The first step in the cyanidation process consists of crushing and milling the gold-bearing ore to a size in which the gold particles are liberated from the host rock or mineral. For practical as well as environmental reasons, the mill­ing is performed on a wet slurry of the ore, usually at a low solids density, in the 5-15% range, and in closed-circuit, with classification by either cyclones or screens. The milled product is then thickened to between 30 and 50% solids (depending on the rheology of the slurry), and advances to the cyanide-leach­ing plant.

Oxygen is introduced to the slurry, usually in the form of compressed air that is blown from the base of a conical-bottomed mixing tank (called a Pa­chuca). The air, therefore, provides not only the oxygen that is necessary to oxidize gold from the metallic state, in which it occurs in nature, to the gold (I) state, but also generates the hydrodynamic conditions needed to keep the slurry fully in suspension. In recent years, mechanically agitated, flat-bottomed leaching tanks have been preferred to Pachucas, particularly since the intro­duction of the carbon-in-Ieach process, and in certain instances pure oxygen, or even stronger oxidants such as hydrogen peroxide, have been used in place of air. .

The slurry flows by gravity through a series of tanks, numbering anywhere from 6 to 20, with a residence time of 1-4 h in each (ank, with the overall residence time being determined by the rate of gold leaching from the partic­ular ore. The various factors affecting the rate ofleaching are discussed briefly below, but the residence time is generally about 24 h for a free-milling ore. Cyanide is usually only added to the second tank in the series, to ensure that the slurry is thoroughly preaerated and the leach solution saturated with ox­ygen prior to contact with cyanide. Modern gold plants monitor and control

130 C.A. FLEMING

the concentrations of cyanide and oxygen in the leach solution throughout the leaching plant, to ensure consistent performance and to minimize cyanide consumption.

The third step in the process involves separating the gold-bearing leach so­lution from the gold-depleted ore. This operation has always been costly and inefficient, and was an important stimulus for the development and rapid implementation of the carbon-in-pulp and carbon-in-leach processes. The methods used are either vacuum filtration (drum or belt filters) or counter­current decantation (CCD) and, in both cases, inefficiencies stem from the difficulty in washing the last traces of gold-bearing solution from the solids without creating water-balance problems in the overall circuit. In general, about 1.5-1.8 m3 of gold-bearing solution is produced for every ton of ore milled, and up to 2% of the leached gold can be lost in the final leach residue.

Once separated from the washed solids, the filtrate or gold pregnant solu­tion is first clarified to remove the last traces of suspended solids, and is then de-gassed under negative pressure to eliminate oxygen from solution. This operation, which is performed in a Crowe tower, is an important step in the overall process, because oxygen would otherwise participate in wasteful side­reactions in the next step in the process. In this next step, gold is recovered from the pregnant solution by reducing it back to the metallic state. A number of powerful reductants, such as sodium borohydride and aluminium metal have been shown to be technically viable in this role. When cost and effi­ciency are considered, however; nothing compares with zinc metal, and all conventional gold plants around the world employ zinc cementation to re­cover gold from the cyanide leach liquor. An excess amount of zinc is re­quired for efficient gold recovery, and the precipitate therefore consists of a mixture of gold and zinc metal. This precipitate is recovered by filtration and the barren filtrate is recycled as wash water to the filtration or CCD plant. In the final step in the process, the excess zinc in the zinc/gold sludge is either converted to zinc oxide in a calcining furnace, or is dissolved in a mineral acid, prior to smelting the gold powder to bullion.

A full description of the conventional process is obviously beyond the scope of this review, and this quick walk through the flowsheet is intended merely as an introduction to the steps in the process where hydrometallurgy is fea­tured. The chemistry involved in these steps is discussed in more detail in the next section. Readers who are more interested in the practical aspects of the conventional gold process are referred to the recent reviews by Young [1, pp. 277-330] and Bosley [2,pp. 331-334].

The chemistry of cyanidation

The action of alkaline cyanide solution on gold metal was first described in a patent by Elkington in 1840, although this was in an electroplating context,

HYDROMETALLURGY OF PRECIOUS METALS RECOVERY 131

and did not refer to primary ore treatment. More than 40 years were to pass before the Forrests and MacArthur conceived the idea of applying this reac­tion to the dissolution of gold and silver from low-grade ores, and were awarded a British Patent (No. 13,174) on cyanide leaching in 1887, and an­other on the use of zinc dust for precipitation, in 1888. The first commercial use of cyanide for gold extraction was at the Crown Mine in New Zealand in 1889, and applications in South Africa and the United States followed soon after.

The chemistry of the dissolution of gold and silver in alkaline cyanide so­lutions has been the subject of considerable investigation ever since these first practical applications. Although the overall features of the reactions involved are well-established, there is still uncertainty regarding the details of some aspects of the mechanism of the process. Undoubtedly, the major advance in the understanding of cyanidation was made by Kudryk and Kellogg [3], who conducted experiments to demonstrate that the dissolution of gold in cyanide solutions is essentially an electrochemical process, with the following overall stoichiometry:

(1)

This reaction consists of cathodic and anodic half-reactions. The anodic re­action involves the oxidation of gold (0) to gold (I):

4Au+8CN-=4Au(CN)i +4e (2)

and is accompanied by the cathodic reduction of oxygen at the surface of the gold particle:

O2 +2H20+4e=40H- (3)

Those interested in detailed mechanistic studies of the gold cyanidation reaction are referred to the definitive studies of Cathro and Koch [4], Fink and Putnam [5] and Nicol [6]. The remarkable success of cyanide as a lixi­viant for gold can be traced at least partially to the enormous stability of the dicyanoaurate ion (P2 = 1038

, [7]). This allows gold to be leached efficiently at very low concentrations of cyanide « 0.0 1 mol dm - 3), and the di­cyanoaurate complex remains in alkaline solution even when the free-cyanide concentration falls to zero. This contributes to the selectivity of the process, because the stability of most of the other metal cyanocomplexes that are pres­ent in cyanide leach solutions are somewhat lower than aurocyanide. This in turn benefits the economics of the process, since the consumption of cyanide via other metal complexation reactions is usually insignificant. The fact that oxygen is an effective oxidant for gold metal in cyanide solution, even at the very low concentrations that are typical of air-saturated aqueous solutions (1.5-3XIO- 4 mol dm- 3, i.e. 5-10 mg I-I) can also be traced to the great stability of the dicyanoaurate ion.

132 C.A. FLEMING

In contrast to cyanide, all the alternative leaching agents that have been proposed (halides, thiourea, thiosulphate, thiocyanate) yield gold complexes that are considerably less stable than aurocyanide (Table 1). Therefore, in order to stabilize these complexes in aqueous solution, and to achieve accept­able leaching rates, it is necessary to employ far higher lixiviant concentra­tions (usually 0.1-1 mol dm- J

), and this, in tum, means that it is necessary to recycle spent leaching solution if unacceptable costs and environmental problems are to be avoided. Oxygen is also unsuitable as the oxidant for gold in these situations, because of its low concentration and the limitation this places on the rate of leaching. It is therefore necessary to resort to more ex­pensive oxidants, such as hydrogen peroxide, chlorine, bromine, ozone or the ferric ion. When all of these factors are considered, it is not surprising that cyanide leaching has dominated gold processing for over 100 years, and will probably continue to do so in the future, despite the fact that the cyanide ion is itself now coming under the close scrutiny of environmental legislators. There are, however, potential niche technologies for the alternati ve lixi viants, and these will be discussed below.

Another important advantage of cyanide over most of the other leaching chemistries (thiosulphate is the only exception) is the fact that the reaction takes place in an alkaline environment. The dissolution of gold requires the cyanide to be present as the free cyanide ion, CN-. Therefore, since the com­position of cyanide solution is determined by the hydrolysis reaction:

(4)

which has a pKa of about 9.3, a high pH (usually in the 10.5-11 range) is necessary to ensure that most of the cyanide is in the ionic form. A high pH is also necessary for safety and economic reasons, as HCN is a volatile and poi­sonous gas, which is purged from the leach slurry during vigorous air agita­tion. An important advantage of leaching under alkaline conditions is that

TABLE I

Stability constants for a selection of complexes of gold (I) and gold (III) [8]

Gold(l) fJ2 Gold (III ) P4 complex complex

Au(CN)i 2x 1038 Au(CN)4" _1056

AU(S203)~- 5x 1028 AuI4" 5x 1047

Au(CS(NH2h)i 2x 1023 Au(SCN)4" 1042

Auli 4x 1019 AuBr4" 1032

Au(SCN)i 1.3 X 1017 AuCI4 1026

AuBri 1012

AuCli 109

HYDROMETALLURGY OF PRECIOUS METALS RECOVERY 133

the dissolution of base metals such as copper, zinc and nickel is substantially reduced, resulting in cleaner effiuents than those generally produced in acid­leach systems. The non-corrosive nature of an alkaline cyanide solution also means that cheaper materials can be used for the construction of a gold plant. The pH is regulated by the addition oflime to the leach slurry, usually in the milling circuit.

The one drawback of the cyanide/oxygen leaching chemistry for the disso­lution of gold is the low solubility of oxygen in water. It can be shown theo­retically [8] that, under most conditions, the rate of gold leaching is controlled by diffusion of oxygen to the metal surface. Factors that playa role, therefore, in the rate of gold dissolution are:

( 1) The concentration of oxygen in water. This, in turn, is a function of air­pressure, and oxygen solubility varies with altitude in the 5-10 mg I - I range for air-saturated solutions. Temperature also plays a role, with oxygen solu­bility varying inversely with temperature. An increase in temperature, there­fore, can result in a decrease in the rate of gold leaching, despite the fact that the diffusivity of oxygen improves. There are several methods of increasing oxygen concentration. One is to employ pressure cyanide leaching [9], as practised in a pipe-reactor at the Consolidated Murchison gold/antimony mine in South Africa [10]. The high leaching rates observed in the Kamyr leaching tower can probably also be traced to the relatively high pressures that are attained at the bottom of a tall tower, and oxygen concentrations in the range of 20-30 mg 1- I have been reported for this process [11 ]. Another is to use oxygen instead of air to increase the partial pressure of oxygen under atmospheric conditions. Alternatively, if the economics warrant it, hydrogen peroxide can be used to supply oxidant to the surface of the gold particle, as proposed by Loroesch [12].

(2) The dispersion of air in the slurry. This, in turn, is determined by fac­tors such as the viscosity of the slurry and the method by which air is physi­cally introduced into the slurry.

(3) The mass transfer of oxygen in the slurry. This is a function of the hy­drodynamic conditions in the mixing tank, and varies with slurry rheology, mixer design and energy input.

( 4) The presence in the slurry of parasitic oxygen-consuming reactions. These species vary from naturally occurring organic molecules (humates, lipids) to inorganic compounds such as the ferrous ion, and their presence can seriously retard the rate of gold leaching. This is usually countered by employing one or more stages of air (or oxygen, or hydrogen peroxide) pretreatment, prior to the introduction of cyanide.

Cyanide concentration is less important than oxygen concentration, and it can be shown theoretically [8] that the rate of gold leaching is independent of free cyanide concentration at levels above about 2.5 X 10 -3 mol dm -3 (60 mg I-I CN - ). At concentrations below 2.5 X 10 - 3 mol dm - 3, the rate of

134 C.A. FLEMING

gold leaching is controlled by the mass transfer of cyanide ion to the surface of the gold particles, rather than by mass transfer of oxygen. This is generally only a factor when there are cyanide-consuming constituents other than gold in the ore. Examples include iron sulphide minerals such as pyrrhotite (which consumes both oxygen and cyanide), copper minerals such as chalcocite, malachite and azurite, and various arsenic and antimony sulphides. A num­ber of sulphur-containing inorganic anions such as sulphide, thiosulphate and polythionate also react with and consume oxygen and cyanide.

Although much work has been done in this area, particularly in attempts to reduce the reactivity of certain copper minerals in cyanide solution, little suc­cess has been achieved, and the only method of ensuring consistently good gold recovery in these situations is to employ high cyanide concentration and tolerate high cyanide consumption.

The final factor that plays a role in the rate of leaching of gold from free­milling ores is the size of the gold particles themselves. It can be shown that, under conditions of oxygen mass-transfer rate control (in air-saturated solu­tion), gold will dissolve at a rate of 3.25 mg cm -2 h -I. It is possible to cal­culate from this that spherical gold particles with a diameter of 44 ~m will take approximately 13 h to dissolve, while 150 ~m particles will take more than 40 h. It is for this reason that most gold plants include gravity separation equipment in the milling circuit if there is coarse gold in the deposit.

The mechanism by which silver metal leaches is also electrochemical, with the same stoichiometry as shown. for gold in eq. (1). The stability of the Ag(CN)2" complex is considerably less than that of Au(CN)2" , however, so the reaction is less favourable, requiring higher oxygen and cyanide concen­trations for comparable leaching rates. Silver is often present in nature as the sulphide Ag2S, and, according to Shoemaker and Dasher [13], reacts with cyanide as shown in eq. (5):

Ag2S+4CN-~2Ag(CN)2" +S2- (5)

This reaction is slow and reversible, and only proceeds efficiently to the right­hand side if the cyanide concentration is kept high (usually> 0.04 mol dm -3) and the sulphide concentration low (usually by simultaneous oxidation to thiosulphate ):

2S2-+202+H20-.S20j- +20H- (6)

Chemistry of cementation

The zinc-cementation process for the recovery of gold and silver was intro­duced in 1890, at the same time as the cyanide leaching process, and consisted essentially of passing the gold cyanide pregnant solution through a bed of zinc shavings. The process was fairly crude and inefficient, and during the follow-

HYDROMETALLVRGY OF PRECIOUS METALS RECOVERY 135

ing 30 years three major modifications were introduced that improved the efficiency of the process. These innovations involved the addition of lead salts to the pregnant solution (1894), the use of line dust in place of line shavings (1897), and deaeration of the pregnant solution prior to cementation (1916). The historical review compiled by Leblanc [ 14] on the recovery of gold from pregnant liquors gives an excellent account of these early developments.

Once the major metallurgical and economic shortcomings of the process had been resolved, the incentive for further research in this area largely fell away. Fortunately, the rekindled interest in electrochemistry that began in the 1960s led to research into the mechanism of cementation, and much of the mystique has been stripped away.

The major reactions are the cathodic deposition of gold and the anodic cor­rosion of line, which occur at the surface of the zinc particles:

Au(CN)i +e=Au+2CN­

Zn+4CN-=Zn(CN)~- +2e

(7)

(8)

Side reactions that influence the economics of the cementation process are the reduction of water and dissolved oxygen:

2H20+2e=20H-+H2 (9)

O2 +2H20+4e=40H- (10)

both of which involve the simultaneous and wasteful consumption of line, by oxidation to Zn2+ and formation of zinc tetracyanide, i.e. if there is oxygen present in solution, the following reaction occurs:

2Zn+8CN- +02 +2H20=2Zn(CN)~- +40H- (11)

This reaction is minimized in a modem gold plant by de-aeration of the preg­nant solution in a Crowe vacuum tower. The consumption of zinc via the hydrogen evolution reaction (eq. (9» can also be controlled to a certain ex­tent by increasing the potential of the cementation reaction, by decreasing the concentration of free cyanide, as discussed by Nicol et al. [8]. In practice, however, if the cyanide concentration is reduced too much, a passivating layer of zinc hydroxide may form on the surface of the zinc particles [8]. The del­eterious effects of excessive and insufficient concentrations of free cyanide were noted by Leblanc [14], who recommended that the concentration of free cyanide should be maintained between 2X 10-3 and 6X 10-3 mol dm- 3

(50 and 150 ppm CN-). The role of dissolved lead ions (usually lead nitrate) in solution is also

apparently associated with a shift in the redox potential at the surface of the linc particles, and Pb2+ concentrations as low as 1-2 ppm have been shown to shift the potential in a positive direction by as much as 0.2 V [15,16]. The mechanism associated with this considerable enhancement is not clear, but

136 CA, FLEMING

similar effects have been found with thallium, mercury and bismuth ions [16]. The addition of lead to above 10 ppm may actually retard the cementation of gold [2,8] so this, too, needs to be carefully controlled.

Finally, the pH of the pregnant solution also influences the redox potential, which shifts in a negative direction with increasing alkalinity. This can be beneficial if the rate of the cementation reaction is not entirely under diffu­sion control, as elucidated by Nicol et al. [8]. Leblanc [ 14] reported that the cementation of gold was practically constant over the pH range 8-11, but that a marked improvement could be obtained by an increase in pH to a value between 11.5 and 11.9.

The zinc metal used for cementation must clearly have a large surface area, because the overall rate of any electrochemical solid-liquid reaction is di­rectly proportional to surface area. Zinc dust fulfils this requirement, but is readily oxidized by atmospheric oxygen during storage. The coating of zinc oxide must be dissolved before cementation can occur:

(12)

which accounts for the frequent necessity for circulation of the pregnant so­lution for considerable lengths of time at the start-up of a cementation reac­tion. High concentrations of free cyanide may be beneficial during this period.

The most common impurities in gold pregnant solutions, including sul­phite, sulphate, thiosulphate, ferrocyanide, thiocyanate, copper, zinc, nickel and cobalt, appear to have little or no effect on the cementation process [ 14,15], although deleterious effects can be observed if the concentrations are very high. Plaskin et al. [17] reported that copper adversely affected the cementation of gold, which almost ceased at a copper concentration of 200 mg 1- I. However, the deleterious effect of copper decreases as the concentra­tion of cyanide is increased, from which it can be inferred that the effect of copper may be due' to a reduction in the concentration of free cyanide as a result of the formation of cyanide complexes. Hancock and Thomas [18] reported that nickel at concentrations higher than 200 mg 1-1 has a slight retarding effect on cementation.

The only species that have a marked deleterious effect on cementation ap­pear to be sulphide ions and soluble compounds of arsenic and antimony [ 14,15,17]. The effects are observed even at very low concentrations (I mg 1-1 and lower). The poisoning effect of sulphide ions is thought to be due to the precipitation of insoluble zinc sulphide on the surface of the zinc parti­cles. No satisfactory mechanism has been prpposed for the effect of arsenic and antimony.

Aluminum is sufficiently electronegative to reduce aurocyanide ions to gold, and its use as a cementing agent was patented by Moldenhauer in 1893. The factor mitigating against its use is the dissolution reaction:

HYDROMETALLURGY OF PRECIOUS METALS RECOVERY 137

(13 )

which requires the pH value of the solution to be above 12 in order to keep AI3+ in a soluble form. Furthermore, calcium aluminate has a low solubility, and the use of sodium hydroxide or sodium carbonate is required to adjust the pH; the use of lime must be avoided. A historical review of the use of aluminum for the cementation of gold has been compiled by Nagy et al. [ 19].

NEW PROCESSES IN GOLD AND SILVER RECOVERY

Free-milling ores

Carbon-in-pulp (ClP) and carbon-in-leach (ClL) The ability of activated carbon to adsorb gold from solutions has appar­

ently been known since 1847 [20], although it was not until 1894 that the use of carbon as a precipitant for gold in cyanide solution was first proposed and patented by Johnson [21]. Unfortunately for Johnson, it was to be another 20 years before the process was used commercially for the first time, at the Yuanmi Mine in Western Australia [22], and a further 60 years before the technology had really come of age as a viable alternative to the zinc cemen­tation process.

Progress over this period, particularly the pioneering work performed by Zadra and others at the United States Bureau of Mines [23,24], has been reviewed by McDougall and Fleming [25].

The most important development in the widespread commercialisation of CIP was the introduction of the process at the Homestake Mine in South Da­kota, in 1973, where about 2200 tons of fine slimes are treated per day [26]. The Homestake Gold Mine always had separate treatment processes for its sand and slime fractions and when, in 1971, the slime filtration plant needed to be replaced, a CIP plant was installed on the advice of the US Bureau of Mines. This highly successful operation transformed the image of the CIP process from an "experimental" and small-scale process to one that could be adapted for the treatment of high-tonnage flows.

The process spread rapidly, firstly in South Africa in the late 1970s, and then throughout the rest of the gold-producing world in the 1980s, and it is now the preferred process for all new gold plants, probably accounting for over 50% of world-wide gold production. One of the important contributions made by the South Africans was their realisation that the process need not be restricted to the treatment of materials that were difficult to filter, such as the slimes fraction of an ore, but could be applied to almost any feed, achieving high overall gold recoveries, with relatively low capital and operating ex­penses. CIP plants were built in South Africa that treated anything from whole ores, to flotation concentrates, flotation tailings, roasted calcines, re-pulped

138 C.A. FLEMING

filter plant residues or, most significantly, old residue dumps. Many hundreds of millions of tons of tailings are scattered throughout the Witwatersrand re­gion of South Africa and, because of the superior economics of the CIP pro­cess compared to zinc cementation, it became feasible for the first time to profitably re-treat this material. Between 40 and 50 million tons of tailings are now being treated in this way every year, yielding about 20 tons of gold per annum from those wastes, and exposing, under the old dumps, hundreds of acres of prime industrial land for development. The economics of the dump re-treatment process require that very large tonnages be processed, and this led to the next major contribution made by the South Africans; namely, the development of the process engineering that was required to take the process from the 60,000 ton/month scale of the Homestake operation, to the 1.5-2 million ton/month scale of the ERGO operation, for example. One of the most important developments in this area was the move away from external interstage screens, as used at Homestake, to internal air-swept screens. In the external screening concept, all of the forward flow of pulp had to be pumped from each adsorption tank up onto a vibrating screen, located on a deck above the tank. The carbon in the screen overflow flowed back into the tank from which it had been pumped, while the pulp underflow advanced to the next tank in line. While this approach was viable on the small scale of the Home­stake plant, it was clearly impractical on the large-scale South African plants. Following the early development of the EPAC (equalized-pressure, air­cleaned) screen in South Africa, a number of new and innovative designs, based on similar principles, appeared on CIP plants around the world, partic­ularly in Australia. Many of these and other developments have been re­viewed in papers by Laxen and co-workers [27,28] and Fleming [29], which also detail the layout of a typical multi-stage, countercurrent CIP adsorption plant.

The final contributions made in South Africa have been on the metallurgi­cal processing side. Fundamental work carried out at the Anglo American Research laboratory (AARL) and at Mintek, resulted, for example, in the development of a successful alternative to the Zadra method of eluting gold and silver from carbon [30] and in the design of a new, more efficient type of electrowinning cell for the recovery of gold from carbon eluates [31 ].

Work carried out at AARL and Mintek also went a long way toward resolv­ing the mechanisms involved in the adsorption [32-34] and elution [35] of gold and other competing species on activated carbon, and of the factors af­fecting adsorption efficiency [36]. Simple mathematical models, based on the kinetics of gold extraction by carbon, were developed [37], which ad­vanced the understanding of several aspects of the countercurrent CIP ad­sorption process. For example, by modelling batch CIP data, it was possible to design parameters such as the flow rate and concentration of carbon and the number of adsorption stages required for optimum metallurgical effi-

HYDROMETALLURGY OF PRECIOUS METALS RECOVERY 139

ciency in a continuous process. Since then, a more rigorous model has been developed, which describes the adsorption process more effectively but which is, by necessity, more complex [38].

The fundamental difference between a CIP flowsheet and the conventional, zinc-cementation flowsheet lies in the fact that, in order to recover gold by reduction on zinc dust, it is first necessary to produce a clarified filtrate, whereas in the CIP process it is possible to extract gold cyanide directly from the slurry. In this way, the costly and inefficient solid-liquid separation pro­cesses of a conventional gold plant (filtration or CCD) are replaced by the relatively simple and inexpensive screening procedures that are used in a CIP plant to recover the carbon granules from the leach slurry.

Early estimates of the capital and operating costs associated with the CIP process indicated that savings of 20-50% could be expected, compared to the conventional process (Table 2), and these estimates have been vindicated, for the most part, on the great many CIP plants operating around the world today. This has meant that lower grade gold ores can be economically treated than would previously have been the case, and this has been a major stimulus to new gold mine development around the world. Another important benefit of the CIP process is the generally improved efficiency of gold extraction that can be achieved. This stems from two factors. Firstly, soluble gold losses from a conventional filtration or CCD plant usually amount to about 1 % of the pregnant solution (i.e. 0.03-0.05 mg 1-1), in the case of easily filterable sol­ids, while even higher soluble losses are suffered in the case of poorly filtering or settling solids. By contrast, soluble losses ofless than 0.01 ppm can usually be achieved on a well-managed CIP plant. Secondly, the additional 5 or 6 hours leaching time in the CIP adsorption contactors usually results in extra gold dissolution. This extra dissolution can be considerable ifthere are "preg­robbing" constituents in the ore. Evidence for extra dissolution is obtained by an analysis of the washed solids in the feed and the discharge from a CIP plant. In most instances, the residue solids are lower in gold than the feed solids, despite the presence of fine abraded carbon from the CIP tanks in the residue [27].

The final advantage of the CIP process is that it is far less vulnerable than

TABLE 2

Comparative costs of carbon-in-pulp and conventional (Merrill Crowe) processes

Process Capital cost Operating cost

Conventional 1.00 1.00 CIP [51) . 0.68 0.77 CIP [62] 0.75 0.94 CIP [42, p. 382) 0.37-0.58 0.88-1.18

140 C.A. FLEMING

the conventional process to impurities such as sulphide, arsenate and anti­monate in the leach liquor, and parameters such as cyanide concentration, pH and oxygen concentration do not have to be carefully controlled, as their influence on the loading of gold on carbon is fairly minor [39].

The mechanism of adsorption of gold cyanide onto activated carbon has been studied extensively over the last 10 years, and the results of these studies have been reviewed recently [34]. Most of the evidence indicates that gold is reversibly adsorbed on the carbon, as the aurocyanide ion, without undergo­ing any chemical change. Adsorption occurs on the extended surface of the porous carbon granules (total surface area -1000 m2 g-I), and is only lim­ited by physical access of the aurocyanide ion into the smaller micropores ( < 10-20 A). The adsorption efficiency of different ions and molecules decreases with increasing hydration and hydrophilicity and, in the case of inorganic metal complexes, is believed to involve charge transfer from the graphitic planes at the carbon surface to the metal atom. Therefore, factors that decrease the hydration and ionicity of the aurocyanide complex, such as increasing acidity (H++Au(CN)2"-+HAu(CNh) or ionic strength (Ca2+ + 2Au (CN) 2" -+Ca [Au (CN h h), enhance the adsorption efficiency on activated carbon. Conversely, an increase in the concentration of hydroxyl or cyanide ions reduces the adsorption of gold on carbon and this, coupled with the reduction in adsorption capacity that occurs with increasing temperature, forms the fundamental basis of both the Zadra and the AARL elution processes.

In practice, only a small fraction of the total adsorption capacity of carbon for gold is utilized in normal CIP operation. It has been shown in the labora­tory, for example, that gold loading values as high as 3 mol kg-I (- 600,000 g t - I ) can be achieved, whereas the loading of gold on carbon in the first stage of a CIP plant seldom exceeds 0.05 mol kg-I (10,000 g t- I ). There are sound practical and economic reasons for this. Firstly, the aurocyanide ion adsorbs relatively rapidly onto carbon in the initial stages of the extraction reaction, during which adsorption occurs in the relatively large macro- and meso-pores. Contact times of weeks or even months are required, however, to fully utilize the total surface area available for aurocyanide adsorption [36]. Secondly, carbon breaks down slowly in the adsorption contactors as a result of the ab­rasive action of the ore in the leach slurry. This carbon is lost from the circuit once it is reduced to a size smaller than the interstage screens, and this would result in significant gold losses if the carbon were loaded to excessively high gold values. This problem is exacerbated by the fact that the rate of the re­versible desorption reaction is slow and, therefore, a highly-loaded carbon granule in the first adsorption contactor would only desorb gold slowly if it fractured and flowed co-currently with the slurry out of the CIP plant [25]. A third reason for keeping gold loadings on the carbon at a relatively low level relates to lock-up of gold in the CIP plant. Carbon inventories on large CIP

HYDROMETALLURGY OF PRECIOUS METALS RECOVERY 141

plants range from 50 to 100 tons or more, and the gold loaded on this carbon is removed from an income-earning capacity for the lifetime of the CIP plant. This can have a critical effect on cash-flow during start-up and, ultimately, on the economic viability of a new operation.

One of the implications of operating CIP plants at gold loadings well below the loading capacity of the carbon is that plant performance is a function more of the kinetic than the equilibrium or thermodynamic characteristics of the carbon. The rate of gold loading is determined by mass transfer of the auro­cyanide ion both in the pulp and within the pores of the carbon granules [36]. Therefore, factors such as the agitation efficiency in the adsorption contac­tors, the viscosity and pulp density of the slurry, and the particle-size of the carbon granules can all have a significant bearing on CIP performance.

The only negative implication of operating CIP plants with fairly low gold loadings is that carbon flow rates are relatively high. This increases carbon handling requirements and, moreover, because the capacity of the elution and reactivation plants are directly proportional to carbon flow rate, capital and operating costs are affected adversely. Consequently, elution and reactivation tended to be the most cost-intensive unit operations in the flowsheet, and this created bottlenecks in the process as engineering companies sought to control capital costs by specifying minimum plant requirements [27].

Regeneration ofloaded carbon to that offresh activated carbon is generally a three-step process. In the first step the loaded carbon is treated with a min­eral acid (usually hydrochloric, although nitric is also used on some plants). The main purpose of the acid-wash is to strip calcium carbonate from the carbon, which can build up to levels of 5% or more in the adsorption circuit. The calcium comes from the lime used to maintain the pH above 10.5, while the carbonate is believed to derive predominantly from the oxidation of cyanide:

CN- + 20H-~CNO- + H20+ 2e

CNO-+2H20+H+~NHt + HC03"

(14 )

(15 )

This reaction is catalyzed by activated carbon [40], and results in a decrease in cyanide concentration in the CIP adsorption tanks of about 50% of the feed value. The acid-wash also removes certain base metals from the carbon, as well as fine slimes of silica, clay, and iron oxides, and results in more efficient precious metal elution in the second step in the regeneration process.

As the adsorption of gold on activated carbon is a thermodynamically re­versible process, chemical and physical factors that inhibit adsorption will enhance elution, and vice versa. A number of these factors are utilized in the elution of carbon on all CIP plants. The most important parameter is temper­ature, since both the kinetics (activation energy 66.5 kJ mol-I) and the ther-

142 C.A. FLEMING

modynamics (exothermic heat of reaction""' 40 kJ mol- I ) of elution improve with increasing temperature. The modem trend is towards temperatures higher than 100°C, with the use of pressurized equipment. A second important pa­rameter is ionic strength. An increase in ionic strength generally favours ad­sorption, and the use of deionized water to achieve enhanced rates of gold and silver elution in the AARL process is a manifestation of this phenome­non. However, the rate of elution also increases with increasing concentra­tions of solvating anions such as cyanide or hydroxide. Therefore, because of the opposing influences of the anion effect and the overall ionic strength ef­fect, the rate of elution passes through a maximum as the rate of the eluting salt is increased [35]. Another important factor that can be utilized to en­hance the rate of elution is the effect of polar solvents, such as acetonitrile, acetone, methanol and ethanol. This effect is attributed to an increase in the activity of the cyanide ion and a decrease in the activity of the aurocyanide ion in polar solvents, as compared with water [41 ], and results in significant enhancement of the rate of elution.

Gold is usually recovered from elP eluates by electrolysis, but in some cases zinc cementation is still preferred. The concentration of gold and silver in a elP eluate is usually 2 to 3 orders of magnitude higher than in the original pregnant leach liquor and, owing to the selectivity of the elP process, the ratio of precious metals to contaminants is also much higher. Consequently, the final gold recovery process is less critical to the overall efficiency and eco­nomics of the process than is the case in the conventional cyanidation/ce­mentation process, and can be "over-designed" with little impact on capital costs. This part of the process has therefore caused few problems, and has not attracted much attention in the way of new developments. Electrowinning and cementation of gold from elP eluates, and the characteristics of the dif­ferent cells that are used commercially, have been reviewed recently by Bailey [42, pp. 550-570] and Nicol et al. [8, pp. 890-894].

The third step in the regeneration process involves the thermal reactivation of the carbon, to either carbonize or volatilize any organic compounds that adsorb onto the carbon from the cyanide pulp in the elP contactors. The carbon is heated to 650-750 o e for up to 30 min, in the absence of air, and is then quenched in water and screened (to remove fines) before recycling to the last adsorption contactor in the elP plant.

There are two other processes involving the extraction of gold and silver from cyanide solution with activated carbon. The first is the carbon-in-Ieach (elL) process. For most gold ores, the time required for efficient gold leach­ing is between 24 and 36 h, whereas the time required for efficient extraction of the gold by carbon is usually only 5-10 h. In elP, the leaching is essentially complete prior to carbon adsorption, and the elP adsorption section is about 5 times smaller than the leaching plant. In principle, the elP adsorption sec­tion could be dispensed with and the carbon added to the leaching tanks rather than to small adsorption tanks at the end of the leaching section, and this is

HYDROMETALLURGY OF PRECIOUS METALS RECOVERY 143

the basis of the elL process. In choosing between elP and elL for a particular application, a number of factors must be considered:

( 1 ) The major advantage of elL over elP is the lower capital cost of the plant, which results from the removal oHive or six elP adsorption tanks and ancillary equipment from the flowsheet. The operating costs associated with the elP section would also be eliminated.

(2) Because the elL leaching tanks are much bigger than elP tanks, elL has a further advantage over elP in that the surface area available for inter­stage screening is much greater, and screening is less likely to be a bottleneck in the process.

(3) A disadvantage of elL is that the carbon concentration in the pulp is lower than in elP for the same carbon inventory. Hence, more pulp is trans­ferred upstream during countercurrent transfer of the carbon, which places an increased load on the pumps or airlifts and on the interstage screens.

(4) The metallurgical efficiency is lower in elL than it is in elP. This is because leaching is incomplete when the pulp encounters carbon in the first adsorption stage of elL. Consequently, the concentration of gold in solution is lower and, to compensate for a decrease in the kinetic driving force, more carbon is needed to match the metallurgical performance achieved by elP. This has been demonstrated by mathematical modelling [37, Part II] and, even for an ore from which the gold is leached at a relatively fast rate, the carbon inventory in elL should be about 25% higher than in elP, for the same metallurgical performance.

(5) Most elP plants have performed well below design specification on start-up, often because the pulp contains poisons that were not exposed dur­ing piloting. The adverse effect of poisons can be countered by the use of a larger carbon inventory or a larger number of adsorption stages. Because a CIP plant has far more flexibility than a CIL plant, it lends itself better to the introduction of these types of changes at some stage after the initial commis­sioning of the plant.

The third process utilizing activated carbon involves extraction of gold and silver from solutions rather than slurries.

The major cost advantage of processes using activated carbon over zinc cementation lies in the ability of granular carbon to extract gold cyanide di­rectly from pulp, thus dispensing with the costly countercurrent decantation or filtration methods of solid-liquid separation. However, there are also in­stances where the recovery of gold from clarified or unclarified solutions by activated carbon is preferred to zinc cementation, and a number of plants around the world recover gold in this way.

The choice between carbon columns and zinc cementation is based on anal­yses of capital and operating costs and consideration of the metallurgical ef­ficiency. As a broad generalization, treatment in carbon columns is more eco­nomical for large volumes of low-grade solutions containing mainly gold,

144 CA. flEMING

whereas zinc cementation is preferred for relatively small volumes of high­grade solutions, particularly those rich in silver.

Obvious examples of large-volume, low-grade solutions are tailings ponds from conventional eeD or filtration plants, and there are a number of plants around the world recovering gold from sources such as this, in packed carbon columns. The pay-back period for this sort of scavenging operation is usually no more than a few weeks!

Another example of a situation in which carbon is usually preferred to zinc is in the treatment of solutions from heap-leaching operations. Here too, the main advantage of carbon over zinc cementation is that the pregnant solution need not be clarified or de-aerated prior to feeding to a carbon column.

The downstream processing to strip gold from the carbon and reactivate the carbon is the same as in the elP and elL processes, although it may not be necessary to reactivate the carbon as frequently because the concentration of organics in solution is usually less than in pulps.

Resin-in-pulp The resin-in-pulp (RIP) process for precious metal recovery is in its in­

fancy in the Western World, although it has apparently been practised widely in the USSR for many years [43], and was used in the uranium industry in the USA, in the 1960s and 1970s.

A detailed description of the mechanisms involved in the adsorption and elution of aurocyanide on ion-exchange resins is beyond the scope of this pa­per, but the chemistry involved has been covered in a recent review [44].

The history of anion-exchange resins in gold processing, which has been detailed in a recent review [8], is not as long as that of activated carbon. The earliest detailed studies were carried out in the UK in the early 1950s and, since then, there have been sporadic bursts of activity in a number of labora­tories around the world. The successful commercialization of the eIP process has undoubtedly contributed to the current surge of interest in the RIP pro­cess, as it has served to demonstrate the very substantial benefits of "in pulp" processing. Moreover, a number of the innovations that have taken place in the engineering oflarge-scale eIP processes will probably benefit the RIP pro­cess, as the broad engineering features of the two processes are likely to be similar.

Up to now the RIP process has not received serious consideration by min­ing or engineering companies during flowsheet evaluation for greenfield op­erations. However, resins have several characteristics that are quite different from carbon, which make them potentially more versatile substrates for the recovery of gold. The following are examples:

(I) Resins are superior to all types of granular activated carbon with re­spect to the rate and the equilibrium loading of aurocyanide. This means that the resin inventory will be lower and the size of the elution plant will be smaller

HYDROMETALLURGY OF PRECIOUS METALS RECOVERY 145

in RIP than in CIP, and this impacts favourably on both the capital and the operating costs of the process.

(2) Resins are eluted at ambient pressure and at temperatures of no more than 60°C, whereas carbon elution has to be carried out at much higher tem­peratures, preferably around 120-130°C, in a pressure vessel. This, too, re­sults in lower capital and operating costs for RIP than CIP.

(3) Activated carbon requires regular thermal reactivation to remove ad­sorbed organic material-a step that is unnecessary with resins. The cost of the equipment required for carbon reactivation can be a significant propor­tion of the total capital cost of a CIP plant, particularly for smaller operations. Resin elution/regeneration costs, on the other hand, are fairly insensitive to the scale of the operation (because the major costs are direct costs, such as the consumption of chemicals) and this indicates that RIP may find its initial market niche in smaller operations.

(4) Another potential capital cost advantage of resins over carbons stems from the fact that RIP adsorption tanks operate efficiently at high concentra­tions of resin in the pulp (20-30% by volume), with no loss in interstage screening duty, and with no increase in resin breakage. CIP plants, on the other hand, generally operate with carbon concentrations of no more than 3-6% by volume. Therefore the adsorption tanks in RIP can be at least 5 times smaller than in CIP, with no loss in metallurgical efficiency, and this contrib­utes to lower fixed costs.

( 5) Resins do not appear to be poisoned by organic species such as flota­tion reagents, machine oils and lubricants, or solvents, all of which can se­verely inhibit the loading of gold on activated carbon. Similarly, species such as hematite and shales or clay-type minerals depress the loading of gold on carbon, but have little effect on resin. Higher gold loadings should be achiev­able on resins than on carbon in situations where any of these species are present in the pulp, and this again reduces the size of the resin elution plant.

( 6) Resins can be used to co-extract other metal cyanide complexes effi­ciently from the gold leach solution, such as those of cobalt, copper, nickel, zinc and iron. Therefore, the technology can be adapted to produce environ­mentally acceptable tailings, whilst at the same time presenting an opportu­nity for recycling excess cyanide to the leaching circuit. This capability is being researched at several laboratories and, if successful, could open the way to treat a variety of deposits that presently cannot be economically processed, because of high cyanide consumption.

(7) Another important consideration in favour of resins is that much can be done during the synthesis of resins to tailor-make them for specific appli­cations, and there has been notable success recently, at Mintek in South Af­rica, in the development of gold-selective resins [45,46].

The many advantages of resins have been amply vindicated in the only op­erating RIP plant in the western world, at the Golden Jubilee Mine in South

146 C.A. FLEMING

Africa. The conversion of that plant from carbon-in-Ieach to resin-in-pulp transformed the operation from near bankruptcy in 1987 to an extremely profitable operation today [47,48].

Since converting to RIP, the mill capacity at Golden Jubilee has been dou­bled, and overall gold recovery has improved from 65% to 85%, resulting in an increase in gold production of more than 200%. What is important from an economic point of view, is that this has been achieved in RIP adsorption tanks that are one-sixth the size of the old CIL tanks, and with a resin inven­tory less than half the old carbon inventory. Operating costs in elution and regeneration have been contained by loading the resin to gold concentrations of 4000-6000 g t- I, from a feed solution value of less than 1 ppm. By con­trast, gold loading on the carbon seldom exceeded 1500 g t- I

, and was usually less than 1000 g t - I. Resin losses due to breakage, at about 8-10 g per ton of ore treated, are also considerably lower than carbon make-up in the old CIL plant (50-100gt- I ).

The rather spectacular improvement in performance obtained with resins in this particular application is, to a large extent, a consequence of the fairly unique feed material being treated, which is heavily contaminated with nat­ural organics (humates and fulvates) and is very viscous owing to its high clay content. There may not be many other deposits that will respond as fa­vourably to RIP relative to CIP, as has the Golden Jubilee deposit, although it has been shown that the efficiency of gold extraction with resins is generally similar to or slightly superior to that of carbon [39]. Moreover, it would seem that, even in those cases where there are no clear-cut advantages for one pro­cess over the other as far as extraction efficiency is concerned, RIP processing still deserves serious consideration during flowsheet evaluation, because of the potential for lower capital and operating costs.

The potential disadvantages associated with the use of resins for gold ex­traction from pulps are mostly concerned with the engineering aspects of the process. Examples include:

(1) Resin beads are smaller than carbon granules and, consequently, both the pre-screening and the interstage screening duties must be performed on finer screens in RIP than in CIP. Since this is already a bottleneck area on many CIP plants, it could present a major engineering challenge on large RIP plants.

(2) The physical strength of resins and their resistance to attrition in the adsorption tanks has been an unknown factor until relatively recently. Since resins are considerably more expensive than activated carbon, high resin losses in RIP would place an intolerable cost on the process. The experience at Golden Jubilee suggests that this is unlikely to be a serious problem on a well­designed and managed RIP plant.

(3) Resins are less dense than activated carbon and tend to accumulate near the surface of the pulp unless it is efficiently agitated. The best methods

HYDROMETALLURGY OF PRECIOUS METALS RECOVERY 147

of mixing and pumping resin/pulp slurries without causing excessive resin breakage need to be established.

( 4) Currently available, commercial, anion-exchange resins are less selec­tive than activated carbon for aurocyanide over the base-metal cyanides that are prevalent in cyanide leach liquors. These complexes can load as strongly as the aurocyanide ion onto anion-exchange resins, and reduce the capacity of the resins for gold. This disadvantage is, of course, very site-specific, but there are many examples of cyanide leach liquors that contain high concen­trations ( > 100 ppm) of copper or zinc, and resins will not compare favour­ably with carbon in these situations. The treatment of solutions and pulps with these compositions will have to await the commercialization of gold­selective resins.

Heap leaching Although the heap leaching process has been used for many years for the

recovery of copper and uranium from oxidized ores, the application of the technique to low-grade gold and silver ores is a relatively recent innovation that was developed by researchers at the United States Bureau of Mines in the late 1960s [49,50]. By opening up vast reserves of low-grade ores to eco­nomic exploitation, heap leaching has probably contributed more than any­thing, apart perhaps from the CIP process, to the surge in gold production since the early 1980s. The process is very low cost [51] and flexible, being well-suited both to small (5-10 t d- I

) and large-scale operation (10,000 t d - I ). Depending on the permeability of the rock being treated, it can be pro­cessed with or without crushing. If there is a large amount of clay in the ore, the fines are usually separated and agglomerated before being piled on the heap. The host rock must be porous and permeable or of small grain size, and the gold and silver must be liberated and non-refractory. Refractory ores and ores containing species that consume excess cyanide (oxidized sulphides of As, Sb, Zn, Fe and Cu, for example) or oxygen (pyrrhotite, organics, etc.) are not well-suited to heap leaching, because of difficulties in controlling the leaching chemistry within the heap.

In essence, the process consists of piling ore to a given height (usually 3-10 m) on a sloping impermeable bed, and spraying dilute cyanide solution onto the top of the heap. After flowing down through the heap, the leach solution is directed along the sloping impervious floor ofthe heap to a collection pond. The pregnant solution, which usually contains 1-3 ppm gold, is pumped through carbon columns to recover the gold, and flows into a second pond where the cyanide strength is restored prior to recycling to the top of the heap. On some heap leach operations, gold is recovered by zinc cementation but, as pointed out previously, carbon adsorption is usually more effective for the treatment oflarge volumes oflow-grade solution.

Recycling of the leach solution to the heap is continued until the gold con-

148 C.A. flEMING

centration in the solution from the bottom of the heap drops to an economic cut-off level, usually about 0.05-0.1 ppm. The kinetics of leaching are ob­viously slower than in a conventional leaching plant, firstly because of the generally larger ore particle-size, and secondly because of the slow rates of percolation of solution and diffusion of cyanide and oxygen. The leaching cycle for a heap is therefore usually several weeks, and can stretch to months, and overall gold recovery can be 20-30% lower than in a milling/agitation leach process.

Once leaching is complete, the barren heap is either discarded to a tailings disposal area if the same leach pad is to be used over and over again, or it is left in place, with a new pad being constructed for the next batch of ore. In certain instances, where the physical characteristics of the ore warrant it, heaps can be built one on top of another, as often practised in copper heap-leach operations. Closure of heap-leach operations is becoming an important envi­ronmental issue, and it is vital to ensure that runoff from abandoned heaps is isolated from surface and ground water sources, at least until seepage of cya­nide from the heap ceases.

Refractory gold and silver ores

General By definition, a refractory ore is one that does not respond well to direct

cyanidation. The actual leach efficiency required to meet this criterion is somewhat arbitrary, although most people in the industry are using a level of 80% recovery, or less, to categorize a refractory ore. Whatever the definition, refractory ores have one thing in common: in order to achieve a reasonable return on investment, it is generally necessary to pretreat the ore in some way to improve gold recovery. The reasons for refractoriness and the different methods of pretreating a refractory ore will be examined and evaluated in this section. Final gold recovery from the pretreated ore is by cyanidation fol­lowed by CIP, CIL or zinc cementation and, in this regard, the flowsheet is no different from one that would be used for a free-milling ore, so these aspects will not be discussed.

The recovery of precious metals from refractory ores has received much attention in the last 10 years as more and more ore bodies are found that do not respond to direct cyanidation. This trend is likely to continue as the oxi­dized free-milling gold reserves close to the earth's surface become mined-out and depleted. Underlying many of these oxidized ore bodies are sulphidic zones that have, until recently, been left in tlie ground owing to their poor response to direct cyanidation.

In response to this trend towards increased refractoriness with existing ore bodies and most new discoveries, there has been a surge in activity amongst research organizations, mining, engineering and consulting companies, to-

HYDROMETAllURGY OF PRECIOUS METALS RECOVERY 149

wards the development of technically and economically feasible pretreatment routes for these ores. These efforts have resulted in the development of sev­eral new processing options and the number of plants that have been built to treat refractory ores is increasing rapidly.

There is an increasing awareness today of the great variability between re­fractory ores, stemming from differences in the mineralogical, metallurgical and chemical properties of these ores. One consequence of this is that it is unlikely that anyone pretreatment route will emerge as a universal process for all refractory ores. To appreciate the validity of this comment, it is neces­sary to consider the mineralogical characteristics of a refractory ore. In very general terms, refractoriness is caused by one or both of the following:

( 1 ) Physical encapsulation of very finely disseminated gold particles (usu­ally less than a few micrometres in size) within a mineral that is unreactive and impervious in cyanide solution.

(2) Chemical interference by one or more constituents of the ore with the cyanide leaching process.

In the mechanism of physical encapsulation, the dominant mineral is arse­nopyrite, which exhibits a strong association with finely disseminated gold in many ore bodies. Other sulphide minerals that are known to trap gold are pyrite, pyrrhotite, realgar and orpiment. In order to liberate the gold from this type of refractory ore, it is generally necessary to chemically oxidize the sul­phides. Until 10 years ago this was done by roasting in air, which resulted in the oxidation of the sulphides to sulphur dioxide. More recently, pressure leaching has been developed to an industrial process and, in most instances, is now preferred to roasting. In this case, sulphide is oxidized to sulphate, most of which reports to the aqueous phase as sulphuric acid. Bacterial leach­ing has been developed to an industrial process on a small scale, and also involves oxidation of sulphides to sulphuric acid in solution. These three pro­cesses are discussed in more detail in the following sections. Oxidation of sulphides with nitric acid has been extensively tested and has a number of attractive features but, apart from one refractory silver ore operation [52] has not been developed to an industrial scale. In addition to these chemical oxidation methods, it is also possible to liberate gold trapped in sulphides physically by grinding the mineral to extremely fine (less than 10 Jlm) parti­cle size. Despite the relatively high energy consumption, this method is con­siderably cheaper than any of the chemical oxidation methods. The process is being extensively tested and marketed in South Africa and Australia, but will always be limited in its applicability to refractory ores containing gold parti­cles that are bigger than a few microns. Unfortunately, these seem to be fairly few and far between.

There are two important mechanisms in which gold recovery is limited through chemical interference from the ore body. The first results from the presence of species that consume cyanide or oxygen-the two essential ingre-

150 C.A. FLEMING

dients for effective gold dissolution. In this case, the simplest expedient is to increase the concentrations of cyanide and oxygen until these species have completely reacted, but this is often not economically viable. The main cul­prits are pyrrhotite and copper minerals such as malachite, azurite, chalcocite and, to a lesser extent, chalcopyrite. In fact, any base metal, such as copper, nickel, zinc or cobalt, will react strongly with cyanide if it is present in an oxidized ore body. The second mechanism is associated with the presence of graphitic, carbonaceous matter in the ore. This is a major component of many refractory ores, particularly in the State of Nevada, USA. Carbonaceous ores reduce gold extraction by virtue of their ability to re-adsorb gold once it has been leached as gold cyanide, a phenomenon that is known as preg-robbing in the industry. The initial metallurgical solution to this problem was the use of chlorine to deactivate the carbonaceous matter under strongly oxidizing conditions [53], and this approach is still followed on certain feed materials at the Carlin and Jerritt Canyon Mines in Nevada. However, for many ores the consumption and cost of chlorine are prohibitive and alternative process­ing strategies have been sought. Blinding and deactivation of the carbona­ceous minerals with an organic compound such as kerosene is a cheaper alter­native to chlorine and has been tested with many ores. It is generally not an effective approach, however. Pressure oxidation is not an obvious choice in these situations as the carbonaceous material has the potential to become even more active as an adsorbent for gold cyanide after this treatment. The effect of preg-robbing can be partially overcome by the use of CIL rather than CIP, provided the activated carbon used in CIL is more "active" than the carbon­aceous matter in the ore. The rate of equilibration of gold between the car­bonaceous matter and the activated carbon can be very slow, however, and this approach is seldom completely satisfactory. Obviously, the complete elimination of graphitic carbon is the most desirable solution. Attempts to achieve this by mineral dressing techniques (gravity, flotation, etc.) have only ever been partially successful. In many instances roasting of the ore will achieve this objective, but the operating parameters have to be carefully chosen, even in this case, to ensure the graphitic carbon is burnt off and is not converted to a more active form.

Apart from the selection of the best processing route to adopt, the other decision that the developer of a refractory gold operation needs to make is whether or not to preconcentrate the gold-bearing ore. This is usually by flo­tation of either arsenopyrite, pyrite, or both, and gravity concentration is also possible, although the opportunities in this area are more limited. Preconcen­tration should meet two objectives for it to b6 economically justifiable: the gold recovery to the concentrate should be sufficiently high (probably> 90%) so that the tailings can be discarded without further treatment, and the mass ratio of concentrate to tailings should be low (less than 1: 10). The other im­portant factor to be aware of is that the major capital and operating costs of

HYDROMETALLVRGY OF PRECIOUS METALS RECOVERY 151

pretreatment, whether by roasting, pressure leaching or bacterial leaching, stem from the amount of sulphide sulphur that must be oxidized, not the concen­tration of sulphide. Hence, the pretreatment costs will be similar for ore and concentrate, and the major benefits of preconcentration only occur further downstream as a result of smaller cyanidation and CIP plants.

The final point of general interest in relation to the treatment of refractory gold ores relates to the choice of the most appropriate lixiviant to leach the precious metals from the pretreated ore or concentrate. Cyanide has been used almost universally up to now, and will continue to be the preferred choice in most cases. However, there has been a great deal of research into alternative leaching chemistries in the last 10 years. The incentive for this work has stemmed largely from concerns relating to the ability of the industry to meet more and more stringent environmental regulations. The validity of this fo­cus is questionable. Firstly, there is every indication that the industry will be able to meet the most stringent limits with technology that is currently avail­able for treating cyanide effiuents. Secondly, it is by no means proven or es­tablished that the alternative leach liquors will be any easier or cheaper to treat and render environmentally benign.

However, the treatment of refractory ores is a potential niche application where alternative lixiviants could become important in the future. The reason for this lies in the fact that preoxidation of sulphide ores, whether by pressure or bacterial leaching, involves the generation of sulphuric acid and ferric ions in solution. In an alkaline cyanide leaching process, these ions must always be neutralized, and the acidic liquors often need to be separated and washed from the solids (to reduce cyanide consumption). This would not be neces­sary when leaching gold with either thiourea or thiocyanate, however, as both sulphuric acid and ferric ions are, in fact, important ingredients in the overall leaching reactions in these two systems. This particular application therefore warrants investigation.

Roasting By far the most commonly used technique for the recovery of gold from

refractory ores is roasting of flotation concentrates prior to cyanidation. When arsenopyrite is present, a two-stage process is usually applied-a non-oxidiz­ing first stage roast at 400-450°C to remove the arsenic as volatile arsenic trioxide, followed by an oxidizing roast at 650-750°C, to produce hematite and sulphur dioxide. Roasting results in a reasonably good recovery of gold from refractory materials, provided close control is maintained over the roasting atmosphere and temperature, since the calcine produced must be sufficiently porous for even the finest gold particles to be accessible to the cyanide leach solution.

In many instances the presence of carbon causes problems in roasting, since the high temperature required to burn off the carbon yields a calcine that is

152 CA. FLEMING

not sufficiently porous to expose all of the gold. Nevertheless, roasting should be the best way of dealing with carbonaceous refractory ores as it is the only method of potentially eliminating all of the carbon. The conditions have to be carefully controlled, however, since any carbon that is not burnt off will be even more active in the calcine than it was in the original ore. The presence of antimony in a roaster feed also has a deleterious effect on subsequent gold recovery, and this is probably associated with the formation ofa low melting­point, impervious, glassy phase that traps gold.

Roasting has developed from fixed-bed hearth and rotary furnaces to fluid­ized-bed type furnaces. The development of fluid-bed roasting has resulted in significant improvements to the process, as a result of more homogeneous reaction conditions and better control of the temperature, atmosphere and residence time of reactants in the roaster.

More recently, there has been a trend towards the roasting of whole ores rather than concentrates-a development that has been accelerated by two important technical innovations, the circulating fluid-bed roaster [54] and the two-stage oxygen roasting process ofIndependence Mining Company [55].

The main technical challenges in roasting lie in meeting the stringent envi­ronmental limits for gaseous discharge, particularly in the treatment of ar­senic sulphide ores. It is generally believed to be necessary to volatilize the arsenic in an ore in order to achieve high gold recovery in the subsequent cyanidation process, although unconfirmed reports indicate that this may not necessarily be true in the oxygen-whole-ore roasting process. Arsenic is vola­tilized as arsenic trioxide, and must then be very efficiently condensed and filtered from the off-gas. This dust is too impure for direct sale to the chemical or pharmaceutical industries, and the cost of refining cannot be justified in most cases. The material is highly toxic, so storage is also an expensive option.

All sulphides are oxidized to sulphur dioxide during roasting. The produc­tion of sulphuric acid from the off-gas is technically feasible if the sulphide grade in the roaster feed is high, but the capital cost of an acid plant is signif­icant, and cannot be justified unless there is a market for the acid nearby. Therefore, the sulphur dioxide is generally neutralized by scrubbing the off­gas in lime, and this can be a significant cost factor, particularly in the roast­ing of whole ore with a high (> 3%) sulphur content. Lime or limestone can be pre-mixed with the ore prior to roasting, and this traps some of the sulphur dioxide in the roaster bed, reducing the demands on the off-gas scrubbers. This option needs to be carefully tested and optimized, however, as there are indications that gold recovery may be adversely affected. The practice of dis­charging untreated gas to the atmosphere is no longer an option on a new plant, anywhere in the world.

Another parameter that needs to be investigated for each ore or concentrate type is the degree of sulphide oxidation that is required. Some ores appear to require as little as 40-50% oxidation to achieve optimum gold recovery. In-

HYDROMETALLURGY OF PRECIOUS METALS RECOVERY 153

creasing the degree of oxidation may lead to only marginal improvement in gold recovery, but produces more sulphur dioxide to be neutralized. The var­ious other sulphide oxidation processes discussed below are no different in this regard.

Pressure leaching There has been a shift in direction in recent years in the treatment of re­

fractory gold ores or concentrates from the traditional roasting method to pressure leaching. This is the result of two factors: firstly, the inability of first­generation roasting plants to meet stringent new environmental limits on the discharge of sulphur dioxide and arsenic trioxide to the atmosphere; and sec­ondly, the fact that gold recovery from roasted calcines is generally inferior to that achievable after pressure leaching. This latter factor is especially impor­tant in the treatment of arsenopyritic ores, where the gold is frequently so finely disseminated that it is not rendered accessible to cyanide, even though a porous calcine is produced. In pressure leaching, the decomposition prod­ucts of the oxidation of both pyrite and arsenopyrite are dissolved species (which then reprecipitate under certain conditions), so any gold occluded in these mineral phases, no matter how fine, is fully exposed during oxidation. The pyrite and arsenopyrite decomposition reactions can be represented as follows:

2FeS2 + 15/202 +H20-+Fe2(S04h +H2S04

2 FeAsS + 702 +H2S04 +2H20-+Fe2(S04h +2H3As04

(16)

(17)

Under the conditions prevailing in an autoclave, 60-95% of the iron repre­cipitates as a basic ferric sulphate (Fe(OH)S04) or as a jarosite, (MFe3(OH)6(S04b where M=H+, Ag+, 1/2 Pb2+) while a significant Portion (80-98%) of the arsenic reprecipitates as the very stable ferric arse­nate complex (FeAs04).

A number of large pressure oxidation plants have been successfully com­missioned in the last 6 years (Table 3), and these have demonstrated to the metallurgical industry that the process is rugged and highly efficient for the pretreatment of refractory gold ores. There are also several plants under con­struction, and many more in various stages of metallurgical testing, engineer­ing design or feasibility assessment. Pressure leaching is undoubtedly the preferred pretreatment route for non-carbonaceous, refractory gold ores at the present time.

The first step in a pressure leaching operation will generally involve precon­ditioning with sulphuric acid to neutralize the carbonates (calcium or mag­nesium) that are often a constituent of refractory gold ores. If this is not done prior to pressure leaching, the acid generated in the autoclave, from the sul­phide oxidation reaction, will decompose the carbonates and produce carbon

154 C.A. FLEMING

TABLE 3

Pressure leaching operations for the treatment of refractory gold ores

Project Status Plant capacity Country (t d- ' )

McLaughin Operational 3,000 ore USA Goldstrike phase I Operational 1,500 ore USA Getchell Operational 3,000 ore USA Mercur Operational 800 ore USA Sao Bento Operational 240conc. Brazil Campbell Red Lake Operational 70conc. Canada Porgera Operational 3,500conc. Papua New Guinea Goldstrike phase 2 Construction 10,000 ore USA Con Construction 100conc. Canada

dioxide in the autoclave. As a result, the partial pressure of oxygen in the autoclave is reduced and this, in turn, lowers the rate of sulphide oxidation. Depending on the relative proportions of sulphide and carbonate in the ore or concentrate, the overall acid balance may be either positive or negative. If it is positive, the acid generated in the autoclave can be recycled to the pre­conditioner by separating the solids and the liquids in the autoclave dis­charge. If it is negative, it will be necessary to import acid to decompose the carbonates. This is obviously undesirable, and would be an incentive to treat concentrate rather than ore, since carbonates can usually be rejected quite effectively to the tailings during flotation. An alternative approach in a situ­ation where the sulphide grade is low and the carbonate grade is high, is to operate the pressure leaching reaction under alkaline conditions, and this is the processing route that has been followed at the Mercur Mine in the USA. The rate of sulphide oxidation is slower under alkaline conditions than it is in acidic solution, and therefore it is necessary to employ longer residence times and higher temperatures to achieve equivalent gold liberation.

Before pumping slurry to the autoclave, the solids density must be adjusted to an optimum value, which is a function of the sulphide concentration in the solids. This can be done before, during or after acid-preconditioning, and could involve either thickening or dilution of the slurry. The relationship between solids density and sulphide concentration is extremely important, since it in­fluences both the heat balance in the autoclave and the ultimate concentra­tion of sulphuric acid in the autoclave solution.

The oxidation of sulphide with oxygen in the autoclave is an exothermic reaction, and the concentration of sulphide required for an autothermal re­action can be calculated from the heats of reaction for the different sulphide minerals. For example, Mason [56] has calculated that a sulphide grade of 5.6% would be needed for autothermal reaction with an arsenopyritic ore, or

HYDROMETAlLURGY OF PRECIOUS METALS RECOVERY 155

12.9% with a pyritic ore, assuming a solids density of 40% and an operating temperature of 220°C. Mixed pyrite/arsenopyrite ores would possess heat­generating capacities between these values. Clearly, therefore, if the concen­tration of sulphide in the feed is higher than these values, some degree of cooling or dilution may be necessary in the autoclave, to maintain tempera­ture. If the grades are lower, which is usually the case in the treatment of whole ore, it is necessary either to preheat the feed slurry prior to pumping into the autoclave, or to inject steam under pressure into the autoclave. Pre­heating of ore is achieved by recovering heat from the autoclave discharge and recycling it to the feed, via a series of heat exchangers operating counter­current to the flow of slurry.

The importance of sulphide grade in relation to sulphuric acid concentra­tion stems from the fact that the oxidation of elemental sulphur (an inter­mediate product in the overall sulphide oxidation reaction) is retarded at higher sulphuric acid concentrations [57]. This results in a build-up of ele­mental sulphur in the autoclave product, which can cause both physical (ag­glomeration) and chemical problems (high cyanide consumption). Interest­ingly, the rate of oxidation of sulphide (to elemental sulphur) apparently increases with acidity at low sulphuric acid concentrations, and is then rela­tively constant at high acidities [57], so gold liberation is not a problem un­less it becomes re-encapsulated in elemental sulphur. In practice, sulphuric acid concentrations are usually maintained in the 20-60 g 1- 1 range, and it is only at concentrations of 100 g 1-1 or more that elemental sulphur formation can become severe. It is not always possible to predict the sulphuric acid con­centration directly from the solids density and the sulphide concentration in the feed, because even if all of the sulphide is oxidized to sulphate, it is usually distributed between sulphuric acid in solution and insoluble sulphate com­plexes in the solids (jarosites, for example), with the distribution influenced by reaction conditions in the autoclave. Laboratory testing is therefore nec­essary to determine optimum reaction conditions, although an empirical re­lationship has been developed by Conway and Gale [58]:

solids density = (0.3[S2-] +0.825)-1 (18)

which can reduce the amount of testing required. The important operating parameters to control in the autoclave are tem­

perature, oxygen partial-pressure, redox potential, residence time and, as mentioned earlier, sulphuric acid concentration. Temperature, oxygen par­tial-pressure and acid concentration all influence the rate of sulphide oxida­tion, whereas redox potential is used to monitor the extent of oxidation. The optimum conditions required for a particular refractory ore are established in laboratory batch and continuous autoclave tests, and the results obtained on a small scale in the laboratory generally reproduce extremely well in a full­scale operation. The size of the autoclave is a function of the residence time

156 C.A. FLEMING

required for sulphide oxidation and, since this can dramatically influence the total capital cost of the facility, and ultimately the economic feasibility of the process, it is important to maximize the rate of oxidation. The temperature in the autoclave ranges from a low of 180°C at the McLaughin plant to 225°C at the Goldstrike plant, and oxygen partial pressures range from 200 to 550 kPa. Under these conditions, residence times in the autoclave range from 60 to 120 min for pyritic and arsenopyritic refractory ores and concentrates.

The discharge from the autoclave is cooled in a series of flashing steps in which, depending on the heat balance in the autoclave, heat is either re­covered and recycled, or is vented as steam to the atmosphere, through a suit­able scrubber. Depending on the acid concentration in the autoclave liquor, and whether or not acid-recycling to the feed is necessary, the autoclave dis­charge slurry will either be neutralized directly prior to cyanidation and CIP (as practised at McLaughin and Goldstrike), or the autoclave liquor will be separated from the solids by filtration or CCD. This is an expensive option because the solid-liquid separation equipment must all be acid-proof, but it does benefit the down-stream cyanidation and CIL or CIP operations in sev­eral ways. Firstly, separation of the autoclave liquor removes base metals that would otherwise consume cyanide and could also cause problems in CIP and elL through formation of slimy hydroxide precipitates. Secondly, the re­moval of most of the acid with the autoclave liquor reduces the amount of gypsum formed during neutralization of the slurry prior to cyanidation. This reduces the viscosity of the slurry and results in improved CIP JCIL perform­ance. Finally, the feed to cyanidation is a lot cooler after solid-liquid separa­tion than when the autoclave discharge is neutralized directly, and the lower temperature also benefits cyanidationjCIL performance significantly.

Bioleaching With the world-wide interest in low-grade refractory gold ores, the "new"

technology of bacterial heap leaching is being actively investigated by several mining companies. The greatest interest, however, is focussed on the treat­ment of high-grade concentrates, by bacterial oxidation of finely-milled solids in stirred-tank reactors. The fundamental difference between the refractory gold applications and those involving copper and uranium minerals lies in the fact that the metal of interest, gold, does not dissolve during bacterial leaching but, as in roasting and pressure leaching, is liberated for subsequent recovery by cyanidation.

The chemical reactions involved are the same as those shown previously for pyrite and arsenopyrite in pressure leacHing (eqs. (16) and (17». Ar­senic and iron dissolve, liberating gold, but unlike the situation in an auto­clave, they do not reprecipitate under the ambient conditions prevailing in a bioleach reactor. The build-up of iron, arsenic and sulphuric acid (and tem­perature) must be controlled, in fact, because they limit the rate of bacterial

HYDROMETALLURGY OF PRECIOUS METALS RECOVERY 157

reproduction. For this reason, bioleaching of sulphide concentrates is con­ducted at low solids densities, usually around 10-15%. The pH needs to be controlled in the 1.5-2.0 range, and the excess sulphuric acid generated by oxidation of sulphides is neutralized by adding lime slurry to the bioleach reactor. The temperature must be controlled in the 35-40°C range for Thioh­acillus ferrooxidans and, because of this limitation, some research work is being done with su/folobus, a bacterial strain preferring higher temperatures. Iron and arsenic levels must also be controlled.

A typical plant treating concentrate will receive a feed containing anywhere from 10% to 30% sulphide and up to 25% arsenic. The iron content will de­pend on the relative amounts of pyrrhotite, pyrite and arsenopyrite, but there should be at least a stoichiometric equivalent of iron to arsenic to ensure com­plete precipitation offerric arsenate rather than calcium arsenate, during neu­tralization after bioleaching. The stability of the precipitate apparently in­creases with increased iron to arsenic ratio, and a molar ratio of 4: 1 is recommended [59]. The feed to bioleaching may be pretreated with fresh or recycled acid if it contains more carbonates than sulphides. The oxidation tanks are usually configured in four stages, with additional tanks employed in the first stage to accommodate the increased oxidation rates and aeration re­quirements in the initial phases of the reaction. Overall residence time in the bioi each reactors is of the order of 3-4 days. The air supply to the reactor must supply both the oxygen needed to oxidize the sulphides and the agita­tion needed to suspend the ore particles, and the demand by the former is much greater than the latter, particularly in the early stages. Consequently, there has been a lot of research into the method of distributing air into the slurry and maximizing mass transfer of oxygen in the slurry. Oxygen transfer efficiency is obviously lower for a process that uses air than for one that uses pure oxygen, such as pressure leaching. However, air is preferred to oxygen as the source of oxygen, firstly because of cost, and secondly because the bacteria also require carbon dioxide to survive. A typical range of theoretical oxygen demands would be 100-500 kg of oxygen per ton of concentrate. The supply of air to the reactor will also remove some of the heat generated by the oxi­dation reaction, but in practice additional heat removal (heat exchangers) will probably be necessary to control the temperature at -- 40 ° C.

Nutrients to stimulate bacterial reproduction are added with the feed to the first reactor. The optimum nutrients and their dosage need to be established for each application, but will usually consist of a mixture of ammonium, po­tassium and phosphate ions. Some of these may occur naturally in the ore, and the balance can generally be supplied in the form of a commercially avail­able fertilizer.

A solid-liquid separation step is generally carried out after bioleaching be­cause of the low solids density of the slurry, and thickening would probably be preferred because oxidized ores do not filter well. Part of the overflow

158 C.A. FLEMING

could be recycled to the feed as a source of make-up acid and bacteria, al­though this is not critical, because the bacteria reproduce very rapidly in the bioleach reactors. The remaining liquor would be neutralized with lime to produce the stable, basic ferric arsenate complex, (FeAs04 )Fe(OHh, and gypsum. The thickener underflow could be washed in a CCO circuit if cya­nide consumption is high, or could be treated directly by lime neutralization, cyanidation and CIP.

An important characteristic of bacterial leaching reactions is that they are autocatalytic, i.e. the bacteria are themselves a product of the sulphide oxi­dation reaction, and much of the progress that has been made in the last 20 years for the treatment of concentrates has been in the development of bac­terial strains that multiply rapidly in the aggressive conditions that prevail in the exothermic sulphide oxidation reaction. This was an essential develop­ment before even small-scale bioleaching of milled concentrate in stirred re­actors could be considered for industrial application. For example, in the early pilot plant work at the Gencor Laboratory in South Africa, it was found that a 10 day residence time was needed to achieve sufficient oxidation of an ar­senopyrite concentrate for 97% gold recovery. When this is compared with the 1-2 h residence time in a typical pressure leaching operation, some insight into the relative sizes of a pressure leaching and a bioleaching plant can be gained. After 2 years of operation of the Gencor pilot plant, the bacteria had adapted and mutated to the extent that the retention time had decreased to 4 days [60]. At that time a decision was taken to build an industrial bioleach­ing plant to treat this arsenopyrite concentrate at the Fairview Gold Mine in South Africa. The plant, which was the first of its kind in the world, was com­missioned in 1986, with a design capacity of lOt of concentrate per day. This represented 40% of the concentrate being produced at Fairview. In the 5 years that the plant has been operating, further improvements and refinements have allowed the capacity of the original plant to be increased to about 17 t d - I,

with gold recoveries typically around 95% (compared with <90% recovery from the roaster calcine at the same mine). Gencor metallurgists are suffi­ciently confident with the process that bioleach capacity is currently being expanded at Fairview to treat all of the concentrate produced, which will al­low the old Edwards roasters to be decommissioned. Gencor are also building a bioleaching plant at their Sao Bento operation in Brazil, which will probably be operated in series with the pressure leaching facility that is already there, to accommodate the planned increase in production at that mine [61 ].

Another bioleaching operation in South Africa is the 20 t d - 1 pilot plant at Vaal Reefs, which oxidizes a pyrite concentrate containing about 5 g t- I of gold, prior to uranium and gold recovery. This operation was developed through ajoint bioleaching research program involving both the Anglo Amer­ican Corporation and Mintek. A third commercial undertaking, at Tonkin

HYDROMETALLURGY OF PRECIOUS METALS RECOVERY 159

Spring in Nevada, was operated for a short while in 1990, but ran into tech­nical difficulties. This plant is currently mothballed while new partners in the venture re-evaluate the technology and the ore reserves. The treatment of concentrate in stirred reactors is also being piloted at the present time at Dickenson Mine in northwestern Ontario. This will be a small-capacity plant, and it is therefore well-suited to bioleaching because of the lower capital costs, on this scale, compared to pressure leaching and roasting. In fact, biological oxidation may well find its niche in the treatment of small tonnages of flota­tion concentrate, particularly arsenopyrite, because of the significant capital cost advantages. The other situation where bioleaching economics may com­pare favourably with other processes, is one in which only partial sulphide oxidation is needed for good gold liberation. Because of the very much slower kinetics of sulphide oxidation in a bioleach reactor, partial oxidation is far easier to control than in a roaster or an autoclave.

CONCLUSIONS

The cyanide leaching process will continue to playa dominant role in the processing of free-milling and refractory (after pretreatment) gold ores for the foreseeable future. New leaching processes may find limited application in niche areas, such as the treatment of acidic slurries generated in pressure leaching or bioleaching.

It is unlikely that any more plants incorporating solid-liquid separation (by either filtration or CCD) will be built, unless gold and silver values are very high (> 20g t- 1 combined Au, Ag). Carbon-in-pulp and carbon-in-Ieach will be the preferred processing routes for the treatment of cyanide leach slur­ries for at least the next 5 years, and possibly longer. Resin-in-pulp has many attractive features, and could gradually replace carbon-in-pulp if the process can be successfully engineered on a large scale.

The trend towards increasing refractoriness in gold and silver ores will con­tinue, and more and more gold plants will incorporate pretreatment process­ing to liberate precious metals. The choice of pretreatment will depend on ore mineralogy and site-specific economic and environmental factors. New pre­treatment technologies such as bioleaching, nitric acid leaching and ultra-fine grinding have many attractive features and will increasingly challenge the "established" processes of roasting and pressure leaching.

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160 C.A. FLEMING

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