AU2016224142B2 - Processing of sulfidic ores - Google Patents

Abstract

A process is disclosed that enables a precious metal to be recovered from a sulfidic material. The process is particularly suited to treating single- and double-refractory pyritic materials such as pyrites and arsenopyrites. The process comprises the preparation of an acidic aqueous metal halide solution that has an oxidation potential that is sufficient to oxidise the 5 sulfidic material, with the metal of the metal halide able to function as a multi-valent species during oxidation of the sulfidic material. The sulfidic material is introduced into the acidic aqueous metal halide solution and the oxidation potential is controlled such that the sulfidic material is oxidised, the metal of the metal halide is reduced to a lower valent state, but such that the precious metal is generally not oxidised. The metal of the metal halide is oxidised to a 10 higher valent state, and a residue that comprises the precious metal is separated from the solution for its subsequent recovery.

Description

Processing of Sulfidic Ores

Technical Field

Disclosed herein is a process for enabling precious metals to be recovered from sulfidic ore materials, especially a precious metal such as gold, and especially from pyritic materials. The process can be applied to so-called "single-refractory" materials (having no or low carbon content, such as arsenopyrites), and can also be applied to so-called "double refractory" materials (those having a relatively high (e.g. > about 2 wt.%) carbon content). The process can help to "unlock" the sulfidic ore, and to liberate the precious metal, to enable the material to then be subjected to conventional precious metal leaching, such as cyanide leaching.

Background Art

Globally there are significant deposits, stockpiles and quantities of sulfidic ore materials (including pyritic ores). Such materials can comprise economically desirable metals to recover, especially precious metals such as gold, silver, platinum and other platinum group metals.

Some of these materials are contaminated with difficult to process contaminants such as arsenic, antimony, bismuth or other heavy metals. Ore treatment may be further complicated when high levels of carbon are present, as carbon associates with and has a high affinity for precious metals such as gold.

Precious metals such as gold can be encapsulated in a sulfidic ore body that has, over time, become oxidised in the environment to form an oxidised upper zone overlaying a generally much larger sulfide zone. The oxide zones of such ore bodies are usually treated using conventional cyanide leaching technology. However, the precious metals in the sulfide zone are not easy to recover economically using conventional cyanide leaching technology, due to excessive cyanide consumption.

Available options for the oxidation of sulfidic materials include roasting, pressure oxidation (POx) and bio-oxidation (Biox). The POx and Biox processes employ a sulfate medium.

Roasting sulfidic ores presents significant problems due to emissions of environmentally toxic sulfur based gases (so-called SOx gases). Where arsenic is present in the ore, poisonous substances such as arsenic trioxide are produced. For these reasons international trends are to move away from roasting of sulfidic ores.

Pressure oxidation of sulfidic materials is employed to avoid the problems of roasting, but requires high pressures (typically greater than 30 bar) and relatively high temperatures (greater than 200°C). Pressure oxidation is also typically carried out in a sulfate based solution.

A two-stage Biox process can be used to solubilise the arsenic, however, the configuration of the leaching process is complex, as is the use of bio-leaching bacteria. In addition, bio-oxidation is notoriously slow. Biox processes are also vulnerable to process interruptions.

US7858056 discloses a leaching process in which a solution is prepared that has a sufficiently high oxidation potential to generate a complex halide species, and to oxidise the sulfidic material and the precious metal(s), thereby rendering the precious metal(s) soluble in the leachate. Specific separation technologies are then employed to recover the precious metal(s).

The above references to the background art do not constitute an admission that the art forms a part of the common general knowledge of a person of ordinary skill in the art. The above references are also not intended to limit the application of the process as disclosed herein.

Summary

Disclosed herein is a process for enabling a precious metal to be recovered from a sulfidic material. The process is particularly suited to treating pyritic ore materials that are contaminated with difficult to process contaminants such as arsenic, antimony, bismuth or other heavy metals, and which pyritic materials may form part of a sulfidic ore body that has an oxidised upper zone overlaying a generally much larger sulfide zone.

The process as disclosed herein enables treatment of what may otherwise be considered as a too difficult or uneconomical to treat contaminated (i.e. refractory) sulfidic materials. In this regard the process is able to break down the matrix of the contaminated material, separate the impurities from a residue that comprises the precious metal to be recovered, and to then pass a clean, purified residue (e.g. filter cake) to a conventional precious metal recovery stage (e.g. cyanide leaching).

The process as disclosed herein comprises the preparation of an acidic aqueous metal halide solution that has an oxidation potential that is sufficient to oxidise the sulfidic material, particularly that of the larger sulfide zone. The metal of the metal halide is selected such that it is able to function as a multi-valent species during oxidation of the sulfidic material.

In the process as disclosed herein typically a suitably prepared sulfidic material (e.g. a gravity or flotation concentrate from bulk sulfidic ore) is introduced into the acidic aqueous metal halide solution. In contrast to the process of US7858056, the oxidation potential of this solution is controlled such that the sulfidic material is oxidised, the metal of the metal halide is reduced to a lower valent state, and the precious metal (e.g. gold) is generally not oxidised. The expression "generally not oxidised" is employed to indicate that, although some e.g. gold may be oxidised (leached) by the solution, this is not the intention of the process. Rather, the process is operated to unlock the sulfidic ore, and to liberate the precious metal, to then enable the resultant material to be subjected to conventional precious metal leaching, such as cyanide leaching.

In one embodiment of the process, the oxidation potential is controlled to be less than an Eh of around 0.65 volts (ref. Ag/AgCl). It has been observed that, at or below this oxidation potential, the precious metal (e.g. gold) is generally not oxidised. However, in the halide environment (e.g. a mixed chloride/bromide solution), and in the presence of the multi-valent metal of the metal halide, it has been discovered that the sulfidic material is still able to be oxidised.

In one embodiment of the process, the oxidation potential is controlled to generally be greater than an Eh of around 0.5 volts (ref. Ag/AgCl). Whilst the oxidation potential may, from time-to-time, fall below 0.5 volts (e.g. to 0.45 volts or less), generally, by controlling the process conditions to ensure an Eh of around or greater than 0.5 volts, the effectiveness and rate of the oxidation are improved.

In the process as disclosed herein, the metal of the metal halide is (re-)oxidised to a higher valent state, such as by the addition of an oxidant (e.g. air, oxygen, etc) into the solution. More particularly, the conditions of the sulfidic material oxidation (leaching stage) can be controlled to provide a suitable balance/equilibrium of, inter alia, the multi-valent species such that, as the sulfidic material is being oxidised and the metal of the metal halide is being reduced to a lower valent state, the oxidant (e.g. air, oxygen) is, inter alia, re-oxidising the metal of the metal halide to its higher valent state. Thus, when this solution is separated from a residue that comprises the precious metal, this solution is ready/suitable to be recycled as the acidic aqueous metal halide solution (i.e. in preparation for further leaching).

In the process as disclosed herein, usually oxygen (typically air) is introduced into the acidic aqueous metal halide solution to, inter alia, facilitate oxidising of the sulfidic material and of the metal of the metal halide to its higher valent state.

In the process as disclosed herein, the residue that comprises the precious metal is separated from this solution. For example, this separated residue may take the form of e.g. a filter cake. This separated residue may then be passed to a subsequent precious metal recovery process (e.g. conventional cyanide leaching).

In one embodiment, the solution that is separated from this residue can, as set forth above, be recycled to enable preparation of the acidic aqueous metal halide solution. The recycled solution can, as set forth above, comprise at least some of the metal of the metal halide solution oxidised to its higher valent state.

With the process as disclosed herein, by controlling the oxidising of the sulfidic material, but leaving precious metal (e.g. gold) generally not oxidised, existing cyanide technology can then be used to recover the precious metal from the residue. This allows the process to be combined or integrated with existing cyanide-based leaching infrastructure, and can also reduce technical risk and capital expenditure through the use of such existing infrastructure.

With the process as disclosed herein, because the precious metal (e.g. gold) is generally not oxidised, there is negligible "preg-robbing" (e.g. masking) by organic carbon which may be present as a contaminant (e.g. at about 0.4%) in the sulfidic material to be treated, because the e.g. gold is generally not dissolved into the solution.

Further, with the process as disclosed herein, when the sulfidic material is contaminated with arsenic, antimony or the like, the arsenic etc can be simultaneously leached and precipitated in a leaching stage of the process (i.e. the process is able to accommodate such impurities by leaching and removal). It can be uneconomic for existing cyanide leach technology to leach such contaminated sulfidic materials.

In addition, it has been observed that the multi-valent metal in a halide environment better enables e.g. air/oxygen to be utilised to oxidise contaminated pyrites such as arsenopyrites, etc through to e.g. scorodite (FeAsO 4) and elemental sulfur.

The process can be operated in a closed loop or recycle mode with attendant economic benefits (e.g. simplicity, low energy consumption, preservation of mass balance etc.). The process can be applied to recover a precious metal such as gold from a range of sulfidic materials, including otherwise difficult to treat ores and concentrates, and including double refractory materials having relatively high carbon content (e.g. carbon-containing arsenopyrites).

In one embodiment, the process may comprise a first leaching stage. In addition to air (oxygen), an acid (typically sulfuric acid, or may be hydrochloric acid or another acid which does not interfere with the process chemistry) may be added to the first leaching stage to ensure an optimum Eh and pH for oxidising of the sulfidic material.

Typically the sulfidic material comprises a pyritic material which may be oxidised in the first leaching stage (i.e. sulfuric acid and air are advantageously optimised to the leaching of many typical pyrites). Other oxidants such as chlorine gas, hydrogen peroxide, etc may also be added to the first leaching stage.

In one embodiment, the process may further comprise a second leaching stage in which the solution conditions are controlled (e.g. the maintenance of a suitable Eh and the control of pH (i.e. raised and monitored relative to the first leaching stage)) to maintain iron in its ferric state, and to maintain a scorodite precipitate (FeAsO 4) so as to facilitate its removal, including as an Fe 2 03 precipitate.

In the process as disclosed herein, each leaching stage may be operated co- or counter currently and, in this regard, each stage may comprise one or more vessels. The entire solution from the first leaching stage can be fed to the second leaching stage.

In one embodiment, when sulfuric acid is added to the first leaching stage, calcium in solution can precipitate as calcium sulphate (CaSO4. To replace the calcium, and also to help in maintaining suitable Eh and pH, a base such as calcium carbonate can be added to the second leaching stage. Oxidation of, in particular, pyritic materials can also produce excess sulfuric acid, whereupon the addition of a base such as calcium carbonate can neutralise such excess acid.

In one embodiment, when the pyritic material comprises arsenopyrite, the oxidation potential in the first leaching stage may be controlled to leach arsenic into solution, and the solution pH may be controlled such that, once leached, the arsenic precipitates as ferric arsenate. In the second leaching stage the solution pH may be controlled to maintain arsenic as a ferric arsenate precipitate, and to also maintain an Fe 2 03 precipitate. Ferric removal may be controlled by regulating limestone addition so as to maintain some iron in solution, which in turn can prevent cupric copper precipitation (i.e. because iron precipitates at a lower pH than copper and buffers the pH whilst it precipitates, thereby acting as a safeguard against copper precipitation). Also, iron in solution can act as one of the multi-valent species, as described below.

Air (oxygen) or other oxidants such as chlorine gas, hydrogen peroxide, etc can also be added to the second leaching stage to control and ensure optimal solution conditions (Eh and pH) to maintain the precipitated forms of Fe 2 03 and FeAsO4, and to regenerate the metal of the metal halide to its higher valent state.

When the pyritic material comprises arsenopyrite, the pH in the first leaching stage can be less than 1 but greater than about 0.5 so as to precipitate the arsenic immediately after it is leached. In this embodiment, the solution temperature can be in the range of about 60-105C. In this embodiment, the pH in the second leaching stage can be less than -2.2 but greater than about 1 so as to maintain the ferric oxide and arsenate precipitates. In this embodiment, the solution temperature in the second leaching stage can be in the range of about 900 C to 1050 C.

In one embodiment, the residue produced by the process (i.e. that comprises the precious metal) can be separated from the solution in a solid-liquid separation stage that comprises a thickening stage and a wash/filtration stage. Other separation methodologies may be employed such as solid/liquid settling, solution evaporation, centrifugation, etc. The product of the wash/filtration stage can comprise filter cake.

In one embodiment, in the thickening stage, a portion of the overflow from e.g. a thickener can be recycled to provide the acidic aqueous metal halide solution. Another portion can be passed to an impurity removal stage.

In one embodiment, in the impurity removal stage, impurities can be precipitated, filtered and disposed of. This can prevent their build up in the process solution. For example, lime and calcium carbonate can be added to the overflow portion to form a hematite/gypsum precipitate, which can then be filtered and disposed of From the impurity removal stage a portion of filtrate can be recycled to enable preparation of the acidic aqueous metal halide solution, and another portion of the filtrate can be passed to the wash/filtration stage.

In one embodiment, underflow solids from the thickening stage can be passed to the wash/filtration stage. In the wash/filtration stage the solids can be washed and filtered before they are passed to precious metal recovery (e.g. cyanide leach). For example, the wash/filtration stage can employ a brine wash and/or a hot water wash to remove metal salts that may be entrained in the residue. The wash/filtration stage can, in this regard, employ a belt filter to enhance metal salts removal.

The filtrate from the wash/filtration stage can be recycled to enable preparation of the acidic aqueous metal halide solution, with washed solids being passed on to precious metal recovery (e.g. to a conventional cyanide leaching process).

In the process disclosed herein, the residue that comprises the precious metal may contain elemental sulphur, which can interfere with precious metal recovery (e.g. it can consume cyanide). Thus, prior to precious metal recovery, the residue may first be subjected to a sulphur removal stage. The sulphur removal stage may comprise one or more of direct oxidation, thermal solvent extraction, flotation. Direct oxidation can oxidise the elemental sulphur to sulphate for easy removal. Thermal extraction of the elemental sulphur can occur via direct dissolution in organic solvent(s), such as xylene, tetrachloroethylene and/or kerosene.

The sulphur removal stage may occur within or as part of leaching (e.g. in, or as part of, the first or second leaching stages). Alternatively, the sulphur removal stage may occur after the residue (e.g. filter cake) has been separated from the solution (e.g. after the wash/filtration stage).

In the process disclosed herein, a multi-valent species (i.e. the metal of the metal halide) is typically selected that has both a relatively high oxidation state to participate in oxidation of the sulfidic material and a relatively lower oxidation state to which it is reduced during oxidation. Advantageously, the multi-valent species can then be regenerated to its relatively high oxidation state (e.g. in the second leaching stage), whereafter the regenerated multi-valent species can be recycled to participate in further oxidation. Advantageously, the regeneration of the multi-valent species can occur during each leaching stage so that the regenerated species can participate in oxidation and be recycled as part of a preferred closed loop or recycle mode of the process, with the attendant economic benefits (e.g. preservation of mass balance, simplicity, low energy consumption, etc.).

In one embodiment, the metal of the metal halide solution may comprise copper and/or iron. Both metals can be present in a pyritic material to be treated in accordance with the process as disclosed herein. Further, and so as to "seed" the process, copper may be added to the acidic aqueous metal halide solution.

Either of copper and iron can effectively act as an electron transfer agent. For example, in the recycled solution a good proportion of the (or each) metal can be in its relatively high oxidation state (e.g. Cu(II) or Fe(III)) and, after material oxidation, be in its relatively lower oxidation state (e.g. Cu(I) or Fe (II)), for re-oxidising to its high oxidation state. In the leaching stages the multi-valent species typically exists as a couple (i.e. in both high and low oxidation states and in a controlled balance/equilibrium).

Other multi-valent species that may be employed in the process include potentially cobalt, manganese and vanadium.

In one embodiment, the halide of the metal halide solution may comprise chloride or, more typically, may comprise a mixture of halides such as chloride and bromide. The concentration of halide(s) may be approximately 5-8 moles per litre.

Typically the precious metal to be recovered comprises gold, the recovery of which can justify the process economics.

When a high level of carbon is present in the sulfidic material (e.g. 2-20 wt.% carbon), a surfactant such as a blinding agent can advantageously be added to the solution during sulfidic material oxidation to prevent a precious metal adsorbing onto carbon in the material. The blinding agent may comprise one or more organic solvents including kerosene, phenol ethers, etc.

A most advantageous application of the process disclosed herein is in relation to the recovery of gold from pyritic ores and concentrates, where typically the contaminant is arsenic, antimony, bismuth, mercury, cadmium, etc and which occur naturally in many as-mined pyritic materials.

Other economically significant metals may additionally be recovered in the process including copper, nickel, zinc, lead etc. In addition, in certain applications, the contaminant may itself be desirable or necessary to recover. For example, the contaminant may be economically valuable or environmentally harmful, prompting its recovery from the contaminant precipitate (e.g. this may be the case for a contaminant such as antimony, bismuth, cadmium etc.).

Brief Description of the Drawings

Notwithstanding any other forms which may fall within the scope of the process as set forth in the Summary, specific forms of the process will now be described, by way of example only, and with reference to the accompanying drawings in which:

Figure 1 schematically depicts a generalised process flow diagram for a first mode for the recovery of a precious metal (gold) from a contaminated sulfidic material arsenopyritee FeAsS);

Figure 2 schematically depicts a generalised process flow diagram for a second mode for the recovery of a precious metal (gold) from a contaminated sulfidic material arsenopyritee FeAsS);

Figure 3 plots a % of gold extraction against different redox potentials employed for the process;

Figure 3 plots a % of gold extraction against different redox potentials employed for the process.

Detailed Description of Specific Embodiments

In the following detailed description, reference is made to accompanying drawings which form a part of the detailed description. The illustrative embodiments described in the detailed description, depicted in the drawings and defined in the claims, are not intended to be limiting. Other embodiments may be utilised and other changes may be made without departing from the spirit or scope of the subject matter presented. It will be readily understood that the aspects of the present disclosure, as generally described herein and illustrated in the drawings can be arranged, substituted, combined, separated and designed in a wide variety of different configurations, all of which are contemplated in this disclosure.

There are a number of factors that can render a gold-bearing ore refractory, as shown in the following table:

Type Causes of Refractory Characteristics

Liberation Physical locking in silicates, sulfides, carbon, etc.

Occlusion Passivation due to formation of a chemical layer.

Chemistry Formation of auriferous compounds e.g. gold tellurides and aurostibnite.

Substitution Elemental replacement by gold in mineral lattice e.g. "solid solution" gold in pyritic ores.

Adsorption Adsorption of dissolved gold by 'active' carbonaceous material in the ore pulp.

The process as disclosed herein was developed specifically to treat concentrates produced from those refractory ores falling into the latter two categories of "substitution" and "adsorption". The major proportion of the world's gold reserves fall into these two categories, which are dominated by iron sulfides such as arsenopyrite and pyrite, occurring either separately or, more commonly, in combination. The process was also applicable when "active" carbon was also present in the ore.

First, the process chemistry will be described for the treatment of refractory gold concentrates containing arsenopyrite and pyrite. Then, specific process flowsheets will be described, followed by experimental Examples of the process.

Leach Chemistry

The basis of the present process was an 'aeration leach' - a copper "catalysed" utilisation of air to oxidise arsenopyrite to scorodite and elemental sulfur, according to the following equations:

FeAsS + 7Cu2 + + 4H20 4 Fe 2 + + H3AsO 4 + S° +7Cu+ + 5H (1)

Cu2 + + Fe 2+ 4 Fe 3+ + Cu+ (2)

Fe 3+ + H3AsO 4 4 FeAsO 4 + 3H (3)

4Cu+ + 02 + 4HW 4 4Cu 2 + + 2H20 (4)

FeAsS + 202 4 FeAsO 4 + S° (5)

Gold was typically "locked" in this arsenopyrite, principally as a lattice-bound species, often referred to as a solid solution, rather than as native gold. Consequently, destruction of the arsenopyrite lattice by chemical oxidation promoted gold liberation, with the process parameters being controlled so as not to generally promote gold oxidation. Oxygen (e.g. in the form of air sparged into the leach at atmospheric pressure) acted through several intermediate steps, as its solubility in the process liquor was low. The oxygen 2 generated a soluble oxidant in the form of cupric ion (Cu +) according to reaction (4), with the reaction taking place at the interface between the air bubbles and the process liquor. Ferrous and cuprous reaction products were subsequently oxidised by further air sparging according to reactions (2) & (4). The action of the Cu 2 +/Cu+ couple was supplemented by the Fe 3 /Fe 2 +

couple, as a small background concentration of iron was always present in the process liquor. In the presence of ferric ion, the arsenic acid of reaction (1) readily formed insoluble ferric arsenate, according to reaction (3), in the complex halide electrolyte. The ferric arsenate was typically crystalline and stable in the environment, enabling its easy separation. The other major sulfide mineral present in the feedstock was "pyrite". It was noted that the "pyrite", as indicated by XRD analysis, may or may not actually be true pyrite (FeS 2 ), and might instead be any combination of a multiple of iron/sulfur analogues, herein collectively termed "pyrite" for XRD, etc purposes. It was further noted that true pyrite does not leach in the present process, even under quite aggressive conditions (very low pH and high redox potential). Conversely, iron sulfur analogues were more or less refractory in the present process, whereby such "pyrites" leached quite well to be no longer detectable in the XRD analysis of the residue. Thus, in the 'aeration leach', the following reactions can also be expected to occur:

FeS 2 + 15Cu2+ + 8H 2 04 Fe2+ + 15Cu& + 2SO42- + 16H (6)

Cu2+ + Fe>4Fe>+ + Cu+ (2)

2Fe 3* + 3H204 Fe 203 + 6H (7)

4Cu+ 02 + 4HW 4 4Cu2+ + 2H20 (4)

4FeS 2 + 1502 +8H 20 42Fe 2 0 3 + 8SO42- + 16H (8)

It was noted that the pyritic sulfur was oxidised all the way to sulfate in contrast to the arsenopyritic sulfur that was only oxidised to its elemental state. Because the pyritess" tended to be more refractory than arsenopyrite, a fine grind size was employed to achieve acceptable reaction kinetics. Reaction kinetics were also influenced by arsenic substitution for a portion of the sulfur in the crystal lattice such that, the higher the arsenic contamination, the more the pyrite reactivity approached that of true arsenopyrite, with an As/S ratio of one.

The reaction proceeded through the Cu 2 +/Cu+ couple as for arsenopyrite in the same liquor used for arsenopyrite oxidation according to reaction (6). Again, Cu+ and Fe2+were oxidised by oxygen sparging according to reactions (2) & (4). The ferric sulfate formed was precipitated as hematite and gypsum by the addition of limestone at a pH of approximately 1 1.5 according to reactions (10) & (11) below.

Because the "pyrite" minerals, unlike other sulfide minerals in the present process, reacted to form sulphate, this increased the oxygen demand, and also generated considerable acid. The sulphate was immediately precipitated from the system as CaSO 4 via the background levels of calcium in the electrolyte. Due to the temperature (~80°C) and the effect of the electrolyte, this CaSO 4 mostly formed as anhydrite, possibly with some bassanite (CaSO4.0.5H 20), and advantageously problematic gypsum (CaSO 4.2H 20) formation was negligible.

The removed calcium was then replaced via the addition of limestone in a second stage of the leach circuit, where acid neutralisation took place. Limestone addition was controlled to maintain soluble iron in the range 2-5g/l, which prevented the precipitation and loss to the leach residue of cupric copper. The net result on the electrolyte composition was neutral across the whole leach circuit, albeit that the reactions were physically separated to different areas of the leach:

Ca2+ + S0 4 2 - + CaSO 4 (9) Ca2+ CaCO3 + 2H+ 4 H 2 0 + CO 2 + (10)

4FeS 2 + 1502 + 8CaCO3 42Fe 2 0 3 + 8CaSO4 + 8C0 2 (11

Oxygen Utilisation

Assuming an acidic complex halide electrolyte that is nominally 100% cupric (Cu 2 +),

the two key reactions affecting the rate of reaction were:

(1) and (6), relating to the leaching of the sulfidic minerals present in the ore concentrate. In these reactions the cupric (Cu 2 +)ions were in turn reduced to cuprous (Cu+) and this caused the observed redox potential to drop (e.g. from >550mV vs Ag/AgC1 to <450mV); and

(4) the re-oxidation of the cuprous back to cupric, which caused the Eh to rise again.

Depending on the types and grades of minerals present and the particle size, the rate of reaction (1) was able to be run faster or slower. Particle size, in particular, was noted to be a dominant factor, and temperature also had a substantial influence.

Conversely, it was noted that the pH and oxygen utilisation rate were dominant factors affecting the rate of reaction (4), as well as the cupric ratio in the electrolyte. Both the acid and the copper levels in solution were noted to influence the rate, via Le Chetalier's principle, while the oxygen availability was most heavily influenced by the form of the oxygen added(02 vsair) and the physical parameters of the test.

In process batch leach tests, reactions (1) and (6) dominated in the early period due to the relative abundance of leachable minerals in the 'fresh' concentrate and the high "cupric ratio" (the ratio of cupric copper/total copper). As the cupric was consumed and the ratio dropped (and excluding the influence of iron in solution, which tended to increase the redox potential), then the redox potential also dropped.

In the experimental testwork, reaction (1) was interpreted to include either of the mineral reactions, arsenopyrite (1) or "pyrite" (6), as well as any other mineral sulfide reactions that might be occurring.

It was further noted that, as the relative abundance of cuprous ions increased following the leaching of the mineral via reaction (1), the rate of reaction (4) increased. Thus, during the main 'body' of leaching, either reaction (1) or (4) will be the rate limiting step. When reaction (1) was rate-limiting, then the cupric ratio tended to stay higher and the redox potential also tended to stay above e.g. -500mV. The aim of the process was to maintain the conditions in this form, and where possible (such as via finer grinding and/or hotter operation) to maximise the rate of this reaction. In doing so, the leach process efficiency was maximised.

However, it was noted that in many cases the rate of mineral leaching was extremely fast, whereby the cupric would be consumed quickly, and then reaction (4) would become rate limiting. In such cases, the redox potential would drop as the cupric ratio dropped. The reaction was particularly noted to be affected by the oxygen utilisation efficiency of the reactor configuration.

Again, it was noted that in larger-scale process reactors, the air bubbles have a long vertical travel distance and appreciable residence time in the turbulent reactor flow. Typical measured oxygen utilisation efficiencies in pilot and demonstration plant reactors are above 30% using compressed air and simple tube outlets (not dispersers). However, in the small 2L beaker tests, oxygen utilisation efficiency was measured to be as low as ~5%, largely due to the short residence time of bubbles, during the short vertical travel. In such cases, reaction (4) tended to become critically rate-limiting, hindering the efficient leaching of the mineral within acceptable residence times. While the problem can be partially overcome by increasing the air flow rate, the increased air flow then had a cooling effect on the reactor.

Most of the small beaker tests were conducted using oxygen rather than air simply to prevent a 'false negative' in testing. Thus, assuming an oxygen utilisation efficiency of 5% in a small beaker and 30% in a commercial process leach tank, it was seen that 5% x 100% 02= 5% 02, whereas 30% x 20.5% 02 (air) = 6.2% 02. In other words, the use of oxygen in a small tank was a reasonable analogue for air in large scale process reactors.

In any case, as the mineral was consumed, the rate of reaction (1) slowed, and as the rate of cupric regeneration (4) exceeded (1), the cupric ratio and therefore the redox potential rose. In the batch tests, a simple means ofjudging the completeness of the reaction was to simply stop the air/oxygen flow once the Eh plateaued above e.g. 550mV. If the Eh remained stable, this indicated that cupric was no longer being consumed by the mineral, and that the leachable mineral was likely fully consumed. At that point, a sample of electrolyte diluted 1/50 in ammonium thiocyanate would show no immediate precipitation, which was indicative that the cupric ratio was 1.

In the experiments, the air/oxygen flow rates were approximate and hence the oxygen utilisation efficiency for the particular system being tested was not precisely measured.

Direct Observation of the Process

AAS measurement of Cu and Fe in solution was employed. Other measurable data such as pH, Eh, temperature, flow rates and periodic liquid assays were measured. In addition to these tests, direct observations were useful forjudging the progress and success of the process leach.

By way of example, colour was used to judge the cupric ratio. In a system with no nickel or iron in solution, then a filtered electrolyte sample with a cupric ratio of 1 was noted to be a clear emerald green. As the ratio dropped, it became darker brown then black. At a cupric ratio close to zero, it became clear pale yellow/brown, and at zero was clear and colourless. However, iron in solution tended to redden the electrolyte (particularly above -30g/L), and so colour was only a partial guide to the chemistry of the system.

2 Ammonium thiocyanate testing followed by Cu analysis gave a quantifiable Cu

+ analysis which was compared to the Cuto togive the cupric ratio, but towards a ratio of 1 it was also able to be used for a simple visual analysis. In this regard, the less visible CuSCN precipitate formation there was, the closer the ratio was to 1. This test, in combination with a measured Eh trend, allowed an operator to make a judgement about the progress of a test run.

In a similar manner to colour, the solution Eh was a very useful measurement, although noted to be influenced by multiple factors. In a system without significant iron in solution and e.g. around 50g/L copper in solution, at a cupric ratio of around 1, the Eh would be -550-58mV (vs Ag/AgCl). At a cupric ratio closer to zero, the Eh was likely to be <100mV. High levels of iron in solution (say >30g/L) tended to increase these figures by 20-4mV.

Halide complexes (such as BrCl2) were noted to form above about 650mV, although could not form in a system using air or oxygen. The formation of such complexes requires a more powerful oxidant source - electricity or a chemical oxidant such as hypochlorite. The choice to use only air in the tests was deliberate. By staying below the halide complex formation zone, the dissolution of gold was minimised (although it was noted that there may be some background levels of dissolved gold to be either recovered, precipitated or recycled with the electrolyte), leaving the gold to therefore be recovered in the cyanide circuit.

A key factor in the batch testing regime was judging the relative rates of reactions (1) and (4). The pH, Eh and other direct observational trends were utilised during the tests, and then the observations were compared against measured data (usually) after a test was complete.

pH Measurement

Apart from Eh, pH was the second important quantitative data measured regularly in the process. However, this measurement was technically challenging.

Firstly, it was noted that pH probes do not directly measure H+ concentration, but rather H+ activity. In theoretical systems at low concentration, concentration and activity were close enough to be roughly equivalent. However, at the very high electrolyte concentrations of the present tests, this was not true - the H+ activity was disproportionately high. As such, the pH in the testing was an empirical measurement yielded by a pH probe, but not the actual concentration of H+ ions.

Secondly, because the tested electrolytes were chemically aggressive, ordinary pH probes tended not to perform well over the medium- to long-term. As such, specialist probes were preferred.

Concentrate Grind Size

Concentrates for use in the process were typically received in the size range of 80% 15-30 microns. Tests indicated that reaction kinetics were significantly enhanced when the concentrates were ground to such a fine size (dependent on the characteristics of each individual concentrate), and in the process regrinding was typically employed. Where arsenopyrite was the sole gold-bearing mineral, a size of 80% less than 25 microns proved more than adequate to achieve good gold liberation and an acceptable leach retention time. Where gold was locked in pyrite, the grind size principally depended on the reactivity of the pyrite which, as previously explained, varied greatly. For a highly active pyrite, the grind employed for arsenopyrite was used, but more refractory pyrite examples required finer grinding. This sometimes extended to an ultra-fine grind with 80% less than 15 microns in a more refractory case.

Impurity Management

In addition to any major contaminants (such as arsenic, antimony etc), the presence of impurities in the feed concentrate (such as Cd, Mn, Mg, etc.) had no detrimental effect on either the leaching or precipitation operations. Nevertheless, a method for the management of impurities was employed to prevent their build-up in the process solution over time. This was achieved via precipitation from a bleed of the regenerated cupric solution with the purified electrolyte returned to the leaching process.

Limestone was added to the bleed to adjust the pH to 3.5, precipitating residual iron and copper, which were removed by filtration and recycled to the leach. Impurities, such as Cd,

Mn and Mg, were removed via slaked lime addition at pH 9 to form insoluble oxides that were recovered by filtration for disposal of an alkaline waste stream.

Determination of Total Gold Extraction

The aim of the process leach was to break down the arsenopyrite and pyrite minerals that should be hosting the majority of the gold. Some gold was noted to also be hosted in the quartz/mica phase(s), but these gangue minerals were not leached in the process (other than the possible liberation of exposed gold at the quartz/mica particle surfaces). Finer grinding was observed to increase this to some extent via the greater surface area.

It was noted that, in a full scale plant, a balance would be established whereby any small amount of leached gold would re-precipitate to the residue, leaving a small recycling load of gold in the process electrolyte. However, in batch tests this small amount of gold stayed in solution, and therefore the calculation of total gold extracted from the concentrate included both the gold extracted in the process and gold extracted by cyanide.

The methodology for calculating total gold extraction, using theoretical data, was:

Gold in concentrate = 60.7 g/t x 1,500g = 91,050 ig

Gold in residue = 38.0 x 1,355.5g = 51,509 ig

Gold extraction = 1 - (51,509 / 91,050)= 44.4%

The residue samples were then subject to a cyanide leach to determine the free gold content (5.0% CN, 0.7% NaOH, 0.7% leach acid, 25% w/w solids).

CN gold extraction = 32.1 %

Total gold extraction = 1 - ((1-44.4%) x (1-32.1 %)) = 62%

Treating Double Refractory Ores (arsenopyrite/pyrite/carbon)

The impact of carbon in the processing of gold concentrates was largely a function of its grade and activity. At the lower range of carbon content, organic additives (blinding agents) could be used to inhibit gold adsorption. However, when the content of carbon started to exceed 3 to 5%, the effectiveness of inhibition was greatly reduced as so-called "preg-robbing" of gold increased. In this instance the destruction of carbon by roasting was the main treatment option practiced in the art. This was noted to be a relatively complex process, as gold extraction from the resulting calcine was affected by the roasting conditions. Further, the optimal conditions for pyrite roasting differed from those of arsenopyrite, necessitating a two stage roasting process.

The present process could be implemented prior to the roasting of selectively leached arsenic and sulfur, to simplify subsequent roasting, which in this instance could become a simpler single-stage process. Further, the removal of arsenic and sulfur was noted to reduce the duty for off-gas scrubbing from roaster operations, because As 2 03 and SO 2 were greatly reduced. The impact would be one of significantly reduced capital and operating costs in the roasting step.

Cyanide Consumption

The consumption of cyanide in extraction of the gold from the process residue was noted to be an important issue. In the process, the oxidation of the mineral to elemental sulfur (or, in the case of pyrite, to sulphate) was catalysed by the presence of cupric in solution. This copper was not allowed to pass through to the cyanide leach or it would cause unacceptable CN consumption. Hence, the wash/filtration stage discussed below. This is also discussed further in the Examples.

It was also noted that, because there were no other key oxidising species in the process leach, after good washing, the only other component of the residue of potential concern might be elemental sulfur (i.e. in some cases elemental sulfur can react with the CN- to form SCN-).

However, it was possible to selectively remove the elemental sulfur from the process residue by either flotation or organic dissolution (i.e. the sulfur dissolved quite readily in both xylenes and tetrachloroethylene). Thus, simply by raising and lowering the temperature, the dissolution and re-precipitation of elemental sulfur was able to be controlled as necessary.

Slurry Suspension/Mixing

The 1.2L tests were conducted in a glass beaker using a baffled lid to aid turbulent mixing, whereby it was easily observable that the solids were fully suspended and that there were no zones of solid accumulation in which mixing would be inadequate. The 6.OL tests were conducted in a titanium tank. The tank included a Rushton-type agitator with a diameter ratio vs the tank of 0.3, and with baffles built into the sides of the tank. Some of the data from tests (described below) suggested that the slurry might not have been fully suspended in the larger tests using 25% w/v solids, which can hinder reaction rates. Tests that used 20% w/v lessened the likelihood of inadequate mixing.

Process Flowsheets

Referring firstly to Figure 1 a single refractory pyritic gold liberation process 10 is schematically depicted. A precious metal concentrate 12 for feeding to the process is prepared by mining, milling and then flotation of a sulfidic ore, or via gravity separation. The concentrate is typically a gold-containing arsenopyrite (and when it has high carbon content it becomes double refractory). The concentrate is ground in a special ball mill, typically to an ultra-fine level of around Pso less than 25rm.

The ground concentrate is fed to a first repulp stage 14 in which it is combined with recycled electrolyte streams 48 and 60 (described below). From the first repulp stage 14 the combined electrolyte is passed to a surge tank 16 to enable suitable flow control through the process. From surge tank 16 the process electrolyte is passed to a first leaching stage 18.

In leaching stage 18 an acidic environment is maintained (preferably a pH of less than ~1 but greater than about 0.5) so as to precipitate the arsenic immediately after it is leached. The acid environment can be achieved solely by the solution recycle, but typically a non contaminating acid such as sulfuric acid 20 is added (or hydrochloric acid, etc), together with an oxidant such as air 22 (although oxygen, chlorine, hydrogen peroxide, etc may be considered). The air stream 22 is optimally sparged into the leaching vessel(s).

The process electrolyte is typically an aqueous cupric, mixed chloride/bromide solution, having a mixed halide concentration of around 5-8 moles/litre. In leaching stage 18, as part of the oxidising of the arsenopyrite and "pyrite" materials, the cupric ion is reduced to cuprous ion (equations (1) and (2)), and the ferrous iron produced in the oxidising of the arsenopyrite is oxidised to ferric ion (equation (2)). The sparged air also acts to regenerate the cupric ion (equation (4)). Thus, in the process, copper and iron each act as electron transfer agents, existing in equilibrium as the Cu 2+/Cu+ and Fe7/Fe 2+couples. Other agents can perform this function, including cobalt, manganese, vanadium, etc.

In the first leaching stage 18 the leaching conditions are controlled to achieve an oxidation potential (Eh) of greater than -0.5 volts and less than about -0.65 volts (ref. Ag/AgCl), to promote oxidation of the arsenopyrite and "pyrite" components of the material (combined reactions (1) to (8) above), but without any substantial oxidation/leaching/solubilisation of gold. Rather, the destruction of the arsenopyrite and pyrite lattices by chemical oxidation is controlled so as to only promote liberation (i.e. "freeing-up") of gold-containing species, which species can then be subsequently leached in a conventional (cyanide) leach. The solution leaching temperature is maintained at around 80-95°C.

Water that evaporates during the first leaching stage 18 may be captured as stream 24 and reused in the process (i.e. as fresh water stream 80, 90).

The leached material is then fed to a second leaching stage 26 where an oxidant such as air 28 (although oxygen, chlorine, hydrogen peroxide, etc may be considered), together with limestone 30 (CaCO3 slurry), are added to control the solution oxidation potential and to raise the solution pH to above -1 and less than -2. The conditions in the second leaching stage 26 are optimised towards the maintenance of stable ferric arsenate (FeAsO 4) and ferric oxide (Fe 2 0 3

) precipitates (i.e. to prevent the solubilising of arsenic and iron).

As a variation to the process flowsheet shown in Figure 1, the iron can be filtered while it is still in solution. This would leave the residue free of iron, which can provide certain advantages in the subsequent cyanide leaching of the residue. In this regard, thickener and filtration stages can be provided after the first leaching stage 18, and another filtration stage (or the previous filtration stage performing double-duty) can be provided after the iron removal from solution in the second leaching stage 26.

The oxidation potential Eh in the second leaching stage 26 is again maintained at less than about -0.65 volts (ref Ag/AgCl), to ensure that gold is not oxidised/leached/solubilised. In addition, the temperature of the solution in the second leaching stage is raised to around 80 900C.

Because excess acid is produced during pyrite oxidation (equations (6) and (8)), not all of which can be consumed when Cu(II) is regenerated (equation (4)), and because calcium in solution precipitates as CaSO4 (equation (9)), the limestone stream 30 (CaCO 3 as a slurry) is added.

Again, water that evaporates during the second leaching stage 26 may be captured and reused in the process as stream 32 (i.e. as fresh water stream 80, 90). In addition, carbon dioxide that is produced by limestone addition 30 (reactions (10) and (11)) may be captured or released to atmosphere 34.

The resultant solids slurry 35 from the second leaching stage 26 is passed to a solid liquid separation stage that comprises thickening stages, an impurity treatment stage and filtration and wash stages.

A first thickening stage comprises a thickener 36 in which flocculant 38 is added to promote settling out of the solids residue that contains the gold-bearing species. That residue is passed as an underflow 40 to an underflow holding tank 42, awaiting the wash/filtration stage.

The overflow stream 44 from the thickener 36 is passed to an overflow tank 46. A major portion of the overflow stream fed to the overflow tank 46 is recycled as recycle stream

48 to the first repulp stage 14, to effectively reuse the acidic complex halide electrolyte. A bleed stream 50 is passed to a first bleed treatment tank 52 for solids removal.

Limestone 54 (i.e. a CaCO3 slurry) is added to the bleed treatment tank 52 to cause the formation of a hematite/gypsum precipitate (reaction (11)). More specifically, limestone is added to the bleed tank 52 to adjust the pH to -3.5, thereby causing residual iron and copper to precipitate. The resultant slurry (i.e. comprising the iron/copper precipitate) is passed as a stream 56 to a second thickening stage that comprises a thickener 58.

The underflow solids from thickener 58 are recycled as slurry 60 to the first repulp stage 14, where it is blended with the recycle solution 48 (i.e. for recycling to the first leaching stage 18). The overflow solution 62 from thickener 58 is passed to a second bleed treatment tank 64 for impurity treatment. In tank 64 impurities such as Cd, Mn and Mg, although having no detrimental effect on the process operations, are removed via the addition of lime 66 (i.e. slaked lime slurry) up to a pH of 9 to form insoluble oxide precipitates to be recovered by filtration. More specifically, the resultant slurry 68 is passed to a first filtration stage 70 which filters the insoluble oxide precipitates for disposal as an alkaline waste stream 72.

The brine filtrate 74 from the first filtration stage 70 is passed to a second filtration/wash stage 76 (e.g. comprising a belt filter), where it is combined with the underflow stream 78 from the underflow holding tank 42. At a first (or early stage) of the belt filter the solids can be separated from a filtrate steam 79, which can in turn be recycled to the overflow tank 46 where it is combined with the overflow stream 44, a significant proportion of which is recycled as solution 48 to the first repulp stage 14.

At a second (or later stage) of the belt filter, fresh make up water 80 (which can include evaporated water streams 24 and 34), can be added to wash the filtered solids and thereby produce a "clean" solids residue/filter cake 82 and a separate "dirty" wash water filtrate 84. The wash stage can remove metal salts that can interfere with subsequent gold (e.g. cyanide) leaching in a cyanide circuit 96.

The resultant washed residue/filer cake 82 is passed to a second repulp stage 86, where it is repulped into a slurry 88 through the addition of fresh water 90, dirty process wash water 84 from e.g. the belt filter of the second filtration/wash stage 76, and lime 92 (i.e. slaked lime slurry) to control (raise) the solution pH. The resultant slurry 94 is passed to the cyanide circuit 96. The fresh water 90 can include evaporated water streams 24 and 34.

In a variation on the process 10 of Figure 1, when the residue that comprises the precious metal contains elemental sulphur, and prior to it being repulped and passed to the cyanide circuit 96, the residue may first be subjected to a sulphur removal stage. This is because sulphur can consume cyanide.

Typically the sulphur removal stage is applied to the washed residue/filter cake 82 from the filtration/wash stage 76 (i.e. once that cake has been separated from the solution) and prior to it being passed to the second repulp stage 86.

The sulphur removal stage can comprise one or more of direct oxidation, thermal solvent extraction and flotation. Direct oxidation (e.g. using hydrogen peroxide) can be employed to oxidise the elemental sulphur to sulphate, to then enable its easy removal (e.g. in a separated solution). Thermal extraction of the elemental sulphur can occur via direct dissolution in organic solvent(s), such as xylene, tetrachloroethylene and/or kerosene. Flotation can be employed to separate sulphur and residual mineral sulphide species, and can comprise e.g. one to two flotation stages, and use e.g. kerosene and pine oil as flotation frothers.

Alternatively, sulphur removal may take place within the leaching stages 1 & 2. For in-leach oxidation of sulphur, the flow sheet (i.e. as shown in each of Figures 1 & 2) can be changed from a co-current leaching arrangement to a counter-current leaching arrangement. The leaching process can then be adjusted such that a leaching solution is prepared that has an oxidation potential that is sufficiently high to oxidise elemental sulphur and thereby solubilise it.

Referring now to Figure 2, where like reference numerals denote like parts, a variation on the process 10 is schematically depicted. In this variation, the process 100 of Figure 2 employs a first solid/liquid separation stage 102 between the first leach stage 18 and the second leach stage 26, and also employs a second solid/liquid separation stage 104 after the second leach stage 26.

The first leach stage 18 can be operated such that the metal salts that can interfere with subsequent gold (e.g. cyanide) leaching in the cyanide circuit 96 can be dissolved in the solution. Thus, these salts pass with the liquid 106 through the first solid/liquid separation stage 102. The separated solids residue stream 108 is thus effectively already "washed".

The separation of the first leach stage 18 and second leach stage 26 also allows for separate iron recovery. This iron residue 110 can be separated out from the stream 112 leaving the second leach stage 26 in the second solid/liquid separation stage 104. This iron residue 110 can represent a by-product, and can provide another potential way to prevent interference with subsequent cyanide processing 96.

The separated liquid solution 114 from the second solid/liquid separation stage 104 is recycled to a bleed treatment tank 116, where impurities such as Cd, Mn and Mg can be removed via the addition of slaked lime (up to a pH of 9) to form insoluble oxide precipitates to be removed from the process. The resultant "treated" solution 118 can be recycled to the repulp stage 14. The process 100 of Figure 2 is thus inherently simpler than the process 10 of Figure 1, although the process 100 of Figure 2 may not be suitable for all feedstocks.

In either of processes 10, 100, where the sulfidic material has a high carbon content (e.g. up to 3-5 wt.%) a masking surfactant may be added to the solution at leaching stages 18 and/or 26 to prevent adsorption onto carbon of any species (including gold or other precious metals) leached into the solution. The surfactant is typically an organic blinding agent such as kerosene, a phenol ether etc.

Materials of construction of process apparatus include fibre-reinforced plastic, rubber lined steel, fluoropolymers, polypropylene and titanium.

Examples

Now that optimal process flowsheets have been described, specific examples that further demonstrate the process will be described.

Experimental testwork was performed on a series of concentrate (sulphidic ore) samples containing arsenopyrite, which were noted to not respond well to conventional cyanide treatment.

Table 1: Concentrate Grades

Element Unit Range Tested

Au ppm 13-60

Ag ppm <0.5-113

Al % 3.8-8.1

As % 1.0-3.0

Ca % 0.0-3.6

Cd ppm <1

Co ppm 24-30

Cu ppm 40-550

Fe % 5.4-7.5

K % 1.2-3.3

Mg % 0.2-1.7

Mn ppm 160-840

Na % 0.0-0.7

Ni ppm 180-530

Pb ppm 37-63

S % 3.2-7.5

Sb % 0.01-4.7

Zn ppm 42-130

Procedure

The testwork involved bench scale batch tests at a 2L and 6L scale. Initial tests were conducted in a baffled 2L beaker with an overhead stirrer and a pipe to deliver air to the mixing zone at the bottom of the vessel.

Liquors were prepared containing 150g/L CaCl2 , 100g/L NaBr, 50g/L Cu(II) and up to 10g/L Fe(III), and then heated to 60-100°C. Fresh concentrate was added at 20% w/v (200g of concentrate per 1.L of liquor). The slurry was agitated to ensure no settling of solids, and where possible (in 2L glass beaker) this was checked visually. Air or oxygen were added to the maximum rate without excessive foaming, typically 0.5-0.8L/min per litre of slurry.

Leach slurries were agitated for 3-9 hours, with measurement of Eh (ORP redox potential vs Ag/AgCl) and pH and typically hourly sampling of up to 1OmL for filtration and liquor analysis. pH was typically maintained -1.0 using HCl.

Control of the operating process used measurable data such as pH, Eh, temperature, flow rates and periodic liquid assays. In addition to this, there were a range of direct observations that were extremely useful forjudging the progress and success of the process leach. As example, colour was, to some extent, able to be used to judge the ratio of Cu(II) to Cu(total) in the process liquor.

Ultimately, a key factor in batch testing was judging the relative rates of reaction of (1) and (4). The pH, Eh and other direct observational trends were utilised during the test itself, and then the observations were compared against the measured data when that became available (usually) after the test was complete.

At the end of the tests, the slurry was filtered using a Buchner funnel. The solids were washed with brine (-100g/L NaCl at pH ~1) to remove metal salts, and then with hot water to remove residual sodium/calcium salts. The resulting solids were dried overnight and weighed.

Laboratory testwork on the concentrate sample indicated that:

• The process as described herein was able to be operated as an additional unit operation to be applied to existing concentration and cyanidation site infrastructure.

• The gold was liberated via the present process by breaking down the host sulphide mineral matrix, but was not substantially leached under preferred (controlled) conditions. The liberated gold was then able to be extracted and recovered by conventional cyanidation of the process residue.

• Up to 99% of the available gold (in the sulfidic mineral phases) was liberated. The total gold recovery after cyanidation was noted to be dependent on the quantity that was locked in non-leachable quartz/mica phases, which would be similarly unavailable to all hydrometallurgical leaching technologies.

Particle Size It was noted that particle size had a strong effect on the reaction rate within the process, as expected with hydrometallurgical technology. This was both due to the surface area for reaction and the 'shrinking core' model of reaction.

Under most circumstances, ultra-fine grinding (to less than 10im) was not recommended to be necessary. The 'typical' recommendation for grinding for the process was 25pm, with larger sizes potentially being acceptable depending on the mineralogy of the concentrate.

Determination of Total Available Gold

An initial test was carried out to determine the total available gold by oxidising the sample at a very high potential using calcium hypochlorite to maintain an Eh>650mV (vs Ag/AgCl) over a period of 5 hours. Under these conditions, the total extractable gold was 68%, with the residual 32% locked in the refractory minerals, quartz and mica.

Oxidant

It was noted that the oxygen utilisation efficiency was strongly affected by the scale of the tests, as well as by the background concentration of copper in solution. Most particularly, the efficiency of the take-up of air at the 2L beaker scale was poor, partially due to the very short vertical distance for rising air bubbles to travel and the associated short residence time.

In many cases, 2L beaker tests were conducted using oxygen rather than air. Assuming an oxygen utilisation efficiency of 5% in a small beaker and 30% in a leach tank, then it can be seen that 5% x 100% 02 = 5% 02, whereas 30% x 20.5% 02 (air) = 6.2% 02. In other words, the use of oxygen in a 2L beaker was a reasonable analogue for air in larger tanks. Larger-scale tests were almost exclusively conducted using air, with the process flowsheet being based on the use of air.

In general terms, more aggressive conditions in the process (higher temperature, lower pH, higher oxidation potential) was expected to yield higher gold liberation and/or faster reaction times, but involve higher costs in terms of reagent additions, energy consumption and materials of construction. The overall aim of the laboratory programme was to determine the leaching conditions that would yield the best combination of highest-practicable gold liberation (not leaching) and lowest practicable operating cost.

Test Results and Discussion

As an example sequence of testing, an as-received concentrate sample containing 14g/t of gold, approximately 2.5% arsenopyrite and 3.5% pyrite was leached in 2L beaker tests over 5 hours at 80°C. Redox potential was varied through the use of different oxidants, air, oxygen and/or calcium hypochlorite.

Example results are shown in Table 2.

Table 2: Effect of Changing Redox Potential Experiment Target Process CN Total

Eh Au Au Au

mV % % %

002 500-550 22.5% 59.2% 68.4%

003 650 31.5% 42.4% 60.5%

004 650 26.7% 48.2% 62.0%

005 750 41.2% 52.2% 71.9%

006 750 50.6% 45.0% 72.8%

007 750 36.0% 44.0% 64.1%

Various observations were made as follows:

Effect of Redox Potential

Within the process, it was generally expected that gold would dissolve above-650mV (vs Ag/AgCl), with gold dissolution facilitated by the presence of ferric ions.

At the same time, higher redox potentials would be expected to facilitate faster reaction times via complete oxidation of copper ions to the cupric state, as well as via direct oxidation of the target sulphide minerals. The same mechanism may also yield higher overall breakdown of leachable minerals, if such breakdown is incomplete at lower redox potentials. While pyrite (FeS 2 ) should not generally be leached in the process, some pyritic iron sulphide analogues may leach, and these refractory analogues are more likely to leach at higher redox potential (>650mV).

As shown in Figure 3, and in Table 1 and Table 3, increasing the redox potential did not increase the total gold liberation after cyanidation. While greater amounts of gold were extracted within the process itself at higher redox potentials (as would be expected), this simply reduced the amount of gold available for subsequent cyanide extraction.

Table 3: Effect of Changing Redox Potential Test Target Process CN Total

Eh Au Au Au

mV % % %

08 500-550 22.5% 59.2% 68.4%

09 650 31.5% 42.4% 60.5%

10 650 26.7% 48.2% 62.0%

11 750 41.2% 52.2% 71.9%

12 750 50.6% 45.0% 72.8%

13 750 36.0% 44.0% 64.1%

14 500-550 44.8% 57.5% 76.5%

Given that the maximum gold liberation of 68-73% across the test series correlated strongly to the total available gold from independent mineralogical studies, these data strongly supported the hypothesis that the process has liberated all of the available gold, and that the remainder of the gold in the sample was trapped in the non-sulphide host minerals, primarily quartz and mica.

The data also suggested that it was not necessary to use the higher redox potentials to maximise attack on the non-arsenical iron sulphide minerals (pyrite is often a mixture of iron sulphides) to liberate the gold contained therein. Less aggressive (and lower cost) conditions appeared to be adequate.

Effect of Residence Time

Over a series of nearly 50 tests across five concentrate samples, it was concluded that the full leaching of the minerals was a matter of sufficient residence time, with various factors influencing this time. In approximate order, these were:

1. total mineral sulphide content in the concentrate;

2. particle size distribution of the concentrate;

3. oxidant type / oxygen utilisation efficiency / background copper tenor;

4. temperature;

5. electrolyte composition; and

6. other sundry factors.

Effect of Electrolyte Composition

The electrolyte composition was important to process economics. Sodium bromide, particularly, was comparatively expensive, and the inevitable loss of soluble salts with the leach residue (due to washing inefficiencies) affected the operating cost. It was also noted that these soluble losses due to washing were the only appreciable salt loss, and that no disposal of electrolyte was envisaged to maintain salt balance.

At the same time, the electrolyte composition had a marked effect on the overall performance of the process. Bromide, particularly, was a stronger chelating agent than chloride, and thus higher bromide levels improved the ability to liberate the gold and to maintain that gold in solution.

A test sequence was conducted at varying electrolyte compositions in 6L of process liquor, using 1.2kg of sample and at 80°C. It was noted that the concentrate sample selected for this test sequence was high in mineral sulphide content (up to 20%). By using a limited leach time on the sample, this allows the influence of other factors contributing to leach kinetics to be more easily identified. The results are shown in Table 4:

Table 4: Electrolyje Testing Results CaCl 2 NaCl NaBr Cu Total CN (g/L) (g/L) (g/L) (g/L) Halide Extraction (M) (%)

039 100 100 30 50 5.4 21%

041 100 100 30 75 6.2 15%

042 100 100 30 75 6.2 45%

046 150 0 100 75 6.0 77%

Effect of Temperature

Temperature was clearly a critical factor to process economics. Higher temperatures were expected to yield faster reaction times, and in some cases also gave higher total extraction. However, higher temperatures increased the energy costs, evaporative losses and materials of construction requirements. Most particularly, above 80°C various inexpensive polymer materials could become compromised and cease to be useful in the process.

As can be seen in these data, in particular Figure 4 and Table 5, higher temperatures appeared to have a small positive effect on total gold liberation from tests on an 'equal' footing (Eh ~550mV, 5 hours, 20% w/v, 'standard' electrolyte composition), but the data were not completely linear. This suggested that other factors were more important than temperature in influencing total gold liberation under the operating conditions.

Table 5: Effect of Temperature Test Temp Process CN Total

Au Au Au

0C % % %

002 80 22.5% 59.2% 68.4%

008 100 22.8% 61.5% 70.3%

009 60 11.0% 57.9% 62.6%

012 80 44.8% 57.5% 76.5%

Effect of Sample Variation

Overall, five samples were tested from the same project. Peak gold extraction results for four samples tested under identical conditions are shown in Table 6. All tests were conducted using the same electrolyte composition (150g/L CaCl 2 , 100g/L NaBr), at 80 0 C in 6L vessels using air. Insufficient sample was available to test sample 'Gravity C388'.

Table 6: Series 5 Sample Variations Cvanidation Results Test Sample Oxidant Time Total CN Extraction (hr) (%)

012 Float Conc 1 Air 6 76.5%

046 Gravity C387a Air 9 77%

047 Gravity C387b Air 9 84%

045 Gravity C390 Air 9 70%

Additional Examples

Additional experimental examples to further demonstrate the process were performed. The experimental testwork was performed on a series of concentrate (sulphidic ore) samples containing arsenopyrite, which were noted to not respond well to conventional cyanide treatment.

Procedure

The testwork involved bench scale batch tests in a baffled 6L beaker with an overhead stirrer and a pipe to deliver oxygen to the mixing zone at the bottom of the vessel.

Liquors were prepared containing 1OOg/L CaCl 2 , 1OOg/L NaCl, 30g/L NaBr, and 100 150g/L Cu(II), and then heated to 80 0C. Fresh ore concentrate was added at 20% w/v (200g of concentrate per 1.L of liquor). The slurry was agitated to ensure no settling of solids. Oxygen was added to facilitate leaching at laboratory scale - noting that oxygen is a laboratory test analogue for the use of air at larger scale.

Leach slurries were agitated for 20 hours (as an excess to requirement, to remove the influence of leach time from the selected tests), with measurement of Eh (ORP redox potential vs Ag/AgCl) and pH, and typically hourly sampling of up to 10mL for filtration and liquor analysis. pH was typically maintained -1.0 using HCl or CaCO 3, as required.

Control of an operating process used measurable data such as pH, Eh, temperature, flow rates and periodic liquid assays. In addition to this, there are a range of direct observations that can be extremely useful for judging the progress and success of a process leach stage. As example, colour can to some extent be used to judge the ratio of Cu(II) to Cu(total) in the process liquor.

Ultimately, a key factor in the batch testing was judging the relative rates of reaction of mineral leaching and catalytic copper re-oxidation. The pH, Eh and other direct observational trends were utilised during the test itself, and then the observations were compared against the measured data when that became available, usually after the test was complete.

At the end of the tests, the slurry was filtered using a Buchner funnel. The solids were washed with brine (-100g/L NaCl at pH ~1) to remove metal salts, and then with hot water to remove residual sodium/calcium salts. The resulting solids were dried overnight and weighed.

Test Results and Discussion

Calculated Au Head Au Extraction from Test # Grade (g/t con) Filter Cake by CN(%)

060 10.7 97%

061 10.6 84%

062 8.1 80%

063 12.1 86%

064 10.7 82%

065 11.0 86.5%

066 12.5 86.5%

067 11.2 88%

From these results it can be seen that, ultimately, greater than 80% gold extraction was possible from the sulphidic ore concentrate containing arsenopyrite. In a number of cases, -90%

and up to 97% gold extraction was possible. Thus, with further process optimisation, consistently high extraction rates for precious metals can be achieved.

Now that specific process embodiments have been exemplified, it will be appreciated by those skilled in the art that the process provides the following advantages:

• The process can be employed to liberate a precious metal such as gold from sulfidic ores and concentrates which are otherwise difficult or impossible to treat using conventional available processes/techniques such as smelting, roasting and cyanide leaching.

• The process can accommodate carbon content in such ores, because the gold is not leached into solution until a subsequent e.g. cyanide treatment stage. In instances where the gold is leached into solution, blinding agents can be employed to prevent precious metal adsorption onto carbon, which can otherwise interfere with precious metal recovery.

• The process enables removal of arsenic, iron and sulfur in readily disposable forms from an original arsenopyrite concentrate, leaving a readily treatable concentrate.

• The process can be employed to remove contaminants from a wide variety of ore and concentrate feedstocks which, once removed, can then be treated using conventional smelting/roasting techniques.

• The process can be used to treat contaminated residues to allow them to be subsequently disposed of with reduced environmental impact.

The process as disclosed herein thus enables treatment of refractory pyritic and arsenopyritic ore bodies. The process is able to break down the matrix of the pyritic or arsenopyritic ore, and separate the impurities/contaminants from the residue that comprises the precious metal to be recovered. This then enables a clean, purified residue (e.g. filter cake) to be passed to a conventional precious metal recovery stage, such as cyanide leaching.

Whilst a number of specific process embodiments have been described, it should be appreciated that the process may be embodied in other forms.

In the claims which follow, and in the preceding description, except where the context requires otherwise due to express language or necessary implication, the word "comprise" and variations such as "comprises" or "comprising" are used in an inclusive sense, i.e. to specify the presence of the stated features but not to preclude the presence or addition of further features in various embodiments of the process as disclosed herein.

Claims (22)

1. A process for enabling a precious metal to be recovered from a sulfidic material that contains the same, the process comprising:

- preparing an acidic aqueous metal halide solution having an oxidation potential that is sufficient to oxidise the sulfidic material, with the metal of the metal halide able to function as a multi-valent species during oxidation of the sulfidic material;

- introducing the sulfidic material into the acidic aqueous metal halide solution and controlling the oxidation potential such that the sulfidic material is oxidised, the metal of the metal halide is reduced to a lower valent state, and the precious metal is generally not oxidised;

- oxidising the metal of the metal halide to a higher valent state;

-separating from the solution a residue that comprises the precious metal for subsequent recovery thereof.

2. A process as claimed in claim 1 wherein the oxidation potential is controlled to be less than an Eh of around 0.65 volts (ref. Ag/AgCl).

3. A process as claimed in claim 2 wherein the oxidation potential is controlled to be greater than an Eh of around 0.5 volts (ref. Ag/AgCl).

4. A process as claimed in any one of the preceding claims wherein the solution separated from the residue is recycled to enable preparation of the acidic aqueous metal halide solution.

5. A process as claimed in claim 4 wherein the recycled solution comprises metal of the metal halide solution oxidised to its higher valent state.

6. A process as claimed in any one of the preceding claims wherein oxygen or air is introduced into the acidic aqueous metal halide solution to oxidise the metal of the metal halide to its higher valent state.

7. A process as claimed in any one of the preceding claims wherein the process comprises a first leaching stage, and wherein the sulfidic material comprises a pyritic material that is oxidised in the first leaching stage.

8. A process as claimed in claim 7, the process further comprising a second leaching stage in which the solution conditions are controlled to maintain iron in its ferric state so as to facilitate its removal, including as an Fe 2 03 precipitate.

9. A process as claimed in claim 8 wherein, when the pyritic material comprises arsenopyrite, the oxidation potential in the first leaching stage is controlled to leach arsenic into solution, and the solution pH is controlled such that, once leached, the arsenic precipitates as ferric arsenate, and in the second leaching stage the solution Eh and pH are controlled to maintain arsenic as a ferric arsenate precipitate.

10. A process as claimed in claim 9 wherein the pH in the first leaching stage is less than 1 but greater than about 0.5 so as to precipitate the arsenic immediately after it is leached, and wherein the solution temperature is in the range of about 80-105 0C; and wherein the pH in the second leaching stage is less than -2.2 but greater than about 1 so as to maintain the ferric arsenate precipitate, and wherein the solution temperature is in the range of about 900 C to 1050 C.

11. A process as claimed in any one of claims 7 to 10 wherein, to facilitate oxidation of the pyritic material, an acid such as sulfuric acid, together with oxygen, are added to the first leaching stage.

12. A process as claimed in claim 11 wherein a base such as calcium carbonate, together with oxygen, are added to the second leaching stage.

13. A process as claimed in any one of the preceding claims wherein the residue comprising the precious metal is separated from the solution in a solid-liquid separation stage that comprises a thickening stage and a wash/filtration stage.

14. A process as claimed in claim 13 wherein:

a portion of overflow from the thickening stage is recycled to enable preparation of the acidic aqueous metal halide solution, and a portion is passed to an impurity removal stage;

in the impurity removal stage impurities are precipitated, filtered and disposed of, with a portion of filtrate being recycled to enable preparation of the acidic aqueous metal halide solution and with another portion of the filtrate being passed to the wash/filtration stage;

underflow solids from the thickening stage are passed to the wash/filtration stage to be washed and filtered, with filtrate being recycled to enable preparation of the acidic aqueous metal halide solution, and with washed solids being passed to precious metal recovery.

15. A process as claimed in claim 14 wherein, in the impurity removal stage, lime and calcium carbonate are added to the overflow portion to form a hematite/gypsum precipitate, which precipitate is then filtered and disposed of

16. A process as claimed in any one of claims 13 to 15 wherein the residue comprising the precious metal that is separated from the solution is subjected to a sulphur removal stage, prior to precious metal recovery.

17. A process as claimed in claim 16 wherein the sulphur removal stage comprises one or more of direct oxidation, thermal solvent extraction, flotation.

18. A process as claimed in any one of claims claim 14 to 17 wherein the precious metal recovery comprises a cyanide leaching process.

19. A process as claimed in any one of the preceding claims wherein the metal of the metal halide solution comprises copper and/or iron.

20. A process as claimed in any one of the preceding claims wherein the metal halide solution has a halide concentration of approximately 5-8 moles per litre.

21. A process as claimed in any one of the preceding claims wherein the halide comprises chloride or a mixture of halides comprising chloride and bromide.

22. A process as claimed in any one of the preceding claims wherein the precious metal to be recovered comprises gold.