Zinc recovery process.

The present invention relates to a process for the recovery of zinc from a zinc ore using a halide-based leaching solution. The invention also relates to a method for removing manganese from zinc and other metal halide solutions, and especially zinc and/or lead halide solutions. Further, the invention relates to a method for removing silver and/or mercury from metal halide solutions, especially zinc or cuprous chloride solutions (eg. in a hydrometallurgical process for zinc or copper production).

An existing process for zinc production involves roasting of a zinc containing ore, followed by acid sulphate leaching and then electrolytic recovery of zinc from the leachate. The roasting process produces sulphur-based pollutant gases which must be removed from the roasting furnace exhaust. In addition, if the ore contains high levels of impurities this can adversely affect the electrolysis and purity of zinc produced. For example, many zinc ores contain significant quantities of manganese. The existing process is constrained in its ability to treat zinc mineral concentrates containing significant manganese, as the manganese leaches into the leachate along with the zinc and cannot be effectively removed. Manganese causes problems in the zinc electrolytic recovery step by depositing on the anodes as MnO₂.

In Australian Patent 669906, the present applicants developed a multi-stage leaching process followed by electrolysis for the recovery of copper. Conversely, US 4,292,147 discloses a process for the recovery of zinc using chloride leaching followed by electrolysis.

The present inventors have also discovered that, when solution, manganese does not deposit on the anode during electrolytic metal recovery from the solution, especially with solutions derived from the leaching of zinc and/or lead ores. This allows for the manganese to be removed from the solution by other means.

Furthermore, in hydrometallurgical processes that have a high solution halide concentration (eg. from 200 to 300 grams per litre NaCl and 10 to 50 grams per litre NaBr), it is desirable to remove deleterious impurities prior to the electrowinning of various metals such as zinc or copper in the process. Many impurities can be removed from the electrolyte solutions of these processes by a stepwise pH adjustment and precipitation (eg. up to around pH 6), however, both silver and mercury are not so readily removed.

US4,124,379 discloses the removal of silver from a cuprous chloride electrolyte, whereby copper metal is added to the electrolyte to reduce cupric ions to cuprous ions followed by contact with an amalgam that exchanges a metal for the silver. The amalgam is formed from mercury metal with one of copper, zinc or iron, preferably copper shot coated or associated with mercury metal. The amalgam requires separate production, requiring a physical association with the mercury and copper, and this is cumbersome and complex, adding to the cost of the process. In addition, the amalgam must then be added to the process and contacted with the silver-bearing electrolyte, increasing the complexity and cost of running the process.

In a first aspect, the present invention provides a process for the recovery of zinc metal from a zinc mineral comprising the steps of: (i) leaching the zinc mineral in a single stage using a solution including a halide species formed from two or more different halides, to leach the zinc into the solution; (ii) passing the solution from (i) to an aeration stage in which air is introduced to oxidise and precipitate any iron present in the mineral; (iii) electrolysing the zinc-bearing solution from (ii) to yield zinc metal and to generate the halide species; and - (iv) returning the electrolysed solution including the halide species to the leaching step (i).

The present inventors have surprisingly discovered that halide species formed from two or more different halides have sufficiently high oxidising potential such that zinc (and other metals such as lead) can be directly leached into solution in a single stage without the need for a multi-stage leaching operation (for example, as disclosed in the applicant's Australian patent 669906).

Typically the halide species is anodically formed (ie. at the anode) in the electrolysis step. Typically the halide species is formed at an oxidation potential lower than that for the formation of many insoluble forms of impurities in the solution. For example, typically the halide species is formed at an oxidation potential lower than that for manganese dioxide. This enables the impurities to be maintained in solution, and then removed from the solution, either in-line or in a separate removal stage.

Preferably the two or more different halides are chlorine and bromine, and preferably the halide species is a soluble halide complex formed at the anode, such as BrCl₂^-, although BrCl gas may also be formed and then used. Such species have been discovered to be highly oxidising of zinc mineral (and other minerals) so that only a single-stage zinc leaching process is required.

Preferably the leaching of the mineral is facilitated by a catalyst, which typically catalyses the oxidation of the mineral by the halide species. In this regard, preferably the catalyst is a metal catalyst, such as copper, which can be present in the zinc ore or introduced into the leaching process (eg. as particulate copper in a first stage leaching).

Preferably when sulphur is present in the mineral, the process comprises the further step (iia) of adding limestone to balance the generation of sulphuric acid as a result of sulphur oxidation in the leaching step (i).

Preferably leachate from the leaching process has these solid precipitates separated therefrom and, prior to electrolysis, any gold and platinum group metals (PGM's) present therein are preferably removed by passing the leachate over activated carbon (typically a column thereof) to adsorb the gold and PGM's onto the carbon.

Preferably prior to electrolysis the leachate is then passed to a series of cementation processes in which zinc dust (typically using a portion of zinc produced in the electrolysis process) is added to the leachate to cement out any copper, silver, lead and other impurities present in the leachate (as a result of being present in the mineral).

Preferably prior to electrolysis and, if necessary, the leachate is then passed to a further iron (and residual metals) removal stage in which limestone and a halide species from the electrolysis step are added to oxidise and precipitate iron as ferric oxide. Preferably in this same stage, the limestone and halide species will oxidise and precipitate at least a portion of the manganese present as manganese dioxide.

Preferably a portion of the electrolysed solution (typically spent catholyte) is removed and processed to remove manganese therefrom. Preferably said portion is a bleed stream from a cathode compartment of an electrolytic cell for the electrolysis process, the bleed stream having added thereto limestone and the halide species from an anode compartment of the electrolysis process, to precipitate manganese dioxide.

In the manganese removal stage, preferably the pH and Eh of the solution are regulated in a manner that favours the formation of the manganese dioxide precipitate over the formation of a precipitate of zinc.

Preferably the pH is regulated by the incremental addition of the limestone to raise the solution pH to a level at which the Eh can be increased by the halide species to a level at or above which MnO₂ formation is favoured. Preferably the amount of limestone added is less than the stoichiometric amount required for MnO₂ formation, such that less of the zinc is precipitated.

Preferably prior to returning the electrolysed leachate to the leaching process and, if necessary, a portion of the catholyte (typically that in which manganese has been removed therefrom) is passed to a magnesium removal stage in which slaked lime is added to firstly remove any zinc (which is returned to the leaching process) and then to remove the magnesium as a magnesium oxide precipitate.

Therefore, in a second aspect the present invention provides a zinc recovery process including the removal of manganese from a zinc halide solution that is subjected to electrolysis to yield zinc metal, including the steps of:

• in the electrolysis, cathodically recovering the zinc metal whilst anodically forming a halide species from two or more different halides at an oxidation potential lower than that for the formation of manganese dioxide;
• removing a portion of the solution and processing it to remove manganese therefrom; and
• returning the processed portion to the zinc halide solution.

By forming a halide species at an oxidation potential lower than that for the formation of manganese dioxide the manganese can be maintained in solution, and this then enables it to be removed therefrom, typically in a separate stage.

The halide solution is a zinc halide leachate resulting from a leaching process in which a mineral concentrate is leached in a halide containing solution to leach zinc into solution.

The leachate is fed to the electrolysis step as electrolyte and, after electrolysis, is returned to the leaching process for further leaching of the mineral concentrate. In this regard, the manganese removal according to the method of the present invention is conducted as part of a closed loop mineral leaching and electrolytic recovery process.

Preferably the portion of solution that is removed and processed to remove manganese therefrom is a bleed stream from the closed loop mineral leaching and electrolytic recovery process, which stream is returned to that process after manganese removal.

The leaching of mineral is facilitated by the anodically formed halide species, which is returned with the electrolyte to the leaching process.

Preferably the two or more different halides are chlorine and bromine, and preferably the halide species is a soluble halide complex formed at the anode, such as BrCl₂^-, although BrCl gas may also be used.

Preferably the leaching of the mineral is facilitated by a catalyst, which typically catalyses the oxidation of the mineral by the halide species. In this regard, preferably the catalyst is a metal catalyst such as copper which is present in or introduced into the leaching process.

The mineral includes zinc or zinc and lead, zing being the metal yielded in the electrolysis step.

Preferably the manganese in said portion of solution is separated therefrom by the incremental addition of an alkali reagent, preferably a reagent that causes the manganese to precipitate as MnO₂ The MnO₂ precipitate can then be separated from the portion of solution before it is returned to the metal halide solution. Preferably the portion of solution is a portion of catholyte from the electrolysis step, typically spent catholyte.

Preferably the alkali reagent is calcium carbonate. When the metal to be recovered also has a tendency to precipitate with the addition of an alkali reagent such as calcium carbonate (eg.-such as zinc) preferably the alkali reagent is added in an amount less than the stoichiometric amount for MnO₂ formation. Preferably a high redox reagent is also added to increase the oxidation potential of the portion of solution to a level which favours the formation of MnO₂. In this regard, preferably the high redox reagent is the halide species (or a gaseous form thereof) anodically formed in the electrolysis step. Alternatively, the high redox reagent can be a hypochlorite or hypobromite salt (such as calcium hypochlorite) that is added to the portion of solution together with the alkali reagent.

Following from this, in a thirds aspect the present invention provides a zinc recovering process including the removal of manganese from a metal halide solution from which at least zinc metal can be yielded, including the steps of:

• regulating the pH and Eh of the solution in a manner that favours the formation of a manganese dioxide precipitate over the formation of a precipitate of the zinc metal; and
• removing the manganese dioxide precipitate from the solution.

Preferably the pH is regulated by the incremental addition of an alkali reagent to raise the solution pH to a level at which the Eh can be increased to a level at or above which MnO₂ formation is favoured. Preferably the amount of alkali reagent added is less than the stoichiometric amount required for MnO₂ formation, such that less of the at least one metal is precipitated.

Preferably the pH is raised by the addition of calcium carbonate, and preferably the Eh is raised by the addition of a high redox reagent as per the first aspect.

Preferably the at least one metal is zinc and/or lead.

Preferably the solution in the third aspect is the bleed stream of the second aspect. The solution of the third aspect is the leachate from the leaching process of the first aspect.

Notwithstanding any other forms which may fall within the scope of the present invention, preferred forms of the invention will now be described, by way of example, and also with reference to the accompanying drawing in which:

• Figure 1 shows a schematic process flow diagram illustrating a closed loop zinc leaching and electrolytic recovery process, including a manganese removal stage;
• Figure 2 shows a Pourbaix diagram for the Mn-Cl-H2O system at 25°C;

A preferred zinc recovery process according to the present invention is schematically depicted in Figure 1 and the description of this process will now be made with reference to Figure 1.

The process was designed to produce high purity zinc metal from complex mixed zinc/lead sulfide concentrates derived from zinc/lead ores. Lead and silver were produced as cement by-products while gold and platinum group metals (PGM's), when present, were produced as bullion.

Concentrates containing significant levels of iron were readily treated, with all leachable iron reporting to the leach residue as heamatite, while sulphide sulphur was converted to the elemental state. High levels of contaminants such as manganese and arsenic were also readily accommodated, with arsenic reporting to the leach residue as the environmentally stable iron arsenate and manganese rejected in a separate residue as manganese dioxide (described below).

The preferred process was based on the electrolytic deposition at the cathode of high purity zinc from a purified sodium chloride - sodium bromide electrolyte. During the electrowinning stage, a halide species was generated in solution at the anode. This species was a mixed halide species, such as soluble BrCl₂^- (hereinafter "Halex"). The species exhibited powerful leaching characteristics when it was recirculated to treat incoming concentrate feed.

The preferred process included three main steps of leaching, purification and electrowinning, as shown in Figure 1. The leaching step included a single leaching stage (reactor) and was combined with a series of subsequent stages (reactors). Zinc concentrate and an oxidant (the halide species) were fed to the leaching stage. Purification consisted of cementation and alkali precipitation steps. Electrowinning employed a number of diaphragm cells with the option for continuous product removal in dendritic form or the production of a conventional cathode plated with zinc.

Electrowinning was an integral part of the preferred process. Zinc metal was electrowon from purified electrolyte, which had a composition of 100 gpl zinc, 50 gpl sodium chloride (common salt, NaCl), 50gpl calcium chloride (CaCl₂) and 110 gpl sodium bromide (NaBr). All other constituents, including 'equilibrium' levels of many other elements (manganese, magnesium etc.), were regarded as impurities.

Electrolysis involved the passage of electric current at 500 A/m² of electrode area, to form high purity zinc on the negatively charged cathode. The feed electrolyte zinc concentration was in turn depleted from 100 to 50 gpl, which was the steady state concentration of the cell.

The spent catholyte continuously permeated through a woven cloth membrane M positioned in an electrolysis cell EC (Figure 1). The catholyte permeated to the positively charged electrode (anode).

Preferably chloride (Cl^-) and bromide (Br^-) were present in solution, and hence there was a preferential formation of the halide species BrCl₂^- (Halex). However, other halide species were able to be formed. This species was considered as a chlorine molecule held in solution by a bromide ion and was observed to be a very powerful lixiviant at an oxidising potential (Eh) of 1000 mV (vs Ag/AgCl). The resultant Halex-bearing solution (electrolyte) from the anode compartment (anolyte) was used for the leaching of zinc sulfide concentrates.

Zinc concentrate and Halex oxidant from the electrolysis process were fed to a single stage leach (the first continuously stirred tank reactor (CSTR) in Figure 1). The CSTR was operated at atmospheric pressure and to impart a temperature to the solution (electrolyte) of 85°C. The leaching (oxidation) reaction initially proceeded without air sparging (aeration (or introduction of air)) until all halex oxidant (BrCl₂^-) was consumed. Significant iron dissolution and oxidation of sulfide occurred in the first CSTR.

Second and third CSTR's were introduced for further aeration to reject iron as a ferric oxide precipitate. Any arsenic present was oxidised in the first CSTR and subsequently precipitated as the environmentally stable iron arsenate.

Recycle copper cement (ie. from a subsequent Cu & Ag cementation stage) was added to the first CSTR to enhance both oxygen uptake and metal extraction. The copper helped to catalyse the oxidation of the zinc concentrate by the halide species.

When leaching was complete limestone was added in a third CSTR to reject (precipitate) any dissolved iron that remained and any sulphate present. The limestone also balanced the generation of sulphuric acid as a result of sulphur oxidation in the first CSTR. Optionally a fourth CSTR was provided to increase residence (reaction) time and allow for maximum precipitation of iron and sulphate.

The leach residue was separated from the zinc-rich pregnant solution by filtration and washed before disposal to landfill at first solid-liquid separation station (S/L).

When gold or PGM's were present they were extracted (oxidised) during leaching and were recovered by feeding the pregnant zinc solution through a column housing activated carbon and onto which the gold and PGM's were adsorbed. This carbon could then be separately eluted with an elutant to recover the gold and other PGM's.

The zinc-pregnant solution was then further purified via a series of cementation reactions with zinc dust reagent. This was a two-stage CSTR facilitated operation, with copper and silver predominantly removed in the first stage, which operated at a temperature of 85°C. In the second stage, excess zinc dust was added to remove the remaining impurities such as cadmium, lead, nickel, cobalt, thallium etc.

After two further solid-liquid separation stages, the now relatively pure zinc laden solution had added thereto, in a CSTR, a small amount of halex vapour from the electrolytic cell anolyte, until an Eh of 700mV was achieved. This ensured that any remaining iron was oxidised to the ferric (Fe³⁺) oxidation state. A subsequent addition of ground limestone to raise the pH to 4.5 precipitated most remaining impurities (eg. Bi, Fe, In, Ge, etc.). These precipitates were removed by filtration and were either discarded or reprocessed to recover economic by-products such as indium.

It was also apparent that at least some of the manganese was removable from the metal halide solution as a manganese dioxide precipitate if the pH was regulated to above 3.2 in this CSTR stage. The pH was raised above this level by the addition of calcium carbonate (typically ground limestone). At this pH, and at a redox potential of 600-800mV vs Ag/AgCl (800-1000mV vs SHE), the manganese entered the favourable MnO₂ stability region of a Pourbaix diagram. The manganese dioxide precipitate was then removed with the precipitated iron from the solution. The Eh was also adjusted (raised) by the addition of the high redox reagent Halex. The specific chemical reactions and reaction conditions for manganese removal are described in detail shortly in the following text for the subsequent manganese removal circuit (for a spent catholyte bleed stream) which also makes use of limestone and halex for this purpose. It is noted that, in the bleed stream to be described, typically the manganese concentration in solution is higher than it was in the present CSTR purification stage.

Any impurities now remaining were observed to not contaminate zinc during electrowinning and were removed in either the manganese or magnesium purification circuits. In particular, because the halide species was formed at the anode at an oxidation potential lower than that for the formation of manganese oxide, this metal did not contaminate the electrolytically recovered zinc.

The purified zinc solution was then electrolysed, as described above, to produce high purity zinc and to regenerate the halide species lixiviant for recycle to the leach. A zinc product was washed and dried under an inert atmosphere prior to melting in a furnace. Some zinc dust produced from zinc in the furnace was used in the cementation stages (described above).

These removal circuits treated a spent catholyte bleed stream, with limestone and halex used in the manganese circuit, and with slaked lime used in the magnesium circuit.

This process was observed to be generally applicable in the treatment of mineral concentrates containing manganese. In the present zinc recovery process the manganese leached into the leachate was removed in equal proportion (ie..removed at least at the rate of its leaching into the leachate) to prevent its accumulation in the process solution.

Up to a maximum of 20% of the total solution flow through the electrolytic cell EC was drawn off as a bleed stream BS from the spent catholyte. This stream was treated with Halex vapour and limestone to precipitate the manganese as manganese dioxide MnO₂.

The vapour pressure of Halex above the electrolyte was observed to be temperature dependent, and the vapour itself likely consisted of BrCl gas.

The manganese removal process is chemically depicted below in equations (1) to (3), where the Halex vapour was assumed to be BrCl.

The electrolyte composition in the spent catholyte was approximately 50gpl NaCl, 110gpl NaBr, 50gpl CaCl₂, and 50gpl Zn²⁺. However, the manganese removal process was also able to be performed on the process electrolyte prior to electrolysis. When using a bleed stream from the spent catholyte of the zinc electrowinning cell, the manganese level in solution depended on the quantity leached into solution and the size of the bleed stream going to Mn removal.

The present inventors demonstrated that manganese was completely removed from the bleed stream as manganese dioxide by the addition of Halex vapour thereto. The electrolyte-vapour reaction was discovered to be extremely rapid in the presence of incrementally added small quantities of limestone.

The manganese precipitation as manganese dioxide is predicted by the Pourbaix diagram shown in Figure 2. As can be seen, manganese does not precipitate at neutral redox potential until above pH 7.5. At high redox potential, 600-800mV vs Ag/AgCl (800-1000mV vs SHE) and pH 3.2, however, the manganese entered the MnO₂ stability region and was able to be precipitated as MnO₂.

A small amount of zinc was also precipitated during the manganese removal. However, the present inventors demonstrated that less zinc was precipitated at lower aqueous zinc levels (eg 25gpl Zn²⁺ instead of 50gpl Zn²⁺), or when the limestone was incrementally added in small amounts to maintain a pH of around 3.2 (as opposed to precipitating the.manganese in-the presence of a stoichiometric excess of limestone).

Thus, by using a spent catholyte containing 50gpl Zn²⁺ and 15gpl Mn²⁺, with the slow addition of limestone, only 7.8% of the zinc in solution was co-precipitated with the manganese. For an 8% catholyte bleed stream, this equated to the loss of less than 0.7% of the total zinc production to the bleed stream precipitate.

Examples of successful manganese removal experiments are shown below. As can be seen, complete removal of the manganese was effected in the first two experiments, while complete removal was not attempted in the third experiment. It should be noted that the intermediate zinc assays in solution were only an approximate guide, as the zinc precipitated from solution was difficult to distinguish from the experimental error associated with this assay. Accurate determinations of the zinc precipitation were determined by direct analysis of the final residue.

The results from Experiments 2 and 3, shown in Tables 2 and 3, reflect the way that the technique was applied to a lead/zinc process in which Halex was anodically formed. The results indicated that for an 8% bleed from the spent catholyte a recirculating manganese concentration at 15gpl was maintained. The co-precipitation of zinc with the MnO₂ precipitation represented a loss of approximately 0.6% of the total zinc production.

Non-limiting examples of manganese removal processes will now be described.

An electrolyte comprising 200gpl NaCl, 150 gpl NaBr, 30 gpl Ca²⁺, 25 gpl Zn²⁺ and 5gpl Mn²⁺ was prepared as a typical bleed stream from a lead/zinc recovery process. 25gpl CaCO₃ was added to the electrolyte in a reaction vessel at the commencement of the removal process. Ca(OCl)₂ was added directly to the electrolyte to oxidise the solution. The temperature of the solution was maintained at a typical process electrolyte temperature of 60°C to 65°C. The results are shown in Table 1 below.

An electrolyte comprising 50gpl NaCl, 110gpl NaBr, 50gpl CaCl₂, 50gpl Zn²⁺ and 15gpl Mn²⁺ was prepared as a typical bleed stream from a lead/zinc recovery process. 85gpl CaCO₃ was added to the electrolyte in a reaction vessel at the commencement of the removal process. Halex vapours were generated externally and pumped into the reaction vessel. The temperature of the solution was maintained at a typical process electrolyte temperature of 60°C to 65°C. The results are shown in Table 2 below.

An electrolyte comprising 50gpl NaCl, 110gpl NaBr, 50gpl CaCl₂, 50gpl Zn²⁺ and 15gpl Mn²⁺ was prepared as a typical bleed stream from a lead/zinc recovery process. CaCO₃ was added to the electrolyte in a reaction vessel in small doses throughout the removal process. Halex vapours were generated externally and pumped into the reaction vessel. The temperature of the solution was maintained at a typical process electrolyte temperature of 60°C to 65°C. The results are shown in Table 3 below.

10.0g of the manganese precipitate from Example 2 was added to 1.0 litre of demineralised water in a reaction vessel. H₂SO₄ was added. The results are shown in table 4 below.

10.0g of precipitate from Example 3 were added to 1.0 litre of demineralised water in a reaction vessel. Halex vapours were generated externally and pumped into the reaction vessel. H₂SO₄ was added. The results are shown in Table 5 below.

Manganese was thus completely removed from the bleed stream as manganese dioxide by the addition of Halex vapour. The electrolyte-vapour reaction was discovered to be extremely rapid in the presence of small quantities of limestone.

A small amount of zinc was also precipitated during the manganese removal. However, less zinc was precipitated at lower aqueous zinc levels (eg 25gpl Zn²⁺ instead of 50gpl Zn²⁺), or when the limestone was added incrementally in small amounts to maintain a pH of around 3.2 (as opposed to precipitating the manganese in the presence of a stoichiometric excess of limestone).

Thus, by using a spent catholyte containing 50gpl Zn²⁺ and 15gpl Mn²⁺, with the slow addition of limestone, only 8.4% of the zinc in solution was co-precipitated with the manganese. For an 8% catholyte bleed stream, this equated to the loss of less than 0.7% of the total zinc production to the bleed stream precipitate.

In the magnesium removal process, spent catholyte from the manganese removal process had, in a series of CSTR's, slaked (or caustic) lime added thereto to cause the precipitation in the first two CSTR's of zinc. This zinc was then returned to the leaching process, typically at the limestone addition stage (as shown in Figure 1). Then, in two further CSTR's more slaked lime was added to cause the precipitation of magnesium oxide and thereby prevent the build up of magnesium in the process. This removal step was only required intermittently, because of the generally lower levels of magnesium in the zinc mineral. The treated spent catholyte was then returned to the leaching process, typically to the first leaching CSTR.

The economics of the preferred zinc recovery process in which Halex was anodically formed were noted be site specific. For a concentrate producer, the loss of 0.6% of zinc production as a result of impurity removal (especially of manganese) represented a significant improvement over prior art smelter losses, where approximately 4% of the contained zinc was deducted before calculating the payable metal content. However, where higher zinc recoveries in the process were desirable, some of the zinc precipitated from solution with the manganese was recovered by re-slurrying the bleed stream precipitate in water and acidifying with sulfuric acid. The zinc was able to be redissolved with approximately 90% selectivity. Allowing for the higher bleed stream ratio required to account for the partial redissolution of the manganese, the zinc loss with the bleed stream precipitate was reduced to 0.1% of total zinc production.

An important feature of the preferred process was that all impurities including manganese, mercury and arsenic were either recovered as saleable by-products or stabilised for disposal to tailings. Another equally important feature was that heat was provided by the exothermic leach reactions. This, coupled with the addition of air to the leach, evaporated water and thereby maintained the water balance at neutral. As such, there was no liquid effluent.

The zinc recovery process according to the present invention had the following clearly identifiable advantages over the prior art roast/leach/electrowin zinc process and other hydrometallurgical processes:

• Significantly lower operating and capital costs;
• Recovery of precious metals in the zinc leach circuit (eg. in the purification of the non-zinc metals from solution, the recovery of saleable copper, lead and silver metals (Cu, Ag, Pb and Fe in the cementation and precipitation stage of Figure 1));
• Tolerance to low grade and "dirty" concentrates;
• Low energy consumption;
• No liquid emissions;
• No production of noxious gases;
• Mild operating conditions of low temperature and atmospheric pressure;
• No need for solvent extraction; and
• No requirement for pure oxygen in aeration stages.