Zinc recovery process

<td>REPUBLIC POLAND</td><td>(12) PATENT DESCRIPTION (19) PL (21) Application number: 369574</td><td>(11) 201418 (13) B1</td>

<td>γίγ</td><td>(22) Date of notification: September 12, 2002</td><td>(51) Int.Cl. C22B 3/04 (2006.01)</td>

<td></td><td>(86) Date and number of the international application:</td><td>C22B 19/20 (2006.01)</td>

<td></td><td>2002-09-12, PCT / AU02 / 01260</td><td>C22B 3/20 (2006.01)</td>

<td>patent Office</td><td>(87) Date and publication number of the international application:</td><td>C25C 1/16 (2006.01)</td>

<td>Polish Republic</td><td>2003-03-20, WO03 / 023077 PCT Gazette No. 12/03</td><td></td>

(54) A method for recovering metallic zinc from the zinc mineral

<td>(30) Priority:</td><td>(73) The right holder of the patent:</td>

<td>2001-09-13, AU, PR7669</td><td>INTEC LTD, Sydney, AU</td>

<td>2001-09-13, AU, PR7667 2001-09-13, AU, PR7670</td><td>(72) Inventor (s):</td>

<td></td><td>John Moyes, Sydney, AU</td>

<td>(43) Application announced:</td><td>Frank Houllis, Sydney, AU</td>

<td>02.05.2005 BUP 09/05</td><td>(74) Representative:</td>

<td>(45) The grant of the patent was announced:</td><td>Łazewska Sławomir,</td>

<td>April 30, 2009 WUP 04/09</td><td>Łazewska i Łazewski Sp.j.</td>

⁽⁵⁷⁾ The invention relates to a process for recovering metallic zinc from a zinc mineral comprising leaching the zinc mineral in a single step using a solution containing a halide compound formed of two or more different halides to leach the zinc into solution, subjecting the zinc-containing solution to electrolysis to obtain metallic zinc, and to produce halide compound, and recycling the electrolysed solution containing the halide compound to the leaching step. Part of the electrolysed solution can be removed as a draw stream from the electrolyser cathode chamber and processed to remove manganese as manganese dioxide precipitate by adding limestone and a mixture of halides from the anode chamber of the electrolyser to this stream. In view of this, the pH and Eh values ​​of the solution can be adjusted in such a way that the precipitation of manganese dioxide is preferable to that of zinc.

PL 201 418 B1

Description of the invention

The present invention relates to a method for recovering metallic zinc from a zinc mineral using a halide-containing leaching solution.

The existing zinc production process involves the calcination of the zinc containing ore followed by a sulfuric acid leach followed by electrowinning the zinc from the leachate. The calcining process produces sulfur-containing gaseous pollutants which must be removed from the off-gases of the calcination furnace. Additionally, if the ore has a high level of impurities, it may have an adverse effect on the electrolysis and purity of the zinc obtained. For example, many zinc ores contain significant amounts of manganese. The current process is limited in the treatment of zinc mineral concentrates with a significant manganese content, since the manganese is leached into the leachate simultaneously with the zinc and cannot be effectively removed. Manganese causes problems in the zinc electrowinning step as it deposits on the anodes as MnO2.

In Australian Patent No. 669,906, the present applicant has developed a multi-stage leaching process followed by electrolysis to recover copper. In contrast, US Patent No. 4,292,147 discloses a zinc recovery process using chloride leach followed by electrolysis.

The present applicant has also found that if certain conditions are maintained in the metal halide solution, manganese does not deposit on the anode during electrowinning of the metal from the solution, in particular from solutions derived from the leaching of zinc and / or lead ores. This allows manganese to be removed from the solution by other means.

Furthermore, in hydrometallurgical processes where the solution has a high concentration of a halide (e.g., 200 to 300 grams per liter of NaCl and 10 to 50 grams per liter of NaBr), it is desirable to remove undesirable impurities prior to electrowinning various metals such as zinc or copper. Many impurities can be removed from the electrolyte solutions in these processes by gradual pH adjustment and precipitation (e.g., to about pH = 6), however, neither silver nor mercury are removed immediately.

US 4,124,379 discloses the removal of silver from a copper (I) chloride electrolyte, wherein metallic copper is added to the electrolyte to reduce the copper (II) ions to copper (I) ions, and then contacted with an amalgam that exchanges the metal for silver. The amalgam is made of metallic mercury with one of the metals copper, zinc or iron, preferably copper, sprayed on or combined with metallic mercury. The amalgam has to be obtained separately, requiring the physical association of mercury with copper, which is cumbersome and complex, and adds to the cost of the process. In addition, the amalgam must then be added to the process and contacted with the silver-containing electrolyte, which increases the complexity and cost of the process.

The present invention relates to a method of recovering metallic zinc from a zinc mineral, characterized in that it comprises:

- leaching the zinc mineral in a single step using a solution containing a halide compound formed from two or more different halides to leach the zinc into solution,

- subjecting the zinc-containing solution to electrolysis to obtain metallic zinc and to produce a halide compound, and

- recycling the electrolysed solution containing the halide compound to the leaching step.

Preferably, the halide compound is produced at the anode in an electrolysis step.

Preferably, the halide compound is produced at an oxidation potential lower than the oxidation potential whereby insoluble forms of impurities are formed in solution.

Preferably, two or more different halides are chlorides and bromides and the halide compound is a soluble halide complex.

Preferably, the halide compound is BrCl^-2 and / or gaseous BrCl.

In another preferred embodiment of the invention, the leaching of the mineral with a halide compound is performed using a catalyst that catalyzes the oxidation of the mineral.

Preferably, the catalyst is a metal catalyst that is present in the zinc mineral or is fed into the leaching process.

Preferably, the catalyst is copper.

PL 201 418 B1

In another preferred embodiment of the invention, the process further comprises one or more additional steps allowing the zinc to be leached in one step and the impurities removed in the additional step (s).

Preferably, the method comprises at least two additional steps which are an aeration step and a sulfate precipitation step.

Preferably, air is introduced in the aeration stage to oxidize and precipitate any iron present in the mineral, and in the sulphate precipitation stage (VI) limestone is added to precipitate sulphate (VI) resulting from the sulfur oxidation in the leaching stage as sulphate ( VI) calcium.

In another preferred embodiment of the invention, the leachate from the pre-electrolysis leachate, if it contains any amount of gold or platinum group metals, is removed by passing the leachate through activated carbon to adsorb gold and platinum group metals thereon.

In yet another preferred embodiment of the invention, the leachate effluent prior to electrolysis is subjected to a series of cementation processes in which zinc dust is added to the leachate to cement any amounts of copper, silver, lead and other impurities present in the effluent.

In a further preferred embodiment of the invention, the leachate effluent prior to electrolysis is subjected to a further iron removal process in which limestone and the halide compound from the electrolysis step are added to oxidize and precipitate the iron as iron (III) oxide.

In another preferred embodiment of the invention, part of the electrolysed solution is removed and processed to remove manganese therefrom.

Preferably, said portion of the solution is a leach stream from the cathode space of the electrolyser, limestone and an anode space halide compound being added to the elution stream to precipitate manganese dioxide.

Preferably, the pH and Eh of the solution are controlled in a manner that favors manganese dioxide precipitation rather than zinc precipitation.

Preferably, the pH is adjusted by the gradual addition of limestone to raise the pH of the solution to a level where the Eh can be raised using a halide compound to a level at or above which MnO2 formation is preferred.

Preferably, less calcium is added than the stoichiometric amount required to form MnO2, such that less zinc precipitates.

In another preferred embodiment of the invention, before the electrolysed effluent is returned to the leaching process, part of the electrolysed effluent from the cathode space is routed to a magnesium removal step where slaked lime is added to first remove zinc which is recycled to the leaching process and then removed. Magnesium is formed in the form of a precipitate of magnesium oxide.

In yet another preferred embodiment of the invention, the manganese contained in said part of the solution is separated therefrom by the gradual addition of an alkaline reagent which causes the manganese to precipitate as MnO2, whereafter the precipitated MnO2 is separated from that part of the solution before it is returned to it. the leaching stage. Preferably, the alkaline reagent is calcium carbonate.

Preferably, if the recovered zinc and / or lead also tends to precipitate upon addition of the alkaline reagent, the alkaline reagent is added in an amount less than the stoichiometric amount needed to form MnO2.

In another preferred embodiment of the invention, a reagent with a high redox potential is also added to said part of the solution to increase the oxidation potential in that part of the solution to a level where the formation of MnO2 is preferred.

Preferably, the high redox potential reagent is a halide compound or its gaseous form formed at the anode in the electrolysis step.

In another preferred variant of the invention, said solution portion is part of the catholyte of the electrolysis step.

It has surprisingly been found that a halide compound formed from two or more different halides has a sufficiently high oxidation potential such that zinc (and other metals such as lead) can be directly leached into solution in a single step without the need for a multi-step leaching operation ( for example, as disclosed in Applicant's Australian Patent No. 669,906).

PL 201 418 B1

Typically the halide compound is formed at the anode in the electrolysis step. Typically a halide compound is formed with an oxidation potential lower than that of manganese dioxide.

This allows contaminants to be retained in solution and then removed from the solution either on an ongoing basis or in a separate removal step.

Such halide compounds have been found to strongly oxidize zinc minerals (and other minerals) such that only a one-step zinc leaching process is required.

The catalyst may be present in the zinc ore or may be introduced into the leaching process, e.g. as copper ground in a first leach step.

The zinc dust added to the effluent in cementation processes may be part of the zinc obtained by the electrolysis process. Copper, silver, lead and other impurities removed in the cementation process are present in the solution due to their presence in the mineral.

The limestone and halide compound added in the iron removal step will oxidize and also cause precipitation of at least some of the manganese present as manganese dioxide.

The mineral treated according to the invention comprises zinc and / or lead as the metal (s) obtained in the electrolysis step.

Due to the formation of the halide compound at an oxidation potential lower than the oxidation potential for the formation of manganese dioxide, the manganese can be retained in solution and this in turn allows it to be removed from the solution, usually in a separation step. Said manganese removal process can be carried out as part of a closed cycle process for the mineral leaching and electrowinning of the metal. Alternatively, in addition to the halogen compound, the high redox reagent used in the manganese removal process may be chlorate (I) or bromate (I) (such as calcium hypochlorite) which is added to part of the solution along with the alkaline reagent.

Preferably, the pH is adjusted by gradually increasing the addition of an alkaline reagent to raise the pH to a level where the Eh value can be raised to a level at or above which the formation of MnO2 is preferred. Preferably, the amount of added alkaline reagent is less than the stoichiometric amount needed to form MnO2 such that less of the at least one metal is precipitated.

The process of the invention also allows removal of dissolved silver and / or mercury from a solution of a metal halide, which metal is capable of forming a deposit with silver and / or mercury. In order to achieve this, it is possible to lower the Eh of the solution with the help of a reducer to the level at which metal precipitation occurs, add an ionic compound to the solution, which reacts with the precipitated metal in such a way that it causes the silver and / or mercury to form a precipitate with the precipitated metal. metal, and remove sediment.

The formation of a metal precipitate and the addition of the ionic compound cement the silver and / or mercury without the need to form and then introduce a separate amalgam, which is unique to the present invention.

Such a process is typically used for the simultaneous removal of silver and mercury, and is of particular use in the treatment of chloride solutions obtained from ores of minerals containing both silver and mercury (which are often found in nature together in many crude ores).

The most preferred application of the process of the invention applies to solutions of zinc chlorides as well as copper (such as from the leaching process of zinc or copper ore). Thus, the metal is then zinc and / or copper and the halide is chloride, and the method is preferably carried out with the copper in the first oxidation state. In at least preferred embodiments of the process, it is desirable that silver and mercury be removed with only a minimal amount of copper as said metal precipitated.

The following Eh values ​​are referenced to the potential of a standard Ag / AgCl electrode. Preferably, the Eh of the solution is lowered to less than 0 mV. The Eh value can be lowered to about -200 mV, but is typically lowered to about -150 mV. A typical copper (I) chloride-containing process electrolyte has an Eh of about +150 mV, and therefore, when the Eh of the solution is lowered to -150 mV, an additional 300 mV driving force of silver precipitation is obtained.

Preferably, the Eh value is lowered by adding to the solution a reducing agent selected from one or more: aluminum metal, zinc metal, metal iron or metal hydride, metal borohydride or metal (tetraoxo) disulfate. Metal aluminum is the most preferred reducing agent because it is economical to use, readily available, and produces a precipitate that is easier to filter through.

PL 201 418 B1

Sodium (tetraoxo) disulfate (III) may also be used as reducing agent to lower the Eh value of the solution, optionally with a final reduction in the Eh value achieved with sodium borohydride (i.e. at the end of the Eh lowering step). By using a non-copper reducer, any mercury present in the removal is not alloyed with more copper.

Preferably the added ionic compound is an ionic mercury compound in both the first and second oxidation states (Hg (I) or Hg (II)), or both at the same time. Other ionic compounds include gold ionic compound etc. which is less economical. The use of an ionic compound such as an ionic mercury compound is preferred over the use of metallic mercury due to the problems defined above in US 4,124,379.

Typically the precipitated mercury is removed from the solution by passing an inert gas through the solution so that at least some of the mercury is removed therefrom along with the gas. For example, an inert gas may be dispersed into solution, typically an inert gas such as nitrogen is used (i.e., due to its availability and low cost).

Once removed from the solution, the inert gas is usually washed to remove the mercury therein before being returned to solution for reuse. Accordingly, the inert gas may be scrubbed with a copper (II) solution (i.e. containing a copper (II) ion) to remove mercury from the gas. Other oxidants such as mercury (II) chloride can be used to leach mercury from nitrogen. Carbon dioxide can also be used.

In addition, the removal of metallic mercury from the solution can be enhanced by heating the solution to cause the additional metallic mercury to evaporate. When an inert gas stripping process is used, an increased amount of vaporized metallic mercury is entrained in the gas bubbles.

Alternatively, metallic mercury can be removed from the solution by passing the solution over activated carbon so that the metallic mercury is adsorbed onto the carbon. In this alternative method, mercury can then be removed from the activated carbon by passing a second solution over the activated carbon, the second solution having an Eh value high enough to convert the adsorbed mercury to an ionic mercury compound to thereby dissolve it in the second solution.

This alternative method is readily applicable to a continuous process and hence more commercially feasible, e.g. an activated carbon column may be used, and then an electrolyte with a reduced Eh value (e.g. down to about -150 mV) may be pumped through the column at continuously. When a limited amount of mercury has been adsorbed onto the carbon, an automatic switch to pump the second solution through the column can be used. The dissolved mercury can be immediately recovered from the second solution.

For the optional removal of silver and / or mercury, zinc chloride and / or copper (I) chloride solutions present in the hydrometallurgical metal recovery processes are used. A typical Eh value in hydrometallurgical processes using such solutions and prior to silver and / or mercury removal is about +150 mV.

Figure 1 is a simplified process flow diagram illustrating the closed cycle of the zinc leaching and electrowinning process, including the manganese removal step, Figure 2 is a Pourbaix diagram for the Mn-Cl-H2O system in Figure 1. 25 ° C, Fig. 3 is a graph of silver and mercury removal as a function of process progress, and Fig. 4 is a graph of mercury removal from an electrolyte containing Cu ions⁺ using nitrogen scattering as a function of time.

EXAMPLES Recovery of zinc

A preferred zinc recovery method according to the present invention is schematically illustrated in Fig. 1, and a description of this method will now be given with reference to Fig. 1.

The process was developed to obtain high purity zinc metal from complex mixed concentrates of zinc and lead sulphides derived from zinc / lead ores. Lead and silver were obtained as by-products of cementation, while gold and platinum group metals (PGM), when present, were obtained as an alloy with silver.

Concentrates containing significant amounts of iron were processed immediately, with all leachable iron going to the leach residue in the form of hematite, while

The sulfur was converted from the sulfide form to the elemental sulfur form. It also adjusted straight to the high level of contaminants such as manganese and arsenic, with arsenic passing into the leaching solution as environmentally stable iron arsenate and the manganese separated as a separate residue as manganese dioxide (described below).

The preferred method is based on electrolytic deposition on a high purity zinc cathode from purified sodium chloride-sodium bromide electrolyte. During the electrowinning step of the metal preparation, the halide compound was formed in the solution at the anode. This complex was a mixed halide compound such as soluble (hereinafter referred to as Halex). The compound showed very favorable leaching characteristics when recycled to feed concentrate processing.

The preferred process included three main steps of leaching, purification and electrowinning as shown in Figure 1. The leaching step comprised a single leaching step (reactor) and was coupled to a series of successive steps (reactors). The zinc concentrate and the oxidant (halide compound) were passed to the leaching stage. The purification consisted of the steps of cementation and basic precipitation. Electrowinning of metal involved a series of diaphragm cells with the ability to continuously remove the dendritic product or to produce a conventional zinc-coated cathode.

Zn electrolysis (zinc electrowinning)

Electrolytic zinc deposition was an integral part of the preferred method. Metallic zinc was electrolytically deposited from a purified electrolyte, which contained zinc in the amount of 100 g per liter, sodium chloride (normal salt NaCl) in the amount of 50 g per liter, calcium chloride (CaCl2) in the amount of 50 g per liter, sodium bromide (NaBr) in 110 grams per liter. All other ingredients, including the equilibrium levels of many other elements (manganese, magnesium, etc.), were considered impurities.

The electrolysis consisted of passing an electric current at 500 A / m through the electrode space to form high purity zinc on the negatively charged cathode. The zinc concentration in the feed electrolyte was reduced from 100 to 50 grams per liter, which was the steady state concentration in the electrolyser.

Oxidizer (leaching agent)

The spent catholyte permeates continuously through the fabric membrane M placed in the electrolyser EC (Fig. 1). The catholyte penetrates into the positively charged electrode (anode).

Chlorides (Cl^-) and bromides (Br^-) and hence favorably formed the halide compound BrCl2^- (Halex). However, other halide compounds could also be formed. This complex was considered to be a chlorine molecule held in solution by the bromide ion and found to be a very strong leach with an oxidation potential (Eh) of 1000 mV (versus Ag / AgCl). The obtained solution containing Halex (electrolyte) from the anode space was used to leach zinc sulphide concentrates.

Leaching stage

The zinc concentrate and Halex oxidant from the electrolysis process were subjected to a single leaching step - the first continuously stirred tank reactor (CSTR) in Figure 1). The CSTR was operated under atmospheric pressure and the temperature of the solution (electrolyte) was adjusted to 85 ° C. The leaching (oxidation) process initially took place without oxygen dissipation (aeration or air injection) until all of the Halex oxidant was consumed. Significant iron dissolution and sulfide oxidation occurred in the first CSTR.

Second and third CSTRs were introduced for further aeration to remove iron as iron (III) oxide precipitate. Any arsenic present was oxidized in the first CSTR and then precipitated as environmentally stable iron (III) arsenate.

The recycled copper containing sludge (i.e. from the next Cu and Ag cementation step) was added to the first CSTR to enhance oxygen uptake and metal extraction. The copper assisted in catalyzing the oxidation of the zinc concentrate by the halide compound.

After leaching was complete, limestone was added to the third CSTR to remove (precipitate) any remaining dissolved iron and any sulfate present. The limestone also made it possible to counterbalance the formation of sulfuric acid (VI) from sulfur oxidation in the first CSTR. A fourth CSTR was optionally used to increase the residence time (reaction) and to allow the precipitation of iron and sulfate (VI) as much as possible.

PL 201 418 B1

Purification step

The leach residue was separated from the zinc-containing mother liquor by filtration and washed in a first solids separation station (S / L) prior to disposal to landfill.

If gold or platinum group metals were present, they were extracted (oxidized) during the leaching and recovered by passing the zinc-containing mother liquor through an activated carbon column onto which the gold and platinum group metals were adsorbed. This carbon could then be washed separately with an eluent to recover the gold and platinum group metals.

The zinc-containing mother solution was then subjected to further purification through a series of cementation reactions with the zinc dust reagent. This was a two step CSTR operation with the copper and silver mostly removed in the first 85 ° C step. In a second step, excess zinc dust was added to remove the remaining impurities such as cadmium, lead, nickel, cobalt, thallium etc.

After a further two liquid-solid separation steps, a small amount of Halex fumes from the anolyte from the electrolyser were introduced into the relatively pure solution with a high concentration of zinc in the CSTR until an Eh of 700 mV was reached. This ensured that any remaining iron was oxidized to iron in the third oxidation state (Fe³⁺). Then, ground limestone was added to raise the pH to 4.5, which caused the precipitation of most of the remaining impurities (e.g., Bi, Fe, In, Ge, etc.). These precipitates were removed by filtration and either discarded or reprocessed to recover valuable by-products such as indium.

It was also seen that it was possible to remove at least some of the manganese from the manganese dioxide precipitate from the metal halide solution if the pH was adjusted above 3.2 in the CSTR step. The pH was raised above this level by adding calcium carbonate (usually - ground limestone). At this pH value and at a redox potential of 600-800 mV against an Ag / AgCl electrode (800-1000 mV against a standard hydrogen electrode [SHE]), manganese was in the preferred MnO2 stability region in the Pourbaix diagram. The manganese dioxide precipitate was then removed from the solution together with the precipitated iron. The Eh value was also adjusted (raised) by adding a high redox potential reagent which was Halex. The specific chemical reactions and reaction conditions for manganese removal are described briefly but in detail in the following text relating to the subsequent manganese removal circuit (for a spent catholyte washout stream) which also uses limestone and Halex for this purpose. It is found that in the described elution stream, the concentration of manganese in the solution is usually higher than in the present CSTR purification step.

It was observed that any residual impurities did not contaminate the zinc during its electrowinning and were removed in the manganese purification circuit or the magnesium purification circuit. In particular, since the halide compound was formed at the anode at an oxidation potential lower than the formation potential of manganese oxide, the metal did not contaminate the electrowinning zinc.

The purified zinc-containing solution was then subjected to electrolysis as described above to obtain high purity zinc and to regenerate the leaching agent, the halide compound, for recycling it to the leaching. The obtained zinc was washed and dried under an inert gas atmosphere before melting it in a furnace. Part of the zinc dust obtained from the zinc in the furnace was used in the cementation steps (described above).

Removal of manganese and magnesium

These stripping circuits treat the spent catholyte wash stream, limestone and Halex in the manganese removal circuit and slaked lime in the magnesium removal circuit.

Generally the process has been found to be applicable to the treatment of manganese containing mineral concentrates. In the present zinc recovery process, manganese leached into the effluent was removed in an equal ratio (i.e., removed at least as fast as it was leached into the effluent) to prevent manganese build-up in the process solution.

Up to a maximum of 20% of the total solution stream flowing through the EC electrolyser was withdrawn as BS wash stream from the spent catholyte. This stream was treated with Halex fumes and limestone to precipitate the manganese as MnO2 manganese dioxide.

PL 201 418 B1

The Halex vapor pressure above the electrolyte was found to be temperature dependent and the vapor itself preferably consisted of BrCl gas.

The manganese removal process is shown below by the chemical equations (1) to (3), assuming that Halex is BrCl.

Me²+ + 2H₂O &lt;: · MnO₂ + 4H + + 2e^- (1)

BrCl + 2e^- q Br^--+ Cl^-_(2)

2H₂O + Me²+ + BrCl&gt; MnO₂ + Br^- + Cl^- + 4H + (3)

2H₂O + Me²⁺ + BrCl&gt; MnO₂ + Br^- + Cl^- + 4H + (3)

2CaCO₃+ 4H +&gt; 2Ca²+ + 2CO₂ + 2H₂AT 4)

Me²⁺ + 2CaCO3 + BrCl&gt; MnO2 + 2Ca²⁺ + Br^- + Cl^- + 2CO2 (5)

The electrolyte composition of the spent catholyte was about 50 grams per liter of NaCl, about 110 grams per liter of NaBr, about 50 grams per liter of CaCl2, and Zn ions²⁺ at about 50 grams per liter. However, the manganese removal process can also be subjected to the process electrolyte prior to electrolysis. When a spent catholyte leach stream from a zinc electrowinning cell is used, the level of manganese in the solution is dependent on the amount leached into the solution and the size of the leach stream directed to the Mn removal.

The present inventors have shown that the manganese was completely removed as manganese dioxide from the wash stream by adding Halex vapors thereto. It has been found that the electrolyte-vapor reaction is extremely rapid in the presence of small amounts of limestone added gradually.

The precipitation of manganese as a manganese dioxide precipitate is predicted in the Pourbaix diagram shown in Figure 2. As can be seen, manganese does not precipitate at a neutral redox potential until the pH is above 7.5. However, at a high redox potential of 600-800 mV against an Ag / AgCl electrode (800-1000 mV against a standard hydrogen electrode [SHE]) and at a pH of 3.2, manganese enters the MnO2 stability region and may precipitate as MnO2.

A small amount of zinc was also precipitated during the removal of manganese. However, as demonstrated by the present inventors, less zinc precipitated at lower levels of zinc in water-soluble form (e.g. at 25 grams per liter of Zn instead of 50 grams per liter), or when limestone was added in small, incremental amounts. to maintain the pH around 3.2 (as opposed to manganese precipitation in the presence of a stoichiometric excess of limestone).

Thus, when using a spent Zn-containing catholyte²⁺ at 50 grams per liter and Mn at 15 grams per liter, and by slowly adding limestone, only 7.8% of the zinc in the solution was co-precipitated with the manganese. For the 8% catholyte leach stream, this equates to a loss of less than 0.7% of the total zinc obtained by precipitation from the leach stream.

Examples of successful manganese removal experiments are shown below. As can be seen, the first two experiments achieved complete removal of manganese, while the third experiment did not attempt to completely remove manganese. It should be noted that indirect tests of the zinc content in the solution were only indicative as the amount of zinc precipitated out of solution was difficult to distinguish from the experimental errors associated with this test. The exact amount of precipitated zinc was determined by direct analysis of the residual residue.

The results of experiments 2 and 3, shown in Tables 2 and 3, reflect the way in which the said technique was used in a lead / zinc recovery process in which Halex was formed at the anode. The results indicate that for the 8% stream of the spent catholyte, the concentration of manganese kept in the circulation was 15 grams per liter. The co-precipitation of zinc together with the precipitation of MnO2 accounted for a loss of about 0.6% of the total amount of zinc obtained.

Examples of manganese removal

Non-limiting examples of manganese removal processes are described below.

Example 1

Electrolyte containing 200 grams per liter of NaCl, 110 grams of NaBr per liter, 50 grams of CaCl2 per liter, Zn²⁺ at 50 grams per liter and Mn²⁺ 15 grams per liter was obtained as a typical wash stream from a lead / zinc recovery process. 25 grams per liter of CaCO3 was added to the electrolyte in a reaction vessel at the start of the removal process. Ca (OCl) 2 was added directly to the electrolyte to oxidize 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.

Example 2

Electrolyte containing 50 grams per liter of NaCl, 110 grams of NaBr per liter, 50 grams of CaCl2 per liter, Zn²⁺ at 50 grams per liter and Mn²⁺ 15 grams per liter was obtained as a typical wash stream from a lead / zinc recovery process. 85 grams per liter of CaCO3 was added to the electrolyte in a reaction vessel at the start of the removal process. Halex vapors were obtained outside the reaction vessel and pumped into it. 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.

Example 3

Electrolyte containing 50 grams per liter of NaCl, 110 grams of NaBr per liter, 50 grams of CaCl2 per liter, Zn²⁺ at 50 grams per liter and Mn²⁺ 15 grams per liter was obtained as a typical wash stream from a lead / zinc recovery process. During the removal process, CaCO3 was added in small portions to the electrolyte in the reaction vessel. Halex vapors were obtained outside the reaction vessel and pumped into it. 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.

Example 4

10.0 g of the manganese precipitate from Example 2 was added to 1.0 liter of demineralized water in a reaction vessel. H 2 SO 4 was added. The results are shown in Table 4 below.

Example 5

10.0 g of the manganese precipitate from Example 3 was added to 1.0 liter of demineralized water in the reaction vessel. Halex vapors were obtained outside the reaction vessel and pumped into it. H2SO4 was added. The results are shown in Table 5 below.

TABLE 1 Results of the experiment 1

<td rowspan="2">Time (min)</td><td rowspan="2">PH</td><td rowspan="2">Eh (mV vs Ag / AgCl)</td><td rowspan="2">Mass of added Ca (OCl) 2 (g)</td><td colspan="2">Zn</td><td colspan="2">Me</td>

<td>Abs</td><td>% removed</td><td>Abs</td><td>% removed</td>

<td> 10</td><td> 4,7</td><td> 480</td><td> 0</td><td> 1018</td><td> 0%</td><td> 297</td><td> 0%</td>

<td> 25</td><td> 4,5</td><td> 470</td><td> 2,2</td><td> 1018</td><td> 0%</td><td> 283</td><td> 5%</td>

<td> 40</td><td> 4,4</td><td> 520</td><td> 4,1</td><td> 1015</td><td> 0%</td><td> 265</td><td> 11%</td>

<td> 55</td><td> 4,4</td><td> 570</td><td> 6,2</td><td> 1016</td><td> 0%</td><td> 257</td><td> 13%</td>

<td> 70</td><td> 4,3</td><td> 550</td><td> 7,9</td><td> 1011</td><td> 1%</td><td> 227</td><td> 24%</td>

<td> 85</td><td> 4,3</td><td> 570</td><td> 10,1</td><td> 1010</td><td> 1%</td><td> 206</td><td> 31%</td>

<td> 110</td><td> 4,3</td><td> 560</td><td> 12,2</td><td> 1010</td><td> 1%</td><td> 186</td><td> 37%</td>

<td> 120</td><td> 4,1</td><td> 610</td><td> 16,4</td><td> 1009</td><td> 1%</td><td> 147</td><td> 51%</td>

<td> 135</td><td> 3,9</td><td> 660</td><td> 25,9</td><td> 1013</td><td> 0%</td><td> 73</td><td> 75%</td>

<td> 160</td><td> 3,8</td><td> 690</td><td> 32,3</td><td> 996</td><td> 2%</td><td> 16</td><td> 95%</td>

<td> 180</td><td> 3,8</td><td> 790</td><td> 38,0</td><td> 994</td><td> 2%</td><td>was not found</td><td> 100%</td>

<td></td><td></td><td></td><td>Mass</td><td colspan="2">Zn content</td><td></td><td></td>

<td colspan="2">Residue</td><td></td><td>24.3 g</td><td colspan="2"> 5,1%</td><td></td><td></td>

<td colspan="4">Zn losses in% of the total amount of Zn obtained</td><td colspan="2"> 0,4%</td><td></td><td></td>

PL 201 418 B1

TABLE 2 Results of the experiment 2

<td rowspan="2">Time (min)</td><td rowspan="2">PH in the reactor</td><td rowspan="2">Eh (mV vs Ag / AgCl)</td><td rowspan="2">32% HCl added to the reactor (ml)</td><td colspan="2">Zn in solution (AAS Abs)</td><td colspan="2">Mn in solution (AAS Abs.)</td>

<td>Abs</td><td>% removed</td><td>Abs</td><td>% removed</td>

<td> 0</td><td> 2,0</td><td> 470</td><td> 0</td><td> 361</td><td> 0%</td><td> 821</td><td> 0%</td>

<td> 30</td><td> 4,6</td><td> 390</td><td> 3</td><td> 339</td><td> 6%</td><td> 791</td><td> 4%</td>

<td> 60</td><td> 4,4</td><td> 500</td><td> 10</td><td> 340</td><td> 6%</td><td> 726</td><td> 12%</td>

<td> 110</td><td> 4,3</td><td> 520</td><td> 20</td><td> 331</td><td> 8%</td><td> 661</td><td> 19%</td>

<td> 170</td><td> 4,3</td><td> 550</td><td> 36</td><td> 311</td><td> 14%</td><td> 460</td><td> 44%</td>

<td> 210</td><td> 4,2</td><td> 540</td><td> 46</td><td> 305</td><td> 16%</td><td> 336</td><td> 59%</td>

<td> 255</td><td> 4,3</td><td> 560</td><td> 62</td><td> 306</td><td> 15%</td><td> 173</td><td> 79%</td>

<td> 280</td><td> 3,7</td><td> 750</td><td> 72</td><td> 290</td><td> 20%</td><td> 66</td><td> 92%</td>

<td> 310</td><td> 3,6</td><td> 790</td><td> 72</td><td> 286</td><td> 21%</td><td> 5</td><td> 99%</td>

<td> 1320</td><td> 4,4</td><td> 770</td><td> 72</td><td> 291</td><td> 19%</td><td>was not found</td><td> 100%</td>

<td></td><td></td><td></td><td>Mass</td><td colspan="2">Zn content</td><td></td><td></td>

<td colspan="2">Residue</td><td></td><td>53.6g</td><td></td><td> 16,5%</td><td></td><td></td>

<td colspan="5">Zn losses in% of the total amount of Zn obtained</td><td> 1,5%</td><td></td><td></td>

TABLE 3 Results of the experiment 3

<td rowspan="2">Time (min)</td><td rowspan="2">PH</td><td rowspan="2">Eh (mV)</td><td rowspan="2">HCl added to the Halex tank (ml)</td><td rowspan="2">CaCO3 added to the reaction vessel (g)</td><td colspan="2">Zn in solution (AAS Abs)</td><td colspan="2">Mn in solution (AAS Abs.)</td>

<td>(PPm)</td><td>% removed</td><td>(PPm)</td><td>% removed</td>

<td> 0</td><td> 1,0</td><td> 520</td><td> 0</td><td> 0,0</td><td> 51,1</td><td> 0%</td><td> 16,3</td><td> 0%</td>

<td> 25</td><td> 3,8</td><td> 600</td><td> 6</td><td> 1,3</td><td> 52,0</td><td> -2%</td><td> 15,9</td><td> 2%</td>

<td> 45</td><td> 3,2</td><td> 710</td><td> 12</td><td> 1,3</td><td> 50,4</td><td> 1%</td><td> 16,0</td><td> 2%</td>

<td> 55</td><td> 2,9</td><td> 750</td><td> 15</td><td> 1,3</td><td> 51,9</td><td> -2%</td><td> 16,4</td><td> -1%</td>

<td> 70</td><td> 2,6</td><td> 780</td><td> 15</td><td> 1,3</td><td> 51,3</td><td> 0%</td><td> 15,7</td><td> 4%</td>

<td> 90</td><td> 3,2</td><td> 720</td><td> 18</td><td> 2,7</td><td> 50,5</td><td> 1%</td><td> 15,8</td><td> 3%</td>

<td> 135</td><td> 3,3</td><td> 700</td><td> 24</td><td> 8,0</td><td> 51,4</td><td> -1%</td><td> 14,3</td><td> 12%</td>

<td> 200</td><td> 3,2</td><td> 700</td><td> 36</td><td> 16</td><td> 52,4</td><td> -3%</td><td> 11,6</td><td> 29%</td>

<td> 260</td><td> 3,3</td><td> 760</td><td> 50</td><td> 25</td><td> 51,6</td><td> -1%</td><td> 9,0</td><td> 45%</td>

<td> 290</td><td> 3,3</td><td> 780</td><td> 57</td><td> 33</td><td> 48,7</td><td> 5%</td><td> 6,5</td><td> 60%</td>

<td> 330</td><td> 3,3</td><td> 780</td><td> 68</td><td> 42</td><td> 47,8</td><td> 6%</td><td> 4,1</td><td> 75%</td>

<td> 1450</td><td> -</td><td> -</td><td> 68</td><td> 42</td><td> 47,7</td><td> 7%</td><td> 3,8</td><td> 77%</td>

<td></td><td></td><td></td><td colspan="2">Mass</td><td colspan="2">Zn content</td><td></td><td></td>

<td colspan="3">Residue</td><td colspan="2">43.2 g</td><td colspan="2"> 9,6%</td><td></td><td></td>

<td colspan="5">Zn losses in% of the total amount of Zn obtained</td><td colspan="2"> 0,62%</td><td></td><td></td>

PL 201 418 B1

TABLE 4 Experiment results 4

<td>Time (min)</td><td>PH</td><td>Eh (mV vs Ag / AgCl)</td><td>Added 9.8% H2SO4 (ml)</td><td>Zn in solution (PPm)</td><td>Zn Dissolved (% of the Initial amount)</td><td>Mn in solution (PPm)</td><td>Dissolved Mn (% of the initial amount)</td>

<td> 0</td><td> 3,5</td><td> 390</td><td> 0</td><td> -</td><td> -</td><td> -</td><td> -</td>

<td> 15</td><td> 6,6</td><td> 230</td><td> 1</td><td> &lt;10</td><td> &lt; 0,6%</td><td> &lt;10</td><td> &lt; 0,4%</td>

<td> 20</td><td> 6,0</td><td> 280</td><td> 3</td><td> &lt;10</td><td> &lt; 0,6%</td><td> &lt;10</td><td> &lt; 0,4%</td>

<td> 30</td><td> 5,7</td><td> 390</td><td> 6</td><td> &lt;10</td><td> &lt; 0,6%</td><td> &lt;10</td><td> &lt; 0,4%</td>

<td> 35</td><td> 5,4</td><td> 440</td><td> 9</td><td> 41</td><td> 2,5%</td><td> &lt;10</td><td> &lt; 0,4%</td>

<td> 45</td><td> 5,1</td><td> 450</td><td> 15</td><td> 97</td><td> 5,9%</td><td> 13</td><td> 0,5%</td>

<td> 60</td><td> 4,3</td><td> 530</td><td> 21</td><td> 228</td><td> 14%</td><td> 51</td><td> 1,8%</td>

<td> 70</td><td> 3,5</td><td> 630</td><td> 27</td><td> 529</td><td> 32%</td><td> 214</td><td> 7,7%</td>

<td> 80</td><td> 2,9</td><td> 710</td><td> 33</td><td> 808</td><td> 49%</td><td> 410</td><td> 15%</td>

<td> 180</td><td> 2,1</td><td> 830</td><td> 40</td><td> 1020</td><td> 62%</td><td> 562</td><td> 20%</td>

<td></td><td></td><td></td><td>Mass</td><td>Zn (%)</td><td>Zn dissolved</td><td></td><td></td>

<td colspan="3">Residue</td><td>4.0 g</td><td> 7,9%</td><td> 82%</td><td></td><td></td>

<td colspan="6">Adjusted Zn loss in% of total Zn obtained</td><td> 0,26%</td><td></td>

TABLE 5 Experiment results 5

<td>Time (min)</td><td>PH</td><td>Eh (mV vs Ag / AgCl)</td><td>added 9.8% H2SO4 (ml)</td><td>Zn in solution (PPm)</td><td>Zn Dissolved (% of the Initial amount)</td><td>Me in solution (PPm)</td><td>Dissolved Mn (% of the initial amount)</td>

<td> 0</td><td> 5,0</td><td> 295</td><td> 0</td><td> 336</td><td> -</td><td> 55,2</td><td> -</td>

<td> 15</td><td> 3,4</td><td> 810</td><td> 4,0</td><td> 578</td><td> &lt; 0,6%</td><td> 77,7</td><td> &lt; 0,4%</td>

<td> 25</td><td> 2,9</td><td> 860</td><td> 4,0</td><td> 607</td><td> &lt; 0,6%</td><td> 80,4</td><td> &lt; 0,4%</td>

<td> 35</td><td> 2,3</td><td> 900</td><td> 6,0</td><td> 646</td><td> &lt; 0,6%</td><td> 88,7</td><td> &lt; 0,4%</td>

<td> 45</td><td> 1,9</td><td> 910</td><td> 10,0</td><td> 696</td><td> 2,5%</td><td> 88,3</td><td> &lt; 0,4%</td>

<td> 55</td><td> 1,6</td><td> 900</td><td> 15,0</td><td> 587</td><td> 5,9%</td><td> 91,0</td><td> 0,5%</td>

<td> 60</td><td> 1,4</td><td> 920</td><td> 25,0</td><td> 460</td><td> 14%</td><td> 95,7</td><td> 1,8%</td>

<td> 75</td><td> 1,4</td><td> 920</td><td> 25,0</td><td> 508</td><td> 32%</td><td> 109</td><td> 7,7%</td>

<td></td><td></td><td></td><td>Mass</td><td>Zn (%)</td><td>Zn dissolved</td><td></td><td></td>

<td colspan="3">Residue</td><td>4.1 g</td><td> 3,6%</td><td> 83%</td><td></td><td></td>

<td colspan="6">Adjusted Zn loss in% of total Zn obtained</td><td> 0,10%</td><td></td>

Thus, the manganese in the form of manganese dioxide was completely removed from the leach stream by interaction with Halex vapors. It has surprisingly been found that the electrolyte-vapor reaction occurs very quickly in the presence of small amounts of limestone.

A small amount of zinc was also precipitated during the removal of manganese. However, less zinc precipitated at the lower level of zinc in a water-soluble form (e.g. with Zn²⁺ 25 grams per liter instead of 50 grams per liter), or when limestone was added in small, gradually increasing amounts to maintain the pH around 3.2 (as opposed to manganese precipitation in the presence of a stoichiometric excess of limestone).

PL 201 418 B1

Thus, when using a spent Zn-containing catholyte²⁺ at 50 grams per liter and Mn²⁺ at 15 grams per liter, and with the slow addition of limestone, only 8.4% of the zinc in the solution was co-precipitated with the manganese. For the 8% elution catholyte stream, this equates to a loss of less than 0.7% of the total amount of zinc obtained by precipitation from the elution stream.

In the magnesium removal process, slaked lime was added to the spent catholyte from the manganese removal process in a series of CSTR reactors to cause zinc precipitation in the first two CSTR reactors. This zinc was recycled to the displacement process, typically in the limestone addition stage (as shown in Figure 1). More quicklime was then added in two successive CSTR reactors to cause the precipitation of magnesium oxide and thus prevent magnesium build-up in the process. This removal step was only required intermittently due to the generally lower levels of magnesium in the zinc mineral. The processed spent catholyte was then recycled to the leaching process, typically the first leach CSTR.

It has been found that the cost effectiveness of the preferred zinc recovery process in which Halex was formed at the anode is site specific. For the concentrate producer, the loss of 0.6% of the obtained zinc due to the removal of impurities (in particular manganese) meant a significant improvement over the losses associated with metal smelting processes known in the art, in which about 4% of the zinc contained was removed before calculating the metal content, which can be sold. However, when a higher degree of zinc recovery in the process was desired, some of the zinc precipitated from the manganese solution was recovered by resuspending the precipitate of the water washout stream and acidifying with sulfuric acid (VI). The zinc could be redissolved with a selectivity of about 90%. With the admission of a larger leach stream ratio, this required taking into account the partial redissolution of manganese, and the zinc losses related to the precipitate from the leach stream were reduced to 0.1% of the total amount of zinc produced.

An important feature of the preferred process was that all impurities including manganese, mercury and arsenic were recovered as by-products in a salable form or in a stable form that could be disposed of as waste. Another equally important feature was that heat was supplied by exothermic reactions in the leaching process. This, in combination with the aeration of the leach solution, caused the water to evaporate and thus kept the water equilibrium at a neutral level. Therefore, there was no sewage (liquid waste) in the process.

Removal of silver and mercury

Preferred embodiments of the silver and mercury removal process for copper (I) chloride leachate will be described below, but it should be understood that the invention can be used with chloride solutions of other metals (including zinc chloride).

In electrowinning, the typically observed Eh value of a copper (I) containing solution in contact with copper metal was about +150 mV (versus Ag / AgCl). It has surprisingly been found that it is possible to chemically lower the Eh value of a solution containing copper (I) ions to -150 mV (relative to Ag / AgCl) with minimal precipitation of metallic copper. It allowed to obtain an additional 300mV driving force of the silver precipitation process. Moreover, the efficiency of silver mercury amalgam formation was not significantly reduced by alloying with copper. Some of the copper was observed to co-precipitate with the silver mercury amalgam, but the amalgam (residual) copper was found to be 1/20 of the amount of residue from prior art processes, greatly simplifying the recovery of silver and mercury.

After removing the silver, it was found that it was possible to remove trace amounts of mercury from the solution containing copper (I) ions by scattering nitrogen therein. After dissipation, the nitrogen was washed, for example with a solution containing copper (II) ions, to remove bubble-entrained mercury gas therefrom. The nitrogen was then returned to the solution to remove mercury further.

Examples of silver / mercury removal.

Example 6

An improved method of removing silver and mercury from chloride solutions has been developed using an improved reducer, i.e. metal aluminum, which has proved to be a cheap and ready-to-use reducer compared to those described below for zinc dust or metal iron.

In addition, the use of metallic aluminum caused the formation of a spongy copper sludge in copper chloride solutions, the sludge after silver and / or mercury cementation was

Easier to filter than the precipitate formed with the use of other reducing agents. Moreover, the aluminum was bound in this precipitate while the zinc dust did not, so that the aluminum could then be removed from the solution together with the precipitate. A thin layer of copper had formed on the metallic iron and its core part did not react, making it a less effective reducing agent.

In an improved method of removing silver and mercury from chloride solutions, metal aluminum was added to the electrolyte in small amounts (0.05 - 0.3 grams per liter of electrolyte), reducing the Eh value in the solution to -120: 150 mV (against Ag / AgCl ) and precipitation of copper with a very developed surface. This copper was observed to form a relatively coarse sponge-like precipitate. It was very reactive with the added mercury ions and was also able to form an amalgam with a chemical composition that was about 7.5% silver, copper in an amount ranging from 10% to 50%, and an equivalent amount of mercury.

The amalgam formed by adding aluminum also had a more uniform particle size and was easier to drain off quickly. This was an improvement over the prior art processes for the separation of solid and liquid phases.

The solution mercury levels at the end of this step were similar to those described in Example 8, and the same mercury removal steps (nitrogen scattering or activated carbon adsorption) were also used as described below.

The improved process was tested on a laboratory scale, prior to a process check, during a 48-hour pilot-scale process, with a continuous process solution flow rate of 20 L / hr. The experimental plant consisted of four tanks, each with a capacity of 20 liters, in which nitrogen was dissipated, with aluminum foil being added to tank 1 at a dosing rate of 0.05 to 0.15 grams per liter of electrolyte (1 to 3 g / hr). and an ionic mercury compound (e.g. mercury (I) or mercury (II) nitrate) added to tank 1 in a weight ratio of 3: 1 relative to the amount of silver supplied (1.5 g / h). The processed electrolyte at 35 ° C contained Cu⁺ at 75 grams per liter, NaCl at 280 grams per liter and NaBr at 28 grams per liter. Fig. 3 shows the averaged results for the silver and mercury contents and for the Eh values ​​over a 48-hour period of operation.

As can be seen in Figure 3, the average silver concentration in tanks 3 and 4 over the 48 hours of operation was 1.5 mg / liter with a minimum concentration of 1.2 mg / liter. The mercury concentration ranges observed were as shown in Table 6.

Table 6

The range of mercury concentration values ​​for the 48-hour period of the process on a semi-technical scale.

<td>Tank No.</td><td>number of tries</td><td>Mean concentration ^ g / l)</td><td>Concentration range ^ g / l)</td>

<td> 1</td><td> 6</td><td> 6311</td><td> 815-24700</td>

<td> 2</td><td> 6</td><td> 1293</td><td> 337-4300</td>

<td> 3</td><td> 6</td><td> 468</td><td> 275 - 773</td>

<td> 4</td><td> 6</td><td> 322</td><td> 277 - 388</td>

Example 7

The reducer in this example was sodium borohydride, but it was observed that other reducing agents such as sodium (tetraoxo) disulfate and metal hydrides (e.g. sodium and calcium hydride, optionally in combination with sodium borohydride, could be used to reduce the amount of borohydride). The most economical method was to use sodium (tetraoxo) disulfate to significantly lower the Eh value (i.e. typically down to a limiting level), with a final reduction in Eh being done with sodium borohydride (i.e. to achieve a preferred lower level of about -150 mV (versus Ag / AgCl)).

Contrary to the use of metallic aluminum as a reducing agent, sodium borohydride caused the precipitation of very fine copper grains forming a spongy structure, and upon addition of an ionic mercury compound, the formation of an amalgam that was difficult and expensive to filter and also less durable. Also, the borohydride itself is both expensive and unstable in water. If not stored under high pH conditions, the borohydrides rapidly hydrolyze, wasting an expensive reagent. Likewise, the very fine amalgam formed after the addition of mercury and sodium borohydride had to be completely prevented from oxidation, otherwise it would re-dissolve rapidly, releasing the silver and mercury ions back into solution.

PL 201 418 B1

Despite these difficulties, it was also possible to use sodium borohydride as a reducing agent, with the addition of an ionic mercury compound (Hg (I) or Hg (II)), a metallic amalgam of copper, mercury and silver was formed, removing both mercury and silver from the solution. and silver.

Silver removal

45 mg of NaBH4 was added to 1.5 liters of electrolyte containing reduced copper and 22 ppm silver, causing the Eh value of the solution to drop from + 120 mV to -135 mV (against Ag / AgCl). 65 mg of mercury in the form of mercury (II) nitrate was added to the electrolyte and the silver content of the solution was reduced from 22 ppm to 4.4 ppm in 40 minutes.

Another 65 mg of mercury was added in the form of mercury (II) nitrate and the silver content had dropped to 1 ppm within 20 minutes. The precipitated solids were filtered and dried and found to contain 7.4% silver, 10% to 47% copper, and an equivalent amount of mercury.

The process was repeated using sodium (tetraoxo) bisulfate to achieve an Eh level below 0 mV, with a final Eh reduction being achieved with smaller amounts of NaBH4.

Example 8

Mercury removal

Eh value of the electrolyte containing Cu ions⁺which electrolyte was derived from the silver removal process of Example 7, was adjusted to -150 mV by adding a reducing agent such as aluminum or NaBH4 thereto. This shifted the normal equilibrium of the solution (i.e. between metallic mercury, Hg ions⁺ and Hg ions²⁺) towards the formation of metallic mercury. Thereafter, nitrogen was flushed into a subjected electrolyte for 1 hour. It was observed that during this time the mercury content of the electrolyte dropped from 0.29 ppm to 0.037 ppm. It was also observed that mercury removal was increased by heating the solution during nitrogen scattering.

In the second study, the initial mercury concentration of 1600 μg / L fell to 150 μg / L within 60 minutes and to 100 μg / L in 120 minutes. The results are shown in Fig. 4.

Example 9

Alternative way to remove mercury

Eh value of the electrolyte containing Cu ions⁺which electrolyte was derived from the silver removal process of example 2 and containing mercury in an amount of 500 µg / l, was adjusted to about -150 mV (versus Ag / AgCl). This solution was then diverted (e.g., pumped under pressure) for through the activated carbon column at 4 ml / min (for the first three hours) and then at 20 ml / min (for the next two hours). Metallic mercury was adsorbed on activated carbon, and its concentration in the solution dropped below 55 µg / l. The metallic mercury was then separated from the activated carbon by passing a second, different, high Eh solution through the column, thereby converting the metallic mercury into a soluble mercury ionic compound to elute from the column along with this second solution. The mercury was then recovered from the second solution.

The experimental procedures described above showed that it was possible to achieve a very effective removal of both silver and mercury without the need for copper (as was the case in the prior art). Moreover, the removal processes were easy to carry out. The simplicity of this process allowed to immediately increase the scale of implementation to an industrial scale and to use it as a continuous process.

The zinc recovery process of the present invention has the following easily identifiable advantages over the prior art roasting / leaching / electrowinning zinc and other hydrometallurgical processes:

- significantly lower operating costs and capital costs,

- recovery of valuable metals in the zinc leaching circuit (e.g. by purifying the solution from metals other than zinc, recovering commercially available copper, lead and silver (Cu, Ag, Pb and Fe in the cementation and precipitation stage in Figure 1)),

- the possibility of using poorly cleaned and unclean concentrates,

- low energy consumption,

- no liquid waste,

- no production of harmful gases,

- gentle working conditions - low temperature and atmospheric pressure,

- no need for solvent extraction, and

- no requirement to use pure oxygen in the aeration stages.

PL 201 418 B1

While the invention has been described with reference to a number of preferred embodiments, it should be recognized that the invention is capable of being embodied in many other forms.