FIELD OF THE INVENTION

The present invention belongs to the field of metal recovery and concerns the production of high purity metal products. It relates to a hydrometallurgical process for recovering metals by means of ammonia leaching and solvent extraction, whereby entrained ammonia is recovered from the organic solvent by means of gas desorption with a sorption gas followed by gas absorption, allowing the recovered ammonia to be recycled back to leaching or to be used otherwise. The invention comprises the selective extraction of one or more target metals such as copper, zinc, nickel, cobalt, silver, gold, platinum, palladium, rhodium, mercury, chromium, cadmium, molybdenum, and rhenium, among others, from a metalliferous feed material, such as bottom ash, metal scrap, waste, ore, concentrates, tailings, or slags, to produce one or more high purity metal products such as metal cathodes, metal salts, metal solutions, metal powder, metal pulps, or other metal compounds. The sorption gas used to desorb the entrained ammonia from the organic solvent preferably contains carbon dioxide, e.g. a mixture of carbon dioxide and air or a combustion gas derived from the combustion of an organic fuel, and preferably has low oxygen gas content, e.g. a mixture of carbon dioxide and nitrogen gas.

Some embodiments of the invention involve in particular the extraction, separation, and purification of two or more metals from secondary waste materials that originate from or comprise bottom ash (e.g. bottom ash deriving from municipal waste incinerators), metallic shredder material (e.g. from scrapped cars), or heavy metal fractions originating from WEEE material (waste electrical and electronic equipment).

BACKGROUND OF THE INVENTION

The main goal of a hydrometallurgical process, and of extractive metallurgy in general, is to achieve high metal recovery at low capital and operational costs. One of the major costs typically involved in a hydrometallurgical process that involves leaching is the consumption of leaching agents per ton of processed metalliferous feed material and per kilogram of final metal product. If the consumption of leaching agents is too high, then the overall process may become economically unattractive. Leaching is a well-known process that involves the dissolution of desired target metals into an aqueous phase, by contacting an aqueous solution that contains leaching agents, called leach solution or lixiviant, with the metalliferous feed material, thereby obtaining a pregnant leach solution (PLS) with dissolved target metals. Common types of leaching techniques are heap leaching, agitation leaching, in-situ leaching, dump leaching, vat leaching, and pressure leaching. Ammonia leaching, in particular, is basic in nature and considers the use of ammonia and/or ammonium salts (e.g. ammonium sulfate, ammonium carbonate, among others) as leaching agents, having the advantage of being more selective towards target metals such as copper, zinc, nickel, and cobalt, among others, and of being less selective towards possible impurities such as iron, aluminum, magnesium, among others. However, the ammoniacal leaching agents are rather more expensive and volatile, particularly at higher concentrations, when compared with other commonly used acidic leaching agents such as sulfuric acid or hydrochloric acid. For the economic viability of the overall process it becomes therefore necessary to minimize the consumption of leaching agents, particularly when dealing with ammonia and related compounds.

An example of ammonia leaching can be found in U.S. patent 4,165,264 "Ammonia leaching" by Satchell, which proposes an improved process for obtaining copper from copper sulfide by leaching with an ammonium carbonate solution, oxygen and recycled gaseous ammonia and carbon dioxide. The proposed process requires the addition of oxygen to oxidize copper sulfide during leaching, the presence of several filtering steps, generating heat to form gaseous ammonia and carbon dioxide from the entire pregnant leaching solution, the addition of a strongly alkaline material like lime to precipitate sulfates in several parts of the process, and the elimination of ammonia before the electrolytic recovery of copper in an acid medium. It is rather complex and does not consider solvent extraction.

Solvent extraction (SX), sometimes also called liquid-liquid extraction or liquid ion exchange, is a well-known process that involves contacting a solvent extraction feed solution, i.e. the pregnant leach solution from the leaching step, with a water immiscible organic solution, called solvent, comprising one or more metal extracting agents or extractants to selectively extract the metal ions of the target metals of interest from the pregnant leach solution into the organic solution, thereby obtaining a loaded organic (solution). The loaded organic, after its separation from the aqueous solution depleted of target metals (called raffinate), is contacted with another aqueous solution (called electrolyte or stripping solution) to strip the target metals from the organic phase into the aqueous phase, from where they are then recovered by some metal recovery method (e.g. electrowinning, precipitation, etc.). Electrowinning (EW) involves the recovery of metals from the aqueous solution by means of an electrolytic cell, wherein an electrodeposition of metals occurs on the cathode(s), and either an oxidation reaction or a metal dissolution on the anode(s). Precipitation involves generating a solid precipitate from the aqueous solution either by cementation, whereby ions are reduced to zero valence with a reducing agent, or by crystallization, whereby the solubility conditions of dissolved metals or contaminants are changed, e.g. by reagent addition, temperature change, or evaporation.

Ammoniacal leaching agents may be consumed by many factors in the process, including the dissolution of target metals and other undesired impurities, by volatilization, by impregnation in the leached waste material, by entrainment in the organic solvent, among others. Of particular concern is the loss of ammonia by entrainment, which may be physical (e.g. aqueous droplets entrained in the organic solvent) or chemical (e.g. metal ammine complexes or ammonium cations loaded onto the organic solvent). Such entrainments are specially undesired when the stripping is performed in acid media, as is often the case. One way to deal with this issue is to include one or more wash stages in the solvent extraction circuit, whereby the loaded organic is contacted with corresponding aqueous wash solutions prior to stripping. This works only as an option if the water balance does accommodate the additional water required to run the wash stages, e.g. as impregnation in the leached waste material or in other process steps.

One alternative or complementary option thereto is the process disclosed in U.S. patent 5,788,844 "Process for removing and recovering ammonia from organic solutions in liquid-liquid process" by Olafson, whereby ammonia is removed from the organic phase by sparging it with an inert gas, preferably air. The concept of sparging is defined as the agitation of a liquid by means of an air or gas entering the liquid, preferably by means of a pipe or other conduit thereby flowing or bubbling through the liquid. The process is typically carried out at ambient temperature, although it is mentioned that the use of higher temperatures, which in some cases may prove to be beneficial, is also contemplated, without giving any detailed description or example. The patent does not mention any other inert gas, besides air, and does not specifically contemplate operating at a pressure that is lower than the atmospheric pressure. Furthermore, once removed from organic phase, the patent does not contemplate any procedure to recover the ammonia from the ammonia-bearing sparged gas in any useful way so as to be recycled or integrated again into the process, e.g. in such a way so as to reuse the ammonia in the leaching step.

When the final metal product corresponds to a solid metal salt (e.g. copper sulfate, zinc sulfate, nickel sulfate, or cobalt sulfate), then its production typically involves the precipitation of crystals from a solution, i.e. crystallization. Such crystallization occurs in two steps, nucleation and crystal growth. Nucleation is the step where some solute atoms or molecules dispersed in a solvent start to gather so as to reach a critical size and form stable clusters, constituting so-called nuclei or crystal seeds on the microscopic scale, characterized by being the atoms or molecules arranged in a periodic manner that defines the crystal structure. Crystal growth is the subsequent step whereby the nuclei increase in size under equilibrium conditions, in such a manner that solute atoms or molecules precipitate out of solution. Depending upon the conditions, either nucleation or growth may be predominant over the other, dictating crystal size. Crystal formation can be achieved by various methods, such as cooling, evaporation, change of pH, change of cations or anions, addition of a second solvent to reduce the solubility of the solute (e.g. ethanol in sulfate media), among several other methods. To induce crystallization, an oversaturated solution must be generated and seed crystals may be necessary for triggering the crystallization.

On this regard, from the prior art U.S. patent 4,561 ,887 "Process for recovering metals from solution through solvent extraction and pyrometallurgical reduction" by Domic & Hein is known, which discloses a process for recovering metals (particularly copper) from solution through solvent extraction and pyrometallurgical reduction. The invention achieves recovering the metal of interest from the organic reagent by means of a saturated aqueous stripping solution in such a way that all the recovered metal precipitates as a metal salt (particularly as copper sulfate) in view of oversaturation. The precipitation of solid crystals is achieved even at ambient temperature and pressure, preferably in the mixer section of a mixer settler unit, wherein the crystals remain temporarily suspended because of the vigorous agitation used for contacting the aqueous and organic phases, and because the crystal particles are small at first. The metal salt crystals grow in size in the aqueous/organic emulsion and, once their size overcomes a certain threshold, the metal salt crystals leave the emulsion by gravity, settling at the bottom of the settler. The precipitation of the metal salt crystals occurs largely at the interface between the organic and aqueous phases, where additional metal ions transfer from the organic liquid to the stripping solution and where the concentration of metal ions is highest. In absence of previously formed crystals or other sites for crystal growth, the solution can become supersaturated until nucleation occurs and new crystals grow, establishing a steady state of crystal formation and settling. Whenever the crystals leave the emulsion, some organic solution is entrained in the settled metal salt slurry. Usually organic entrainments are reduced by operating the mixer settler unit in organic continuity. The metal salt crystals are separated from the stripping solution by removing the solution from the slurry of crystals collected from the bottom of the settler. The metal salt is further separated from liquid on a filter or centrifuge to obtain a reasonable dry filter cake. The separated crystals are then washed or rinsed with water in a wash stage, displacing stripping solution, redissolving precipitated impurities and removing traces of organic liquid. The metal salt crystals are thereafter dried, either by mild heating maintained below about 25°C or even as high as above 250°C for complete dehydration (of copper sulfate crystals). The resulting dried metal salt crystals are considered to be an intermediate product in the process of the invention, and further subjected to a pyrometallurgical reduction to produce a pure metal product (e.g. copper).

U.S. patent 6,869,520 "Process for the continuous production of high purity electrolytic zinc or zinc compounds from zinc primary or secondary raw materials" by Martin, Diaz & Garcia proposes a process to produce high purity electrolytic zinc or zinc compounds, wherein the organic entrainment may be removed with activate charcoal to minimize organic losses, from the final solutions originating from the solvent extraction system, i.e. from the raffinate solution originating from extraction, from the acid solution originating from stripping, and/or from the spent hydrochloric acid solution originating from regeneration. This patent considers stripping the metals into an electrolyte, which is taken outside of the mixer settler unit for crystallization, i.e. it does not consider a simultaneous stripping and crystallization.

Examples of other patents where metal salts are crystallized outside of the stripping (mixer settler) unit are U.S. patent 4,563,256 "Solvent extraction process for recovery of zinc" by Sudderth, Sierakosky & Lewis, U.S. patent 4,314,976 "Purification of nickel sulfate" by Stewart, Odle & Brunson, and U.S. patent 4,401 ,531 "Process for the production of electrolytic zinc or high purity zinc salts from secondary zinc raw-materials" by Martin, Regife & Nogueira.

U.S. patent 8,323,596 "Method for extracting zinc from aqueous ammoniacal solutions" by Johnston, Sutcliffe & Welham discloses a process for the extraction of zinc from an ammoniacal zinc solution by solvent extraction and its recovery as a zinc (end) product. In the patent, the zinc product is obtained from an aqueous zinc solution in a separate recovery stage besides the stripping stage, i.e. not in a simultaneous manner.

U.S. patent 3,966,569 "Method of recovering metal from metalliferous waste" by Reinhardt & Ottertun divulges a process for the recovery and separation of zinc and iron (and eventually also copper, nickel, and chromium) from metal-containing waste by means of a solvent (liquid-liquid) extraction process that considers two separate stripping stages (called washing operations), wherein at least one of them considers very high free sulfuric acid concentration (500 to 1000 g/L H2SO4), where a crystallization by cooling is considered in a separate stage, i.e. not together with stripping.

U.S. patent 3,929,598 "Recovery of copper and zinc from low-grade non-ferrous materials" by Stern, Jansen & Vance describes a process for the recovery of copper and zinc from low-grade non-ferrous materials by ammoniacal leaching and solvent extraction (liquid ion exchange) to produce electrolytic copper and zinc carbonate precipitate. Alternatively, the copper could be precipitated by sulfur dioxide to copper sulfite and then decomposed to copper metal and the zinc could also be obtained by electrowinning. The copper is extracted in a first solvent extraction circuit and the zinc is then extracted either by precipitation from the resulting raffinate solution or by a second solvent extraction circuit. The crystallization is not considered together with the stripping stages.

Other examples of patents that disclose processes for the solvent extraction, separation, and recovery of two different metals, in particular copper and zinc, are U.S. patent 3,853,981 "Liquid ion exchange process for the recovery of metals" by Hadzeriga, U.S. patent 3,989,607 "Solvent extraction and electrowinning of zinc and copper from sulfate solution" by Bush & Bailey, and U.S. patent 4,362,607 "Recovery of copper and zinc from complex chloride solutions" by Ritcey, Price & Lucas.

The particular case of recovering metals from secondary streams such as metal-comprising waste streams is considered more challenging as these streams comprise a mixture of metals. Recovery methods typically involve leaching a desired metal (or several desired metals) from the secondary stream. In case one or more other metals are present that are leached more easily, the presence of these metals interferes with the leaching of the metal(s) of interest. For instance, by treating a stream comprising a mixture of zinc and copper present as an alloy (e.g. brass), it is hard to selectively leach copper and zinc. Additionally, the presence of multiple metals also complicates the (selective) recovery of desired metal or metals after the leaching stage. The recovery and separation of such metals can be achieved by combined or separate solvent extraction circuits after the leaching operation. In a combined solvent extraction circuit the separation can be achieved by separate stripping stages. Some examples of patents dealing with metal recovery from secondary streams are the already mentioned U.S. patent 6,869,520, U.S. patent 3,929,598, and U.S. patent 3,966,569.

As can be appreciated from the previous patents, no mention was found in the prior art of a process and apparatus for the recovery of metals by means of ammonia leaching and solvent extraction using gas desorption and absorption with a sorption gas preferably containing carbon dioxide and preferably low in oxygen gas content, so as to reduce the consumption of ammoniacal leaching agents by recycling them back to leaching, allow the makeup of anions (particularly of carbonate ions) in the leach solution, avoid any potential oxidation (and hence degradation) of the organic extractant in solvent extraction, adjust the pH in the leach solution as well as in the absorption solution so as to improve leaching efficiency and the absorption of ammonia in the absorption solution, and act as an alternative to a wash step for the loaded organic solution when the water balance does not accommodate the additional amount of water that would be required to be disposed of in this case.

One of the objectives of the present invention is to overcome the drawbacks associated with the prior art, or to at least provide a useful alternative thereto.

SUMMARY OF THE INVENTION

The present invention discloses a hydrometallurgical process for the recovery of one or more target metals from a metalliferous feed material to produce one or more metal products by using ammonia leaching and solvent extraction, whereby entrained ammonia is recovered from the organic solvent by means of gas desorption (also called gas stripping) with a sorption gas, preferably containing carbon dioxide and preferably low in oxygen gas content, followed by gas absorption (also called gas scrubbing), allowing the recovered ammonia to be recycled back to leaching, e.g. to the leach solution or the pregnant leach solution.

Preferably said metalliferous feed material is chosen from the group consisting of bottom ash, metal scrap, water, ore, concentrates, tailings, and slags.

In one embodiment of the invention, the process comprises the steps of: (a) leaching the feed material with an ammoniacal leach solution to produce a resultant pregnant leach solution containing dissolved metal ions of the target metal(s) and a resultant leached waste material with reduced content of target metal(s); (b) contacting the pregnant leach solution with a water immiscible organic solution comprising one or more metal extracting agents or extractants to selectively extract the metal ions from the pregnant leach solution into the organic solution, thereby providing a resultant loaded organic solution containing target metal(s) and a resultant raffinate solution with reduced content of target metal(s); (c) desorbing entrained ammonia from the loaded organic solution with a sorption gas, thereby providing a resultant enriched gas stream containing gaseous ammonia and a resultant desorbed loaded organic solution with reduced content of entrained ammonia; (d) absorbing the gaseous ammonia from the enriched gas stream in an aqueous absorbent solution, thereby providing a resultant aqueous absorption solution with increased content of dissolved ammonia that is recycled back to the leaching step (a), to the extraction step (b), or to both, and a resultant exhaust gas stream that is barren of ammonia; (e) contacting the desorbed loaded organic solution with an acidic stripping solution to strip the metal ions from the organic solution, thereby providing a resultant pregnant electrolyte stream with increased content of target metal ions and a resultant stripped organic solution with reduced content of metal ions; and (f) recovering target metal(s) from the pregnant electrolyte stream in the form of one or more final metal products.

In other embodiments of the invention, the gas desorption step (c) and the gas absorption step (d) involve contacting the solution flow with the gas stream in a countercurrent manner. In yet other embodiments, the gas desorption step (c) is operated at a higher temperature (above ambient temperature), a lower pressure (below atmospheric pressure), or both, whereas the gas absorption step (d) is operated at a lower temperature (below ambient temperature), a higher pressure (above atmospheric pressure), or both.

In yet another embodiment of the invention, the exhaust gas stream from the gas absorption step (d), either in whole or in part, is recirculated back to the gas desorption step (c) as part of the sorption gas.

In another embodiment of the invention, the process further comprises a steam stripping step, whereby the absorption solution from the gas absorption step (d) is contacted with steam, thereby providing a resultant enriched steam with increased content of ammonia and a resultant depleted solution with reduced content of ammonia, which can be returned back to the gas absorption step (d) as the absorption solution or sent to other process steps. In this embodiment, the process further comprises a condensation step, whereby the enriched steam from the steam stripping step is changed from gas phase to liquid phase, by cooling, compression, or both, thereby providing a condensed enriched solution with higher ammonia concentration than the absorption solution, which is returned to the leaching step (a), to the extraction step (b), or to both.

In yet another embodiment of the invention, the process further comprises the step of contacting the desorbed loaded organic from the desorption step (c) with an acidic, neutral or basic (loaded organic) wash affluent solution so as to remove further impurities, leaching agents, or both from the loaded organic, thereby providing a resultant (loaded organic) wash effluent solution with increased content of impurities and a resultant washed loaded organic with reduced content of impurities, leaching agents, or both, prior to sending the (washed) loaded organic to the stripping step (e).

In even another embodiment of the invention, aimed in particular to produce high purity metal salt crystals, the stripping step (e) involves an acidic stripping solution that is saturated with dissolved metal ions of the target metal(s), so as to strip the metal ions from the loaded organic solution and simultaneously precipitate salt crystals of the target metal(s) in the stripping solution, thereby providing a resultant stripped organic solution with reduced content of metal ions, a resultant slurry of metal salt crystals, and a resultant saturated stripping solution that is recirculated again to the stripping step (e) to contact additional loaded organic solution; and wherein furthermore the metal recovery step (f) involves recovering metal salt crystals and comprises the additional steps of: (g) separating the solid phase from the liquid phase of the slurry of metal salt crystals, thereby obtaining resultant metal salt crystals and a resultant liquid solution; (h) washing the metal salt crystals with water to remove impurities; and (i) drying the metal salt crystals.

As an option for the previous embodiment, prior to the solid/liquid separation step (g), the process may comprise the additional step of: (j) separating and recovering entrained organic solution from the slurry of metal salt crystals.

As another option for said embodiment, after the solid/liquid separation step (g) and prior to the water wash step (h), the process may further comprise the step of: (k) washing the metal salt crystals with an acid solution to minimize the precipitation of impurities.

In even other embodiments, the process further comprises the extraction and recovery of secondary and/or tertiary target metals, which are leached together with the (primary) target metals and are then either extracted in separate solvent extraction circuits or are extracted together in the same solvent extraction circuit and stripped in separate stripping steps or stripped together while crystallizing primary target metals and subjecting the other target metals to another metal recovery step, e.g. electrowinning. In these cases, the resulting aqueous stripping solutions are subjected to further metal recovery steps such as electrowinning, precipitation (crystallization, cementation), or other metal recovery procedures.

In one such embodiment, aimed to recover one or more secondary target metals remaining in the raffinate (I) solution from the extraction (I) step (b) so as to produce one or more metal products II, the process comprises the additional steps of: (I) contacting the raffinate (I) solution from the extraction (I) step (b) with a water immiscible organic (II) solution comprising one or more metal extracting agents or extractants (II) to selectively extract the metal ions from the raffinate (I) solution into the organic (II) solution, thereby providing a resultant loaded organic (II) solution containing secondary target metal(s) and a resultant raffinate (II) solution with reduced content of secondary target metal(s); (m) desorbing entrained ammonia from the loaded organic (II) solution with a sorption gas (II), thereby providing a resultant enriched gas (II) stream containing gaseous ammonia and a resultant desorbed loaded organic (II) solution with reduced content of entrained ammonia; (n) absorbing the gaseous ammonia from the enriched gas (II) stream in an aqueous absorbent (II) solution, thereby providing a resultant aqueous absorption (II) solution with increased content of dissolved ammonia that is recycled back to the leaching step (a), to the extraction (I) step (b), to the extraction (II) step (I), or to any combination of them, and a resultant exhaust gas (II) stream that is barren of ammonia; (o) contacting the desorbed loaded organic (II) solution with an acidic stripping (II) solution to strip the metal ions from the organic (II) solution, thereby providing a resultant pregnant electrolyte (II) stream with increased content of secondary target metal ions and a resultant stripped organic (II) solution with reduced content of secondary target metal ions; (p) recovering secondary target metal(s) from the pregnant electrolyte (II) stream in the form of one or more final metal products II.

In another embodiment, related to the previous one, the process comprises absorbing both the enriched gas (II) stream containing gaseous ammonia from the gas desorption (II) step (m) and the enriched gas (I) stream containing gaseous ammonia from the gas desorption (I) step (c) in an aqueous absorbent (l+ll) solution, thereby providing a resultant aqueous absorption (l+ll) solution with increased content of dissolved ammonia that is recycled back to the leaching step (a), to the extraction (I) step (b), to the extraction (II) step (I), or to any combination of them, and a resultant exhaust gas (l+ll) stream that is barren of ammonia.

In other embodiments of the invention, either one or both of the gas desorption (I or II) step(s) (c or m), either one or both of the gas absorption (I or II) step(s) (d or n), or any combination of them involves contacting the solution flow with the gas stream in a countercurrent manner. In yet other embodiments, the exhaust gas stream from either one or both of the gas absorption (I or II) step(s) (d or n), either in whole or in part, is recirculated to either one or both of the desorption step(s) (c or m). In even other embodiments, either one or both of the gas desorption (I or 11) step(s) (c or m) is operated at a temperature above ambient temperature, a pressure below atmospheric pressure, or both, whereas either one or both of the gas absorption (I or II) step(s) (d or n) is operated at a temperature below ambient temperature, a pressure above atmospheric pressure, or both.

In yet other embodiments, the dissolved ammonia content in either one or both of the aqueous absorbent (I or II) solutions from the gas absorption (I or II) step(s) (d or n) is further concentrated by using a steam stripping step, whereby said absorption solution is contacted with steam, thereby providing a resultant enriched steam with increased content of ammonia and a resultant depleted solution with reduced content of ammonia, which can be returned back to either one or both of the gas absorption (I or II) step(s) (d or n) as the absorption (I or II) solution or sent to other process steps, and a condensation step, whereby the enriched steam from the steam stripping step is changed from gas phase to liquid phase, by cooling, compression, or both, thereby providing a condensed enriched solution with higher ammonia concentration than the absorption solution, which is returned to the leaching step (a), to the extraction (I) step (b), to the extraction (II) step (I), or to any combination of them.

In even other embodiments, the aqueous absorbent (I or II) solution of either one or both of the gas absorption (I or II) steps (d or n) comprises, in whole or in part, the raffinate (I) solution from the extraction (I) step (b), the raffinate (II) solution from the extraction (II) step (I), the pregnant leach solution from the leaching step (a), or any combination of them.

In other embodiments, the process further comprises a wash the step of: (q) contacting the desorbed loaded organic (I or II) solution(s) from either one or both of the gas desorption (I or II) steps (c or m) with an acidic, neutral or basic wash affluent solution so as to remove further impurities, leaching agents, or both from the loaded organic (I or II) solution(s), thereby providing a resultant wash effluent solution with increased content of impurities, leaching agents, or both and a resultant washed loaded organic with reduced content of impurities, leaching agents or both, prior to sending the (washed) loaded organic (I or II) solution(s) to the stripping (I or II) step(s) (e or o).

In yet other embodiments, aimed to recover one or more tertiary target metal(s) from the loaded organic solution, the process further comprises the steps of: (r) contacting the desorbed (or washed) loaded organic solution with an acidic stripping (III) solution, which is less acidic than the stripping solution in the stripping (I or II) step(s) (e or o), so as to strip one or more tertiary target metal(s) from the desorbed loaded organic, thereby providing a resultant pregnant electrolyte (III) solution with increased content of tertiary target metal ions and a resultant stripped organic (III) solution with reduced content of tertiary target metal(s), prior to sending the stripped organic (III) solution to the stripping (I or II) step(s) (e or o); (s) recovering tertiary target metal(s) from the pregnant electrolyte (III) solution in the form of one or more final metal products III.

In even other embodiments, one or multiple stripping (I, II, or III) step(s) (e, o, or r), in any combination, involve an acidic stripping solution that is saturated with dissolved metal ions of one or multiple target metal(s), so as to strip the metal ions from the loaded organic solution and simultaneously precipitate salt crystals of one or multiple target metal(s) in the stripping solution, thereby providing a resultant stripped organic solution with reduced content of metal ions, a resultant slurry of metal salt crystals, and a resultant saturated stripping solution that is recirculated again to the stripping (I, II, or III) step (e, o, or r) to contact additional loaded organic solution; and wherein furthermore the metal recovery (I, II, or III) step (f, p, or s) involves recovering metal salt crystals and comprises the additional steps of: (t) separating and recovering entrained organic solution from the slurry of metal salt crystals; (u) separating the solid phase from the liquid phase of the slurry of metal salt crystals, thereby obtaining resultant metal salt crystals and a resultant liquid solution; (v) washing the metal salt crystals with water to remove impurities; and (w) drying the metal salt crystals.

The main claimed novelty of the present invention involves the removal and recovery of entrained ammonia from the loaded organic solution, in a solvent extraction circuit after ammoniacal leaching, by means of gas desorption and absorption with a sorption gas, preferably containing carbon dioxide and preferably low in oxygen gas content, so as to reduce ammonia entrainment to the stripping step, reduce consumption of ammoniacal leaching agents by recycling them back to leaching, allow the makeup of anions (particularly of carbonate ions) in the leach solution, avoid any potential oxidation (and hence degradation) of the organic extractant in solvent extraction, adjust the pH in the leach solution as well as in the absorption solution so as to improve leaching efficiency and the absorption of ammonia in the absorption solution, and act as an alternative to a wash step for the loaded organic solution when the water balance does not accommodate the additional amount of water that would be required to be disposed of in this case, i.e. the essential novelty of this invention relates mainly to the step (c) of gas desorption and the step (d) of gas absorption within a solvent extraction circuit. Other claimed novelties involve the optional use of higher temperature, lower pressure, or both in the gas desorption step (c), and the optional use of lower temperature, higher pressure, or both in the gas absorption step (d). Other claimed novelties involve the optional use of higher carbon dioxide content, lower oxygen gas content, or both in the sorption gas for the gas desorption step (c) and the gas absorption step (d).

Further claimed novelties of the present invention involve the overall process scheme (i.e. the combination of the different unit operations that constitute the process as a whole), and the production of high purity metal salts by means of removing organic entrained solution from the slurry of metal salt crystals after simultaneous stripping and crystallization and/or by the use of an acid wash step (k) after the solid/liquid separation step (g) and prior to the water wash step (h) to minimize the precipitation of impurities.

Yet other claimed novelties of the present invention, regarding the overall process scheme, involve the recovery and separation of primary and secondary target metals (or even tertiary target metals), allowing the production of high purity metal salts.

Other claimed novelties of the invention, within the overall process scheme, involve the use of a steam stripping step and a condensation step to further increase the concentration of ammonia in the absorption solution, prior to returning it to leaching.

Even other claimed novelties of the invention involve the use of gas desorption and absorption equipment in an ammoniacal leaching and solvent extraction plant and, in some embodiments related to the production of metal salts, the optional use of a mixer- thickener unit for stripping and crystallization, and even for organic recovery. Other claimed novelties include, for some embodiments of the invention, the implementation of an acid wash within a centrifuge for the metal salt crystals and the use of a column-type or rotating drum reactor for leaching.

Among the major advantages of the present invention are: (i) The invention allows the removal and recovery of entrained ammonia from the organic solution in a solvent extraction circuit after ammoniacal leaching, reducing the entrainment of ammonia to the stripping step and reducing the consumption of ammoniacal leaching agents by recycling them back to leaching, (ii) The invention presents an alternative to a wash step for dealing with entrained ammonia in the organic solution of a solvent extraction circuit, which is particularly suitable when the water balance does not accommodate the additional amount of water that would be required to be disposed of in such a case, (iii) The use of a sorption gas that contains carbon dioxide in preferred embodiments of the invention allows as much the recycling of ammonia as well as the replenishment of carbonate in the leach solution, contributing to the makeup of leaching agents, specifically of ammonia and ammonium carbonate (as well as of carbonate ions, in general), (iv) The carbon dioxide contained in the sorption gas for some preferred embodiments of the invention allows also adjusting the pH in the leach solution and in the absorption solution, improving leaching efficiency as well as the absorption of ammonia in the absorption solution (which is enhanced by the dissolution of carbon dioxide in the absorption solution), (v) The use of a sorption gas with low oxygen gas content in some preferred embodiments of the invention avoids the risk of possible oxidation (and hence degradation) of any extracting agent in the organic solution of solvent extraction, (vi) In some embodiments, the operation with a sorption gas at a higher temperature, a lower pressure, or both in the gas desorption step (c) improves gas desorption kinetics and allows a higher ammonia content in the enriched gas stream, (vii) In some embodiments, the operation with a sorption gas at a lower temperature, a higher pressure, or both in the gas absorption step (d) allows a higher ammonia concentration in the absorption solution, (viii) The concentration of ammonia in the absorption solution can be further increased in some embodiments of the invention by the use of steam stripping and condensation steps, (ix) In some embodiments, the invention allows the production of a solid metal salt at ambient temperature and atmospheric pressure directly from the loaded organic solution by crystallization within the stripping stage, i.e. by simultaneous stripping and crystallization, (x) In some embodiments of the invention, the use of an acid wash of the metal salt crystals inhibits the precipitation of undesired impurities in the metal salt crystals, allowing the impurities to be washed away more easily and thus increasing the purity of the final metal salt crystal products, (xi) Other embodiments of the invention allow the recovery and separation of different target metals, i.e. primary, secondary, and/or tertiary target metals from each other, by using either appropriate separate solvent extraction circuits, appropriate separate stripping stages within the same solvent extraction circuit, or an appropriate metal recovery step (e.g. electrowinning) after the stripping and crystallization stage within the same solvent extraction circuit, (xii) Some embodiments of the invention are particularly well suited for the extraction, separation, and purification of metals from secondary waste materials that originate from or comprise bottom ash (e.g. bottom ash deriving from municipal waste incinerators), metallic shredder material (e.g. from scrapped cars), or heavy metal fractions originating from WEEE material (waste electrical and electronic equipment).

BRIEF DESCRIPTION OF THE DRAWINGS

Figure 1 is a process flow sheet of one embodiment of the invention, showing the steps of leaching, extraction, gas desorption, gas absorption, stripping, and metal recovery.

Figure 2 is a process flow sheet of another embodiment of the invention, showing the steps of leaching, extraction, gas desorption, gas absorption (recirculation of exhaust gas to gas desorption), stripping, and metal recovery.

Figure 3 is a process flow sheet of yet another embodiment of the invention, showing the steps of leaching, extraction, gas desorption, gas absorption, steam stripping, condensation, stripping, and metal recovery.

Figure 4 is a process flow sheet of even another embodiment of the invention, showing the steps of leaching, extraction, gas desorption, gas absorption, loaded organic wash, stripping, and electrowinning.

Figure 5 is a process flow sheet of yet even another embodiment of the invention, showing the steps of leaching, extraction, gas desorption, gas absorption, stripping & crystallization, organic recovery, solid/liquid separation & wash, and drying.

Figure 6 is a process flow sheet of another embodiment of the invention, showing the steps of leaching, extraction, gas desorption, gas absorption, stripping I, metal recovery I, stripping II, metal recovery II, stripping III, and metal recovery III. Figure 7 is a process flow sheet of yet another embodiment of the invention, showing the steps of leaching, extraction, gas desorption, gas absorption, stripping I & crystallization, solid/liquid separation & wash, drying, stripping I I, and metal recovery I I .

Figure 8 is a process flow sheet of even another embodiment of the invention, showing the steps of leaching, extraction, gas desorption, gas absorption, stripping & crystallization, solid/liquid separation & wash, drying, and metal recovery I I .

Figure 9 is a process flow sheet of yet even another embodiment of the invention, showing the steps of leaching, extraction I , gas desorption I , stripping I , metal recovery I , gas absorption l+l l, extraction II , gas desorption II , stripping II , and metal recovery II .

DETAILED DESCRIPTION OF THE INVENTION

The present invention discloses a hydrometallurgical process and apparatus for the recovery of one or more target metals from a metalliferous feed material to produce one or more metal products by using ammonia leaching and solvent extraction, whereby entrained ammonia is recovered from the organic solvent by means of gas desorption with a sorption gas, preferably containing carbon dioxide, followed by gas absorption.

Target metals may include, but are not limited to, copper (Cu), zinc (Zn), nickel (Ni), cobalt (Co), silver (Ag), gold (Au), platinum (Pt), palladium (Pd), rhodium (Rh), mercury (Hg), chromium (Cr), cadmium (Cd), molybdenum (Mo), rhenium (Re), among others.

The metalliferous feed material, also referred to as metal bearing feed material, metal containing feed material, or simply feed material, may comprise, but is not limited to, bottom ash, metal scrap, waste, ore, concentrates, tailings, slags, or any other solid (feed) material containing one or more target metals. The feed material may be dry or wet, or even in the form of a slurry or pulp. It may have any particle size distribution and may even consist of a run-of-mine (ROM) material.

Some embodiments of the invention are particularly well suited for the extraction, separation, and purification of metals from secondary waste materials that originate from or comprise bottom ash (e.g. bottom ash deriving from municipal waste incinerators), metallic shredder material (e.g. from scrapped cars), or heavy metal fractions originating from WEEE material (waste electrical and electronic equipment).

The final metal products may comprise any product produced from the target metal(s), including, but not limited to, metal cathodes, metal powder, metal salt crystals, metal bearing solutions, metal bearing pulps, metal alloys, metal oxides, metal sulfides, and combinations thereof, among others. The metal salt crystals, in particular, may comprise sulfates (with anions SO4 ^2" , HSCV), carbonates (with anions CO3 ^2" , HCO3 ^" ), chlorides (with anion CI ^" ), phosphates (with anion PO4 ^3" ), nitrates (with anion NO3 ^" ), acetates (with anion CH3COO ^" ), oxalates (with anion C2O4 ^2" ), among many others. Some examples of high purity metal products are copper (Cu) cathodes, zinc (Zn) cathodes, nickel (Ni) cathodes, cobalt (Co) cathodes, copper sulfate pentahydrate (CuSCvSh O), zinc sulfate heptahydrate (ZnSCv7H20), nickel sulfate heptahydrate (NiSC h O), and cobalt sulfate heptahydrate (CoSC h O), among others.

The sorption gas (also called scrubbing gas) employed for gas desorption and absorption may comprise any gas able to desorb and transfer ammonia from the organic solution and, preferably, may comprise any gas containing carbon dioxide (CO2), so as to adjust the pH of the absorption solution to the right value and replenish carbonate ions in the leach solution. The sorption gas may comprise any combustion gas derived from the combustion of an organic fuel such as wood, coal, charcoal, petrol, diesel, kerosene, methane, propane, ethane, methanol, ethanol, hexane, benzene, paraffin wax, naphthalene, polyethylene, polypropylene, polystyrene, organic waste, among many other hydrocarbons. The sorption gas may also correspond to any mixture of a combustion gas and an inert gas, or of carbon dioxide and an inert gas, in particular to any mixture of carbon dioxide and air, or of carbon dioxide and nitrogen (N2). The content of carbon dioxide in the sorption gas is preferably in the range between 400 and 100000 ppm, and more preferably in the range between 1000 and 50000 ppm. The sorption gas may also comprise air (which already contains trace amounts of carbon dioxide), or any mixture of air and an inert gas, in particular it may comprise any mixture of air and nitrogen (N2). The sorption gas, preferably, has a low oxygen gas (O2) content, so as to avoid any risk of possible oxidation of any extracting agent in the organic solution of solvent extraction. Such an oxidation may eventually cause degradation of extracting agents (hence increasing their consumption rate), higher viscosity of the organic solution, higher phase disengagement time, more entrainments, and crud, among other difficulties. Although, standard mixer-settlers (e.g. for ketoxime- and aldoxime-based extracting agents) typically operate in contact with air (which contains oxygen gas), or sometimes even with the introduction of air into the mixer (as aeriation), without any noticeable oxidation being observed. The content of oxygen gas in the sorption gas is preferably below 10 v%, and more preferably below 1 v%. The sorption gas may also comprise any gas containing sulfur dioxide (SO2), which would also allow adjusting the pH of the absorption solution to the right value and replenish now sulfate ions (instead of carbonate ions) in the leach solution, but it has the big drawback that it is known that it reduces (and therefore degrades) the organic extractant (particularly when dealing with ketoxime- and aldoxime-based extractants). As another alternative, the sorption gas may also comprise any gas containing chlorine gas (C ), which would also allow adjusting the pH of the absorption solution to the right value and replenish now chloride ions (instead of carbonate ions) in the leach solution, but it has the big drawback that it is known that it oxidizes (and hence degrades) the organic extractant (particularly when dealing with ketoxime- and aldoxime-based extractants, where the formation of chlorine gas in EW is known to completely destroy the entrained organic extractant). Hence, the sorption gas more preferably corresponds to a mixture of carbon dioxide and nitrogen gas. It may also comprise a mixture of carbon dioxide, nitrogen gas, a combustion gas, and (preferably in a rather small amount) air.

In one embodiment of the invention, shown in Figure 1 , the process comprises the step of leaching (100) the metalliferous feed material (5) with a (lean) leach solution or LLS (10) containing ammoniacal leaching agents (10) to produce a resultant pregnant leach solution or PLS (20) containing dissolved metal ions of the target metal(s) and a resultant (leached) waste material (15) with reduced content of target metal(s). The pregnant leach solution (20) is sent to an extraction step (200) wherein the pregnant leach solution (20) is contacted with a water immiscible organic solution (35) comprising one or more metal extracting agents or extractants so as to selectively extract the metal ions from the pregnant leach solution (20) into the organic solution (35), thereby providing a resultant loaded organic (LO) solution (40) containing target metal(s) and a resultant raffinate solution (30) with reduced content of target metal(s), which is returned back to the leaching step (100) as the leach solution (10). The loaded organic solution (40) is sent to a gas desorption step (300) wherein the loaded organic solution (40) is contacted with a desorption gas (45) (i.e. the sorption gas that enters the gas desorption step) to remove and recover entrained leaching agents (i.e. ammonia) from the organic solution (40), thereby providing a resultant desorbed loaded organic solution (50) with reduced content of entrained ammonia and a resultant enriched gas stream (55) with increased content of ammonia. The enriched gas stream (55) is sent to a gas absorption step (600), wherein the enriched gas (55) is contacted with an aqueous absorbent solution (65) to transfer ammonia and, if present, carbon dioxide from the enriched gas stream (55) to the absorbent solution (65), thereby providing a resultant exhaust gas stream (60) with reduced content of ammonia and, in some cases, of carbon dioxide and a resultant aqueous absorption solution (70) with increased content of ammonia, ammonium ions, or both, as well as, in some cases, carbonate ions, which is returned back to the leaching step (100). As an alternative, the absorption solution (70) may be returned also to the extraction step (200), or to both the leaching step (100) and the extraction step (200). The desorbed loaded organic solution (50) is sent to the stripping step (400), wherein the (desorbed) loaded organic solution (50) is contacted with a spent electrolyte (or stripping) solution (75) to strip the metal ions from the loaded organic solution (50) into the spent electrolyte (75), thereby providing a resultant pregnant electrolyte (PE) solution (80) with increased concentration of metal ions and a resultant stripped organic (SO) solution (35) with reduced content of metal ions, which is returned back to the extraction step (200) to contact additional pregnant leach solution (20). The pregnant electrolyte (80) is sent to a metal recovery step (500) to produce the final metal product (85), returning the resultant spent electrolyte (SE) solution (75) with reduced content of metal ions back to the stripping step (400).

In another embodiment of the invention, the absorbent solution (65) for the gas absorption step (600) may comprise the raffinate solution (30), in whole or in part, returning thereafter the resultant absorption solution (70) back to the leaching step (100) as the leach solution (10), or as part of it.

In yet another embodiment of the invention, the absorbent solution (65) for the gas absorption step (600) may comprise the pregnant leach solution (20), in whole or in part, sending thereafter the resultant absorption solution (70) to the extraction step (200) as the solvent extraction (SX) feed, or as part of it.

In even another embodiment of the invention, as shown in Figure 2, the exhaust gas stream (60), either in whole or in part, may be recirculated back from the gas absorption step (600) to the gas desorption step (300). The part of the exhaust gas stream that is not recirculated back from the gas absorption step (600) to the gas desorption step (300) becomes a bleed of the recirculating sorption gas stream. Such a bleed may also comprise the enriched gas stream (55), although not as a preferred option. Desorption gas (45) is added as makeup for gas entrainment losses (particularly of carbon dioxide) in the solutions, particularly in the absorption solution (70). The sorption gas makeup is added preferably to the gas desorption step (300), but may be added also to the gas absorption step (600) or elsewhere to the gas recirculation loop.

In yet other embodiments of the invention, the concentration of ammonia in the absorption solution from the gas absorption step (600) is increased prior to recycling the ammonia in solution back to the leaching step (100), to the extraction step (200), or to both. In one such embodiment of the invention, as illustrated in Figure 3, the process further comprises a steam stripping step (1100), whereby the absorption solution (70) from the gas absorption step (600) is contacted with steam (96), i.e. water vapor, thereby providing a resultant enriched steam (97) with increased content of ammonia and a resultant depleted solution with reduced content of ammonia, which can be returned back to the gas absorption step (600) as the absorbent solution (65), in whole or in part, or sent to other process steps. In this embodiment, the process further comprises a condensation step (1200), whereby the enriched steam (97) from the steam stripping step (1100) is changed from gas phase to liquid phase, by cooling and/or compression to the saturation limit when the molecular density in the gas phase reaches its maximal threshold, thereby providing a condensed enriched solution (98) with higher ammonia concentration than the absorption solution (70), which is recirculated to the leaching step (100), to the extraction step (200), or to both.

In another embodiment of the invention, as shown in Figure 4, the desorbed loaded organic solution (50) is sent to a loaded organic (LO) wash step (700) to remove entrained impurities, leaching agents (e.g. ammonia), or both prior to the stripping step (400), wherein the desorbed loaded organic (50) is contacted with an acidic, neutral or basic (preferably slightly acidic) aqueous loaded organic wash affluent solution (42), thereby providing a resultant washed loaded organic solution (51) with reduced content of entrained impurities and/or leaching agents, and a resultant aqueous loaded organic wash effluent solution (43) with increased concentration of impurities, leaching agents, or both.

In yet another embodiment of the invention, the stripped organic solution (35) may also be subjected to a stripped organic (SO) wash step to remove entrained impurities (e.g. acid), wherein the stripped organic solution (35) is contacted with an acidic, neutral or basic (preferably neutral or slightly acidic) aqueous stripped organic wash affluent solution prior to be returned to the extraction step (200).

The metal recovery step (500) may involve any process to produce one or more final metal products (85) from the pregnant electrolyte solution (80). It may involve electrowinning (as shown in Figure 4), precipitation (crystallization or cementation), smelting, or combinations thereof, among others. Precipitation by crystallization may involve cooling, evaporation, change of pH, change of cations or anions, addition of a second solvent to reduce the solubility of the solute (e.g. ethanol in sulfate media), among several other methods. In some embodiments, e.g. when the final metal product (85) is a metal solution and, in particular, the pregnant electrolyte solution (80) itself, then the spent electrolyte (75) is a fresh solution that is added to the stripping step (400) and is not returned from the metal recovery step (500). In other embodiments, e.g. when the final metal product (85) involves metal cathodes or metal salt crystals, one or more metal product refining steps may be applied, e.g. involving electrorefining (for metal cathodes) or re-crystallization (for metal salt crystals).

In another embodiment of the invention, the metal recovery step (500) involves, in particular, simultaneous stripping and crystallization, as shown in Figure 5, to produce metal salt crystals as final metal product (85). In this embodiment, the stripping step becomes a simultaneous stripping & crystallization step (400) wherein the desorbed loaded organic (50) is contacted with a saturated stripping solution (90) that is saturated with dissolved metal ions of the target metal(s), to strip the metal ions from the desorbed loaded organic solution (50) and precipitate salt crystals of the target metal(s) in the stripping solution (90), thereby providing resultant precipitated metal salt crystals (31) in the saturated stripping solution (90) and a resultant stripped organic solution (35) with reduced content of metal ions. A slurry of the precipitated metal salt crystals (31 ), the remaining portion of the saturated stripping solution (90), and the stripped organic solution (35) are then separated. The remaining portion of the saturated stripping solution (90) is recirculated back to the stripping and crystallization step (400) to contact additional desorbed loaded organic solution (50) and the stripped organic solution (35) is returned back to the extraction step (200) to contact additional pregnant leach solution (20). To minimize the accumulation of impurities that might contaminate the metal salt crystals, a small amount of bleed solution (33) can be bled from the recirculating saturated stripping solution (90). This bleed solution (33) may be added to the leach solution (10), it may be added to or become the absorbent solution (65), it may be disposed of together with the leached waste material (15), it may be sent to other process steps, or it may be added to other process solutions. An acid solution (32) can be added to the recirculating saturated stripping solution (90) to provide hydrogen ion for stripping the target metal(s) from the desorbed loaded organic solution (50) and to replenish the metal salt anion(s) removed by crystallization of the metal salt. The bleed solution (33) is withdrawn from the saturated stripping solution (90) preferably before adding the acid solution (32). The slurry with the precipitated metal salt crystals (31) may be sent to an optional organic recovery step (800) to separate and recover entrained organic solution (62) from the slurry of metal salt crystals (31), thereby providing a resultant metal salt slurry (61) with reduced content of entrained organic solution. The metal salt slurry (61), or, alternatively when the organic recovery step (800) is omitted (as illustrated in Figures 7 and 8), the slurry with the precipitated metal salt crystals (31), is then sent to a solid/liquid separation & wash step (900), wherein the solid phase is separated from the liquid phase of the slurry, thereby providing metal salt crystals and a liquid solution, and where thereafter the metal salt crystals are washed with water (71) to remove impurities, separating again the solid phase from the liquid phase, thereby obtaining washed metal salt crystals (84) and a wash solution (72) containing the separated liquids. The wash solution (72) is returned preferably to the recirculating saturated stripping solution (90). The amount of water (71) added to wash the metal salt crystals preferably balances the amount of water lost as crystallization water entrapped within the washed metal salt crystals (84) and the water withdrawn as the bleed solution (33). The metal salt crystals (84) are sent to a drying step (1000) wherein the washed metal salt crystals (84) are dried to remove water moisture, thereby providing high purity metal salt crystals (85) as final metal product.

In another embodiment of the invention, related to the previous embodiment and as illustrated in Figure 7, the process comprises, in the solid/liquid separation & wash step (900), washing the metal salt crystals with an acid solution (73) to avoid or at least minimize the precipitation of impurities after the first separation of solids from liquids and prior to the washing with water (71). The liquid phase from both the acid wash and the water wash, after its separation from the solid phase, is then returned as the wash solution (72) preferably to the recirculating saturated stripping solution (90).

In yet other embodiments of the invention, with the aim of producing metal salt crystals of even higher purity, the process further comprises, after the solid/liquid separation & wash step (900), a re-crystallization step, wherein the washed metal salt crystals (84), or, alternatively, the dried metal salt crystals (85), are dissolved in a (preferably aqueous) solution and then precipitated again as metal salt crystals, e.g. by cooling, evaporation, change of pH, change of cations or anions, addition of a second solvent to reduce the solubility of the solute, among several other methods, prior to separating the resulting metal salt crystals of higher purity from the liquid, thereby providing a resultant solution with increased content of impurities and resultant metal salt crystals of higher purity. Further re- crystallization, solid/liquid separation & wash, and/or drying steps may be implemented prior to the final drying step (1000). A re-crystallization step may become necessary to remove entrained impurities and, in particular, remaining entrained organic solution (or at least remove its smell).

In even other embodiments, the invention further comprises the recovery and separation of secondary and/or tertiary target metals, which may include, but are not limited to, copper (Cu), zinc (Zn), nickel (Ni), cobalt (Co), silver (Ag), gold (Au), platinum (Pt), palladium (Pd), rhodium (Rh), mercury (Hg), chromium (Cr), cadmium (Cd), molybdenum (Mo), rhenium (Re), among others. The secondary and tertiary target metals are leached together with the (primary) target metals. In some embodiments, they may be also extracted together with the primary target metals into the organic solution, but are then separated from the primary target metals by separate stripping steps. In one embodiment, as shown in Figures 6 and 7, the secondary target metals may be separated by a stripping (II) step (410) after the stripping (I) step (400) and prior to returning the stripped organic (II) solution (35) to the extraction step (200), whereby the stripped organic (I) solution (36) from the stripping (I) step (400) is contacted with an aqueous stripping or spent electrolyte (II) solution (76) of preferably higher acidity (i.e. of lower pH) than the one of the spent electrolyte (I) solution (75) (or of the saturated stripping (I) solution (90), depending on the embodiment), thereby providing a resultant pregnant electrolyte (II) solution (81) with increased content of secondary target metals that is sent to a metal recovery (II) step (510) to produce a metal product II (86) and a resultant stripped organic (II) solution (35) with reduced content of secondary target metals that is returned to the extraction step (200). In another embodiment, as shown in Figure 6, the tertiary target metals may be separated by a stripping (III) step (420) after the extraction step (200) and prior to the stripping (I) step (400), after the gas desorption step (300), whereby the loaded organic (40), the desorbed loaded organic (50), or both are contacted with an aqueous stripping or spent electrolyte (III) solution (77) of preferably lower acidity (i.e. of higher pH) than the one of the spent electrolyte (I) solution (75) (or of the saturated stripping (I) solution (90), depending on the embodiment), which may even be of alkaline nature, thereby providing a resultant pregnant electrolyte (III) solution (82) with increased content of tertiary target metals that is sent to a metal recovery (III) step (520) to produce a metal product III (87) and a resultant stripped organic (III) solution (37) with reduced content of tertiary target metals. Optionally the stripping (III) step can be carried out before the gas desorption step (300). The resultant stripped organic (III) solution (37) with reduced content of tertiary target materials is sent either to the gas desorption step (300) or to the stripping (I) step (400), depending on the embodiment.

In even other embodiments of the invention, wash steps and/or gas desorption steps may be implemented prior, between, or after the stripping steps (II) and (III), as well of course for the stripping step (I). The aqueous electrolyte or stripping solutions from the stripping steps (II) and (III) are sent to corresponding metal recovery steps (I I) and (III) to extract secondary and/or tertiary target metals and produce corresponding metal products II and III. The metal recovery steps (II) and (III) may involve electrowinning, precipitation, crystallization or other metal recovery procedures.

In yet another embodiment of the invention, as shown in Figure 8, when secondary target metal(s) are leached, extracted, and stripped together with the primary target metal(s) and when producing metal salt crystals as final metal product (85) for the primary target metal(s), the process further comprises a metal recovery (II) step (510) to recover secondary target metal(s) from the saturated stripping solution that is recirculated to the stripping & crystallization step (400) so as to produce a metal product II (86). The saturated stripping solution with increased content of secondary target metal(s) that is sent from the stripping & crystallization step (400) to the metal recovery (II) step (510) is denoted by saturated pregnant electrolyte (II) solution (81), whereas the saturated stripping solution with reduced content of secondary target metal(s) that is returned from the metal recovery (II) step (510) to the stripping & crystallization step (400) is denoted by saturated spent electrolyte (II) solution (76). In this embodiment, the metal recovery (II) step (510) preferably involves electrowinning, although it may also involve precipitation (crystallization or cementation), among others. A portion of the saturated spent electrolyte (II) solution (76) may be recycled back to the saturated pregnant electrolyte (II) solution (81), e.g. to lower the content of primary target metal(s) below saturation if the secondary target metal(s) coincide with them, i.e. if the metal recovery (II) step (510) involves the primary target metal(s). The bleed solution (33) is preferably withdrawn from the saturated spent electrolyte (II) solution (76) after the metal recovery (II) step (510) and not before, preferably before adding the acid solution (32). For example, the primary target metal can be zinc and the secondary target metal can be copper, being both extracted from the pregnant leach solution (20), e.g. by an extractant such as an aldoxime (e.g. LIX 860N-I), being both stripped together in the stripping & crystallization step (400), but being zinc simultaneously crystallized therein (e.g. as zinc sulfate) whereas copper is subjected to electrowinning in the metal recovery (II) step (510) following thereafter, which allows keeping the copper concentration in the spent electrolyte (II) solution (76) below saturation while keeping the zinc concentration at saturation level. In another example, both the primary target metal and the secondary target metal can be copper (or zinc), being subjected to stripping & crystallization (400) and thereafter to electrowinning (510), allowing thus to simultaneously produce copper salt crystals (e.g. copper sulfate) and copper cathodes (or, when the target metal is zinc, e.g. zinc sulfate and zinc cathodes).

In other embodiments, secondary and tertiary target metals leached together with the primary target metals may be recovered in one or more separate solvent extraction circuits, either before or after extracting the primary target metals. Such separate solvent extraction circuits may involve the same or different extracting agents and/or organic solvents. The stripping solutions from such parallel solvent extraction circuits are subjected to further metal recovery steps such as electrowinning, precipitation, crystallization or other metal recovery procedures. In other embodiments, wash steps and/or gas desorption steps may be implemented prior or after the stripping steps of the parallel solvent extraction circuits. The wash solutions of the wash steps from the parallel solvent extraction circuits may be connected.

In some embodiments of the invention, related to the recovery of multiple

(primary, secondary, tertiary, etc.) target metals in separate solvent extraction circuits, a single gas absorption step can be used for several separate solvent extraction circuits. As shown in Figure 9, primary and secondary target metals are leached together in the leaching step (100), but are then treated in separate solvent extraction circuits I and II so as to produce respectively a final metal product I (85) and a final metal product II (86). However, both the enriched gas (I) stream (55) from the gas desorption (I) step (300) and the enriched gas (II) stream (56) from the gas desorption (II) step (310) are sent to a gas absorption step (600) that is shared by both solvent extraction circuits, returning thereafter the absorption solution (70) back to the leaching step (100).

Leaching

The leaching step (100) involves the dissolution of desired target metals into an aqueous phase by contacting the (lean) leach solution (10), which contains (ammoniacal) leaching agents, with the metalliferous feed material (5), thereby obtaining a pregnant leach solution or PLS (20) with dissolved target metals and a (leached) waste material (15) with reduced content of target metals, which can be disposed off (e.g. in tailings) or subjected to other process steps.

The ammoniacal leaching agents used during the leaching step (100) may comprise ammonia (NH3), ammonium salts such as ammonium carbonate ((NhU^COs), ammonium bicarbonate (NH4HCO3), ammonium sulfate ((NhU^SC ), ammonium chloride (NH4CI), ammonium nitrate (NH4NO3), ammonium hydroxide (NH4OH), ammonium thiosulfate (NH ₄ )2S203, among other leaching agents, and mixtures thereof. In some embodiments, the (lean) leach solution (10) may comprise in addition leaching agents selected from the group consisting of carbon dioxide (CO2), sulfur dioxide (SO2), carbonic acid (H2CO3), sulfurous acid (H2SO3), sulfuric acid (H2SO4), hydrochloric acid (HCI), nitric acid (HNO3), formic acid (HCOOH), acetic acid (CH3COOH), oxalic acid (HOOCCOOH), hydrogen cyanide (HCN), carbonates in general (CO3 ^2" ), among others, and combinations thereof.

The leaching agents may be added to the leach solution (10) or directly to the metalliferous feed material (5). Some leaching agents may even be added during the leaching step (100).

In addition, oxidizing or reducing agents may be added together with the leaching agents to the leaching step (100), which may comprise oxygen (O2), air, oxygen- nitrogen mixtures, chlorine (CI2), hydrogen peroxide (H2O2), sulfur dioxide (SO2), hydrogen sulfide (H2S), sodium hypochlorite (NaCIO), elemental zinc (Zn°), ferric compounds (Fe ³⁺ , e.g. in combination with complexing agents such as EDTA), among others, and combinations thereof.

The leaching step (100) may involve heap leaching, agitation leaching, in-situ leaching, dump leaching, vat leaching, pressure leaching, or any other kind of leaching technique. Prior to leaching, the metalliferous feed material (5) may be subjected to comminution (e.g. crushing, grinding, etc.), separation/concentration (e.g. screening, sieving, etc.), drying, roasting, blending, agglomeration, curing, oxidation, reduction, among many others process steps. Leaching may be thus performed preferably either in a tank, a vessel, a reactor, a column, a heap, a pile, a dump, a vat, or in situ, among other alternatives, and more preferably in a heap, a stirred tank reactor, a rotating drum reactor, a column-type reactor, or any combination thereof. In a particularly preferred embodiment of the invention, particularly when dealing with secondary waste materials related to bottom ash (e.g. from municipal waste incinerators), shredded metallic devices (e.g. scrapped cars), or WEEE material (waste electrical and electronic equipment), the leaching operation is performed in a column-type or rotating drum reactor, which enables efficient mixing of the leach solution and the metalliferous feed material. Leaching is performed preferably at a pressure in the range between 0.2 and 2 atm and at a temperature in the range between 2°C and 60°C, more preferably at a pressure in the range between 0.8 and 1.2 atm and at a temperature in the range between 10°C and 40°C, and most preferably at atmospheric pressure and ambient temperature.

In this description, ambient temperature refers to a temperature in the range of

10°C to 30°C, for example between 15°C to 30°C and atmospheric pressure refers to a pressure of about 1 atm.

The leaching step (100) may be carried out in a single or in multiple leaching stages, which are operated in batch or continuous, and may involve different leach solutions which contact consecutively the metalliferous feed material (5).

In one embodiment of the invention, which is particularly well suited for the treatment of secondary waste materials that originate from or comprise bottom ash, the metalliferous feed material (5) preferably comprises larger particles, i.e. having a particle size in the range of 0-40 mm, more preferably in the range of 1-20 mm, and most preferably up to about 4 mm. As an alternative it is also possible to process separate streams of a small fraction, for example, streams of particles having a particle size in the range of 0-4 mm, preferably in the range of 1-3 mm, and a large fraction, typically in the range of 4-12 mm, preferably in the range of 5-10 mm, by using a sieve to obtain these separate streams. These separate streams may then be leached separately.

Particle size as used in this description refers to smallest size of a particle; for example, for a particle having two different dimensions, it is considered that the particle size is the smallest dimension. When refereeing to a sieve to obtain the separate streams, reference is made to a method in which the material is allowed to pass through a series of sieves of progressively smaller mesh size; therefore, separating the materials in streams of particles having a specific particle size.

Conventionally, when leaching in stirred-tank reactors, large-sized particles hamper the stirring and, consequently, leaching is slowed down and/or becomes unselective. It was found that by using a column-type or rotating drum leaching reactor, leaching a secondary waste material could be carried out in a more effective manner, particularly in the case of a copper- and zinc-bearing material.

A metalliferous feed material (5) comprising particles that have specific desired size can be obtained from ballistic separation devices, Eddy-current separators, density separators, heavy medium separators, magnetic fluid separators, gravity separators, jigs, or sensor sorting, among others, e.g. particles that have a particle size of up to 20 mm can be obtained from the ballistic separation device described in the international patent application WO 2009/123452 "Separation-apparatus" by Berkhout & Rem, and commercially available as the ADR from Inashco B.V., Rotterdam, the Netherlands. In a preferred embodiment, related to the treatment of secondary waste material, the metalliferous feed material (5) comprises a non-ferrous secondary material, preferably a non-ferrous waste material, such as non-ferrous waste material originating from or comprising bottom ash (e.g. bottom ash originating from municipal waste incinerators), shredded metallic devices (such as cars), or WEEE material.

In one embodiment of the present invention, when dealing with secondary waste materials, more desired target metals such as copper and zinc are substantially selectively leached from the metalliferous feed material (5) while more noble metals such as gold and silver remain in the residue. In this embodiment, metals such as aluminum and iron are typically removed before the leaching step by conventional methods such as Eddy-current and/or magnetic separation. Therefore, although in this case the metalliferous feed material (5) may contain metals other than copper and zinc, the resulting pregnant leach solution or PLS (20) typically essentially only contains copper and zinc as the target metals present in the PLS. Essentially containing only copper and zinc means in the context of the present invention that other metals may be present in the PLS in an amount of less than 100 ppm. For instance, amounts of lead of about 62 ppm, of tin of about 6 ppm, of iron of about 2 ppm and other metals, generally in amounts of less than 1 ppm may be present, but these amounts may be considered negligible relative to the amount of copper and zinc, which are typically present in the PLS in amounts of e.g. about 2-30 and 1-10 g/L respectively. The amount of other metals in the PLS may vary depending on the composition of the PLS (e.g. ammonia concentration, pH, etc.). As such, essentially containing means that generally more than 30 wt%, preferably more than 70 wt%, and more preferably more than 90 wt% of the metals present in the PLS are zinc and copper. The metal components other than copper and zinc are typically discarded as residual materials in this embodiment. The selective leaching according to the present invention is preferably carried out with a (lean) leach solution or LLS (10) comprising an aqueous ammoniacal solution. For instance, in the case of copper, it was found that copper is leached less efficiently at high pH, because copper hydroxide is formed when the pH is too high. The copper hydroxide tends to precipitate in the solution, thereby rendering the leaching less efficient, while copper ammonia complexes remain soluble in the solution. As such, in this case the aqueous ammoniacal solution preferably has a pH in the range from 7.0 to 1 1.0, and more preferably in the range of 8.0 to 10.5. This solution can be obtained by dissolving ammonia and an acid in water. In addition, adding an acid is preferable since hydrogen ions are consumed in the leaching reaction (e.g. according to the reactions 2Cu + O2 + 4H ⁺ → 2Cu ²⁺ + 2H ₂ 0 and 2Zn + 0 ₂ + 4H ⁺ → 2Zn ²⁺ + 2H ₂ 0).

It was found that the non-ferrous secondary material as described for the embodiment herein above is characterized by a high permeability, allowing high flow rates of the leach solution through the packed material, due to which a column-type reactor is favorable to use. This improves the efficiency of the present invention, because less dead zones are created and the wetting of the material is better.

In other embodiments, a rotating drum leaching reactor is used, since it enables efficient mixing of the leaching solution and the material.

In the case of a metallic copper- and zinc-bearing feed material that comprises a copper concentration of less than 70 wt% and a zinc concentration of more than 10 wt%, it may be preferred that the leaching comprises a pre-leaching stage to selectively leach zinc and produce an early pregnant leach solution. This results in an overall increased efficiency of the process for this particular case, because metallic zinc protects the cathodic copper from oxidation, which inhibits the dissolution of copper. The early pregnant leach solution, in this case, comprises zinc as the substantially only metal, remaining copper in the residual material. However, the pre-leaching stage does not necessarily leach all zinc from the zinc- and copper-containing feed material and the pre-leaching stage is therefore typically followed by one or more subsequent leaching stages to produce another pregnant leach solution. The early pregnant leach solution may be joined with the raffinate (I) solution (30) that is obtained by the selective solvent extraction of the pregnant leach solution, produced in the leaching stages after the pre-leaching stage. Alternatively, the early pregnant leach solution may be processed separately to recover the zinc contained therein. In this embodiment, the pre-leaching stage is preferably carried out in a column- type or rotating drum leaching reactor since this allows good control over the leaching parameters (e.g. contact time between the leaching solution and the material) and facilitates selective zinc leaching.

The rotating drum leaching reactor comprises a barrel that is adapted to be driven by a motor at relatively slow rotational speeds of typically up to 50 rpm, preferably up to 10 rpm. This relatively slow rotation provides good mixing of the feed material to be leached (5) and the lean leach solution (10), but also result in a relatively low consumption of energy for mixing when compared to e.g. stirred-tank reactors. The barrel can comprise one or more baffles to divide the barrel into partitions such that the back mixing of solids and liquids during continuous operation is limited. In addition, the barrel may comprise lifters to facilitate the tumbling of the solids and to limit sliding of the zinc- and copper-containing material in the barrel upon rotation. Examples of a suitable rotating drum leaching reactor are described in U.S. patent 3,400,871 "Apparatus for continuous metal extraction" by Davis and U.S. patent 6,613,271 "Apparatus and methods for recovering valuable metals" by Lewis Gray, which are both incorporated herein.

The rotating drum reactor further comprises an opening for feeding the zinc- and copper-containing material. This opening can be covered by a lid (preferred for use in batch processes) or may be connected to a filling hopper (preferred for use to continuous feed the material). An unloading system that is adapted to unload the PLS (20) is typically placed at the other side of the barrel, in particularly if the rotating drum reactor is adapted for use in a continuous process.

The leaching step (100) may comprise a solid/liquid separation step after the leaching and before the pregnant leach solution (20) is continued in the process. The solid/liquid separation step can, for instance, be carried out with a filter or a centrifuge that is placed in the outlet of the leaching reactor.

Solvent extraction

The extraction (200), such as solvent extraction, of target metal(s) involves contacting the pregnant leach solution (20) from the leaching step (100) with a water immiscible organic solution (35) comprising one or more metal extractants so as to selectively extract the metal ions from the pregnant leach solution (20) into the organic solution (35), thereby providing a resultant loaded organic solution (40) containing target metal(s) and a resultant raffinate solution (30) with reduced content of target metal(s), which is returned back to the leaching step (100).

The organic solution (35) in the extraction step (200) is immiscible in water and comprises one or more metal extracting agents dissolved in a water-immiscible organic solvent or diluent. Such solvents may include, but are not limited to, aliphatic and aromatic hydrocarbons such as kerosene, benzene, toluene, xylene, and the like, among others, and mixtures thereof. Preferred solvents for the use in metal recovery in the present invention are aliphatic and aromatic hydrocarbons having flash points of 65°C and higher, and solubility in water of less than 0.1 % by weight. Aliphatic hydrocarbons are typically preferred over aromatic hydrocarbons due to the possibility of higher ammonia entrainments for the latter. These solvents are also essentially non-toxic, chemically inert, and the costs thereof are currently within practical ranges. Some commercially available suitable organic solvents are, e.g. Escaid 120 (sold by Exxon Mobile Corporation, having a flash point of about 103°C and an aromatic content < 0.50 wt.%) and ShellSol D100 (sold by Royal Dutch Shell, having a flash point of about 103°C and an aromatic content of about 0.02 wt.%).

The metal extracting agents or extractants of the organic solution may comprise oximes such as 5-nonylsalicylaldoxime (also called C9 aldoxime, NSAO, or simply aldoxime, sold as LIX 860N-I by BASF), 5-dodecylsalicylaldoxime (also called C12 aldoxime or DSAO, sold as LIX 860-I by BASF), 5-nonyl-2-hydroxy-acetophenone oxime (also called C9 ketoxime, HNAO, or simply ketoxime, sold as LIX 84-I by BASF), 5-dodecyl-2-hydroxy- acetophenone oxime (also called C12 ketoxime), 5,8-diethyl-7-hydroxy-dodecan-6-oxime (sold as LIX 63 by BASF), or the recently developed reagents described in U.S. patent 8,475,748 "Metal solvent extraction reagents and use thereof by Virnig, Bender & Emmerich from BASF, namely 3-methyl-5-nonylsalicylaldoxime, 3-methyl-5-nonyl-2- hydroxyacetophenone oxime, 3-methyl oxime, 3-methyl ketoxime, and 3-methyl aldoxime, among others, and mixtures thereof. The extractants may also comprise organophosphorus compounds such as di-2-ethylhexyl phosphoric acid (also called D2EHPA, DEHPA, or HDEHP), dinonyl phenyl phosphoric acid (also called DNPPA), 2-ethylhexyl phosphonic acid mono-2-ethylhexyl ester (also called HEH/EHP, sold as PC88A by Daihachi Chemical Industry and as lonquest 801 by Rhodia), bis-2,4,4-trimethylpentyl phosphinic acid (sold as Cyanex 272 by Cytec Solvay), bis-2,4,4-trimethylpentyl-dithiophosphinic acid (sold as Cyanex 301 by Cytec Solvay), bis-2,4,4-trimethylpentyl-monothiophosphinic acid (sold as Cyanex 302 by Cytec Solvay), trioctylphosphine oxide (also called TOPO, sold as Cyanex 921 by Cytec Solvay), trialkylphosphine oxides (a mixture of 4 trialkylphosphine oxides, namely trihexylphosphine oxide, dihexylmonooctyl-phosphine oxide, dioctylmonohexyl- phosphine oxide, and trioctylphosphine oxide, sold as Cyanex 923 by Cytec Solvay), triisobutylphosphine sulfide (sold as Cyanex 471X by Cytec Solvay), octyl-phenyl-N,N- diisobutyl-carbamoylmethylphosphine oxide (also called CMPO), octyl phenyl acid phosphate (also called OPAP), tributyl phosphate (also called TBP), among others, and mixtures thereof. The extractants may also comprise carboxylic acids, such as pivalic acid (sold as Versatic 5 by Hexion) and neodecanoic acid (sold as Versatic 10 by Hexion), sulfonic acids, such as dinonylnaphtylsulfonic acid (called DINNSA), β-diketones (e.g. sold as LIX 54-100 by BASF), among others, and mixtures thereof. The metal extracting agents or extractants may comprise any of the previously mentioned reagents, among other related compounds (e.g. modified extractants), as well as any mixture or blend thereof.

In some embodiments of the invention, particularly when copper is a target metal, preferably a ketoxime-, aldoxime-, or β-diketone-based extracting agent is used (e.g. a ketoxime, an aldoxime, an aldoxime-ketoxime blend, a modified aldoxime, a hydrophobic β-diketone, a β-diketone-ketoxime blend, etc.), sold e.g. as the LIX series of solvent extraction reagents by BASF or as the ACORGA series of copper extractants by Cytec Solvay.

In other embodiments, particularly when zinc is a target metal, preferably an aldoxime- or phosphoric-acid-based extracting agent is used (e.g. an aldoxime, a modified aldoxime, D2EHPA).

In other embodiments, particularly when both copper and zinc are target metals, preferably ketoxime- and aldoxime-based extracting agents are used. A ketoxime (e.g. LIX 84-I) is a particular good extracting agent for extracting copper and a poor extracting agent for zinc, and is thus preferably used for selective copper extraction. An aldoxime (e.g. LIX 860N-I) is an even stronger extracting agent for copper but is also capable of extracting zinc and therefore less selective for copper than a ketoxime. In one embodiment of the invention, separate solvent extraction circuits can be implemented for copper and zinc, using a ketoxime as extractant in the first circuit to extract the copper first, from the pregnant leach solution (20), and then using an aldoxime as extractant in the second circuit to extract the remaining zinc, from the raffinate (I) solution (30). In another embodiment, a combined (single) solvent extraction circuit can be used to simultaneously co-extract both copper and zinc, using an aldoxime-ketoxime blend (e.g. LIX 984N) as extractant. To optimize the extraction, the blend ratio of aldoxime and ketoxime may be modified for each specific application.

The target metal(s) are preferably extracted from the pregnant leaching solution (20) into the organic solution (35) by mixing the two (immiscible) solutions and, after settling, both solutions are separated by gravity to produce a loaded organic solution (40) with increased content of target metal(s).

The (solvent) extraction step (200) is preferably carried out continuously in rather traditional mixer-settler units. It may involve a single or multiple extraction stages, and may be connected in any configuration (e.g. series, parallel, series-parallel, interlaced, etc.). Some configurations may involve several pregnant leach solutions feeding different extraction stages as well as several raffinates returned to different leach stages or operations. Solvent extraction is preferably carried out under ambient temperature and atmospheric pressure. Gas desorption

The gas desorption (300), sometimes also called gas stripping or gas sparging, involves contacting the loaded organic solution (40) from the extraction step (200) with a desorption gas (45) (i.e. the sorption gas that enters the gas desorption step) to remove and recover ammonia from the organic solution (40), thereby providing a resultant desorbed loaded organic solution (50) with reduced content of entrained ammonia and a resultant enriched gas stream (55) with increased content of ammonia.

Gas desorption can be conducted in a mixer, a settler, a tank, a vessel, a pipe, a packed column (also called a packed tower), a trayed tower (also called a plate column or a tray column), a (circulating) bubble column/tank and many other types of embodiments. Gas desorption is operated preferably in a countercurrent flow manner, by forcing the gas flow and liquid flow in opposite directions. Preferably the organic solution is fed at the top of the desorber and preferably this solution is removed at the bottom of the desorber. The gas is preferable introduced at the bottom of the desorber and removed at the top. In this way a countercurrent flow can be obtained.

In a preferred embodiment gas desorption is performed in a packed column in which random or structured packings are present (e.g. Raschig rings, Pall rings, cross flow structures, etc.). Preferably the solution is fed at the top of the column by means of a liquid distributor, which distributes the solution evenly over the packing. The surface area of the packing material preferably has hydrophobic properties, due to the use of an organic solution. This ensures that the organic solution is evenly distributed over the packing surface and that the gas-liquid contact surface area is maximized.

In another preferred embodiment gas desorption is performed in a plate or tray column. In this type of desorber, the column is divided into one or multiple horizontal sections in which the solution and the gas are in near equilibrium with each other (e.g. the ammonia concentration in the solution is in equilibrium with the concentration in the gas). Multiple plates and/or trays can be used in order to increase the efficiency of the desorber. An example of a tray type, which can be used, is the so called "bubble cap tray", which is often used in desorption columns.

In even another preferred embodiment the gas desorption is conducted in a tank, which may involve mixing the gas and the liquid. Preferably a gas distributor is present at the bottom of the desorber, through which the gas is distributed and blown through the liquid. Mixing can be enhanced by introducing a mixer or by using a circulating bubble column/tank type of desorber, in which the column/tank is divided into three vertical sections. Gas is blown upwards through the middle section, allowing the liquid to circulate over the middle and side sections (i.e. liquid is blown upwards in the middle section and comes downwards again along the side sections).

In some embodiments the gas desorption is performed at a temperature above ambient temperature, a pressure below atmospheric pressure, or both. When operating at a higher temperature, it should not be so high so as to degrade significantly the organic extractant (e.g. preferably not above 65°C for aldoxime or ketoxime) and when operating at lower pressure it should not be so low so as to significantly evaporate the organic solution. When operating at a lower pressure, the lower pressure is preferably applied in the gas stream outlet, i.e. in the exit point of the enriched gas stream (55) from the gas desorption step (300), so as to generate a vacuum that improves desorption efficiency, e.g. using a fan or a vacuum pump. When operating at a temperature above 10°C, this temperature is preferably achieved by heating the entering gas stream (i.e. the desorption gas (45)), the entering liquid stream (i.e. the loaded organic solution (40)), or both. The temperature is preferably in the range of 10°C to 80°C, such as for example between 15°C and 80°C, more preferably in the range of 20°C to 65°C, and the pressure is preferably in the range of 0.01 atm to 1 atm. The preferred, but not required, use of a sorption gas with low oxygen gas content in some preferred embodiments of the invention avoids the risk of possible oxidation (and hence degradation) of any extracting agent in the organic solution. Gas absorption

The gas absorption (600), sometimes also called gas scrubbing or gas washing, involves contacting the enriched gas (55) from the gas desorption step (300) with an aqueous absorbent solution (65) to transfer ammonia and, if present, carbon dioxide from the enriched gas stream (55) to the absorbent solution (65), thereby providing a resultant exhaust gas stream (60) with reduced content of ammonia and, in some cases, of carbon dioxide and a resultant aqueous absorption solution (70) with increased content of ammonia, ammonium ions, or both, as well as, in some cases, carbonate ions.

Gas absorption can be conducted in a mixer, a tank, a (pressurized) vessel, a pipe, a packed column (also called a packed tower), a trayed tower (also called a plate column or a tray column), a (circulating) bubble column/tank, among many other types of absorbers. Gas absorption is operated preferably in a countercurrent flow manner, by forcing the gas flow and liquid flow in opposite directions in the absorber. Preferably the solution is fed at the top of the absorber and preferably the solution is removed at the bottom of the absorber. The gas is preferable introduced at the bottom of the column and removed at the top. In this way a countercurrent flow can be obtained.

Gas absorption can be conducted in the same preferred embodiments as the gas desorption process, as desorption and absorption operate in a similar manner. The main difference corresponds to the direction in which a compound is transferred. In gas absorption a compound is transferred from the gas phase into the liquid phase and in gas desorption a compound is transferred from the liquid phase into the gas phase.

If an aqueous absorbent solution is used, then the surface area of the packing material preferably has hydrophilic properties, and if the absorbent solution is an organic solution, then the surface area of the packing material in the absorber preferably has hydrophobic properties. This allows the liquid to be evenly distributed over the packing surface and maximizes the gas-liquid contact surface area.

In some embodiments the gas absorption is performed at a temperature below ambient temperature, a pressure above atmospheric pressure, or both. This is opposite from the gas desorption, in which a higher temperature, a lower pressure, or both is sometimes preferred, because the transfer of the absorbed compound takes place in the opposite direction than in gas desorption (absorption: gas to liquid, desorption: liquid to gas). When operating at a higher pressure, the higher pressure is preferably applied in the gas stream inlet, i.e. in the entrance point of the enriched gas stream (55) to the gas absorption step (600), so as to generate an overpressure that improves the partial pressure of ammonia and hence the absorption efficiency, e.g. using a compressor. When operating at a lower temperature, the lower temperature is preferably achieved by cooling the entering gas stream (i.e. the enriched gas (55)), the entering liquid stream (i.e. the absorbent solution (65)), or both. The temperature for gas absorption is preferably in the range of 0°C to 30°C, more preferably in the range of 1 °C to 20°C, even more preferably between 1 °C and 15°C; for example between 1 °C and 10°C, while the pressure is preferably in the range of 1 atm to 10 atm.

The efficiency of the absorption of ammonia in the aqueous absorption solution decreases when the pH of the absorption solution is increased. This is because ammonia becomes more volatile if the pH is increased.

The pH of the aqueous absorption solution (70), which is leaving the absorption step (600) and which is recycled to the leaching step (100), is preferred to be equal (or at least similar) to the pH of the leach solution (10). It is most preferred that the pH value of the absorption solution (70), which is leaving the absorption step (600) is lower than the pH of the leach solution (10), in order to compensate for the increase in pH during leaching (e.g. according to the reactions 2Cu + 0 ₂ + 4H ⁺ → 2Cu ²⁺ + 2H ₂ 0 and 2Zn + 0 ₂ + 4H ⁺ → 2Zn ²⁺ + 2H2O). Hydrogen ions (H ⁺ ) are consumed during the absorption of ammonia (according to the reaction NH3 + H ⁺ → NhV), due to which the pH of the absorption solution (70) will increase and due to which the efficiency of the absorption of ammonia will decrease. The pH of the absorption solution (70) is preferably, but not necessarily, adjusted by adding a certain amount of carbon dioxide to the sorption gas (e.g. by adding pure carbon dioxide or a combustion gas). The carbon dioxide will dissolve in the absorption solution (70), which decreases the pH of the solution, because H ⁺ ions are produced (according to the reaction CO2 + H2O→2H ⁺ + CO3 ^2" ). For these reasons the pH of the aqueous absorption solution (70), which is leaving the absorption step (600) is preferably not above 10. Steam stripping and condensation

Steam stripping (1 100) involves contacting the absorption solution (70) from the gas absorption step (600) with steam (96), thereby providing a resultant enriched steam stream (97) with increased content of ammonia and a resultant depleted solution with reduced content of ammonia, which can be returned back to the gas absorption step (600) as the absorbent solution (65). Condensation (1200) involves changing the enriched steam (97) from the steam stripping step (1 100) from gas phase to liquid phase by cooling and/or compression, thereby providing a condensed enriched solution (98) with higher ammonia concentration than the absorption solution (70).

Steam stripping can be conducted in a mixer, a tank, a (pressurized) vessel, a pipe, a packed column (also called a packed tower), a trayed tower (also called a plate column or a tray column), a (circulating) bubble column/tank and many other types of strippers. Steam stripping is operated preferably in a countercurrent flow manner, by forcing the steam flow and liquid flow in opposite directions in the stripper. Preferably the liquid is fed at the top of the absorber and preferably the liquid is removed at the bottom of the absorber. The steam is preferable introduced at the bottom of the column and removed at the top. In this way a countercurrent flow can be obtained.

Steam stripping is performed in the same preferred embodiments as the gas desorption process since steam stripping is a desorption process in which steam is used as the desorption gas.

Steam stripping is generally operated at elevated temperatures (above ambient temperature). Elevated temperatures are needed in order to keep the steam in the gas phase (above 100°C at 1 atm). With respect to pressure, either pressures below atmospheric pressure or pressures above atmospheric pressure can be used. Pressures below atmospheric pressure can be used in order to decrease the temperature that is needed to maintain steam in the gas phase (e.g. 45°C at 0.1 atm). This, for example, is preferred if the absorption solution degrades or evaporates at elevated temperatures. Pressures above atmospheric pressure can, for example, be used in order to add more energy/steam into the stripper, without changing the temperature of the steam between the inlet and the outlet of the stripper too much (e.g. if a large amount of liquid is fed into the stripper, which needs to be heated by the steam) or if the temperature of the steam is not allowed to be too high (e.g. due to degradation of the liquid), but if still lots of energy/steam is needed inside the stripper. Preferably, the temperature is in the range of 40°C to 200°C and the pressure is in the range of 0.1 atm to 5 atm.

In a condenser the gas stream is cooled down or pressurized, so that condensation occurs. Condensation can be conducted in a shell and tube heat exchanger, a packed column (also called a packed tower), a trayed tower (also called a plate column or a tray column), or in any other type of equipment containing an internal structure with a lot of surface area on which the gas phase can condense. In case of a packed column, the gas phase can be condensed by contacting and cooling it with a liquid flow.

Steam stripping as well as condensation may involve a single stage, but also multiple stages, which can be present in the same embodiment (e.g. multiple plates or trays in a column). Multiple stages are needed if the remaining liquid phase (in the case of steam stripping) or if the remaining gas phase (in the case of condensation) still contains too much of the compound, which need to be stripped/condensed. Stripping

The stripping (400) of target metal(s) involves contacting the (desorbed) loaded organic solution (50) with a spent electrolyte (or stripping) solution (75) to strip the metal ions from the loaded organic solution (50) into the spent electrolyte (75), thereby providing a resultant pregnant electrolyte solution (80) with increased concentration of metal ions and a resultant stripped organic solution (35) with reduced content of metal ions.

The target metal(s) are preferably stripped from the loaded organic solution (50) into the spent electrolyte (75) by mixing the two (immiscible) solutions and, after settling, both solutions are separated by gravity to produce a pregnant electrolyte (80) with increased content of target metal(s).

The stripping step (400), when not involving a simultaneous crystallization (to produce metal salts crystals), is preferably carried out continuously in rather traditional mixer-settler units. It may involve a single or multiple stripping stage(s), preferably connected in a series configuration, although other configurations may also be used if required.

Stripping is preferably carried out under ambient temperature and atmospheric pressure. In some embodiments, to improve stripping kinetics, the temperature may be increased, preferably not too high to avoid excessive degradation of the extractant (e.g. preferably not above 60°C for aldoxime or ketoxime). However, with increased temperature the stripping equilibrium isotherm tends to deteriorate (e.g. for aldoxime- or ketoxime-based extractants).

Organic wash

An (organic) wash step, e.g. (700), involves the use of an aqueous wash solution for the removal of entrained impurities, leaching agents, or both from an organic solution, typically from the loaded organic solution (50) prior to the stripping step (400), but it may also be used for the stripped organic solution (35) prior to the extraction step (200), or for other organic solutions (e.g. between stripping steps for different target metals).

Wash steps are preferably carried out continuously in rather traditional mixer- settler units. They may involve a single or multiple wash stages. The organic wash is preferably carried out under ambient temperature and atmospheric pressure.

Metal recovery

Metal recovery (500) involves any process to produce a desired final metal product (85) from the pregnant electrolyte solution (80) from the stripping step (400).

Metal recovery may involve electrowinning, precipitation (crystallization or cementation), smelting, or combinations thereof, among others. Precipitation by crystallization may involve cooling, evaporation, change of pH, change of cations or anions, addition of a second solvent to reduce the solubility of the solute, among several other methods.

In preferred embodiments of the invention, the metal recovery step (500) involves electrowinning (to produce metal cathodes), crystallization (to produce metal salts), or just solution storage (to produce a pregnant metal solution).

Stripping 8$ crystallization

The simultaneous stripping and crystallization (400) of target metal(s) is achieved by contacting the desorbed loaded organic solution (40) with an acidic aqueous stripping solution (90) that is saturated with dissolved target metal(s), e.g. 20-70 g/L Cu or 150-250 g/L Zn, so as to strip one or more target metal(s) from the desorbed loaded organic solution (40) and precipitate metal salt crystals. In some embodiments of the invention, the acidic stripping solution may not be saturated at first (e.g. if it was previously subjected to electrowinning in order to produce another metal product next to the metal salt crystals), but will become saturated (and then oversatu rated) as the target metal(s) are stripped from the desorbed loaded organic (50).

For example, in the case of copper sulfate, zinc sulfate, nickel sulfate, or cobalt sulfate production, the acidic stripping solution (90) preferably comprises sulfuric acid.

The stripping and crystallization step (400) is preferably carried out in a rather traditional mixer-settler unit or in a mixer-thickener unit. The mixer is preferably located at a higher position so that the metal salt crystals produced in the emulsion during mixing may settle down in the settler or thickener following after the mixer. The mixer may involve one or more mixing tanks, e.g. a primary mixer and an auxiliary or secondary mixer. Typically, as the mixers often require maintenance, particularly due to crystals sticking on the mixing impeller, one or more spare mixers are desirable in the design, to allow operating continuously even during maintenance of one of the mixers.

The settler after the mixer may be of traditional design or it may involve a certain slope or geometry at its bottom to allow an easy recollection of the metal salt crystals. In a preferred embodiment, the settler corresponds rather to a traditional thickener, allowing the metal salt slurry (31 ) to exit from the bottom of the thickener by an underflow and the solutions, i.e. the saturated stripping solution (90) and the stripped organic solution (35), to exit from the top by appropriate overflow weirs, as in a traditional settler, to be collected respectively by an aqueous launder and an organic launder.

Stripping and crystallization is preferably carried out under ambient temperature and atmospheric pressure. In some embodiments, to improve crystallization kinetics and/or the stripping equilibrium isotherm, the temperature may be lowered below ambient temperature. Organic recovery

During stripping and crystallization (400), when the metal salt crystals leave the emulsion by gravity and settle at the bottom, some organic solution is entrained in the settled metal salt slurry (31). After separating the metal salt slurry (31) from the saturated stripping solution (90) and the stripped organic (35), an organic recovery step (800) may optionally be carried out, whereby entrained organic solution (62) is separated and recovered so as to yield a resultant metal salt slurry (61) with reduced content of entrained organic solution.

The recovery of entrained organic solution is performed by some means of mechanical agitation, e.g. by blowing air through the settled crystals and/or by shaking them with a rake (for example, in a thickener), allowing the rise of organic solution droplets, which are then recovered from the top (for example, by an overflow). In a preferred embodiment, the entrained organic solution is recovered in a thickener by means of a rotating rake and/or by means of blowing air through the settled crystals. The organic recovery (800) may be even implemented in the same thickener or settler of the mixer-thickener or mixer-settler unit of the stripping and crystallization step (400). The organic recovery (800) may also be implemented in tanks or similar equipment by introducing air and/or with a suitable mechanical device. The recovered entrained organic solution (62) may be added to the circulating organic solution, e.g. the loaded organic (40), the desorbed loaded organic (50), or the stripped organic (35), it may be sent to organic treatment (e.g. by using clay), or it may be disposed of for other uses.

Solid/liquid separation & wash

The resultant metal salt slurry (61) from the organic recovery step (800), or in its absence the metal salt slurry (31) from stripping & crystallization (400), is subject to a solid/liquid separation S wash step (900). The solid/liquid separation involves removing the liquids from the slurry and retaining the solids, i.e. the metal salt crystals, which are washed thereafter with water (71) and the solids are then again separated from the liquids, so as to obtain a reasonable dry cake. In some embodiments of the invention, an acid wash is implemented prior to the water wash, i.e. the metal salt crystals are washed with an acid solution (73) to minimize the precipitation of impurities and allow them to be better washed off thereafter with water (71), prior to a solid/liquid separation in between (as an alternative, the water may wash off directly the acid solution without a solid/liquid separation in between). The liquids, i.e. the wash solution (72), obtained from the solid/liquid separation & wash (900), are preferably recycled back to the saturated stripping solution (90), but may also be sent to other process steps or to waste disposal. The amount of acid solution added to (acid) wash the metal salt crystals is preferably also considered in the water balance so that together with the wash water (71) they should compensate the amount of water lost as crystallization water entrapped within the metal salt crystals (84) and the water withdrawn as the bleed solution (33).

The solid/liquid separation and the involved wash operations may be implemented by filtration, centrifugation or similar solid/liquid separation processes, e.g. using equipment such as a belt filter, a vacuum filter, a filter press, a screen filter, a disc filter, a plaque filter, a centrifuge, or other similar solid/liquid separation equipment. The wash operation is preferably performed within the solid/liquid separation equipment, but may also be performed by using separate wash equipment or in a separate wash stage. Preferably a centrifuge, either continuous or batch, is used for solid/liquid separation and wash. The acid wash, if desired, may also be performed within the centrifuge. In other embodiments, the solid/liquid separation may be performed by just draining the liquids from the metals salt crystals.

Drying

After the solid/liquid separation & wash step (900), the metal salt crystals (84) are sent to a drying step (1000) wherein the metal salt crystals are dried to remove residual water moisture.

The drying of the metal salt crystals may be performed in a kiln, an oven, or similar heating equipment. Drying may even involve solar sun-drying. Preferably, drying is performed in a rotary kiln. Preferably, a temperature between 20°C and 100°C is used for drying. In some embodiments higher temperatures may be used, e.g. above 250°C when the crystallization water entrapped within the metal salt crystals is also desired to be removed. For the purpose of clarity and a concise description, features are described herein as part of the same or separate embodiments, however, it will be appreciated that the scope of the invention may include embodiments having combinations of all or some of the features described.

It is to be understood that in this invention the preferred embodiments, conditions, flow patterns, materials, metals, and reagents are not limited to those particular mentioned, as these may vary. It is also to be understood that the terminology used herein is for the purpose of describing particular embodiments only, and is not intended to limit the scope of the present invention in any way. EXAMPLES

The invention may be illustrated with the following examples. In the examples, concentration units for liquids are expressed as mass concentration (10000 ppm equal 1 w/v% equal 10 g/L), whereas concentration units for gases are expressed as volume concentration (10000 ppm equal 1 v/v% equal a volume fraction of 0.01).

Example 1

A setup of two columns was arranged of 20 cm effective solution height with 5 cm inside diameter each. In the first column 0.5 liters of copper and zinc loaded organic with 30 v/v% LIX 84-I and 70 v/v% Shellsol D100 were stirred by introducing 2.5 L/min of air at the bottom of the organic column. LIX 84-I from BASF, is water insoluble 2-hydroxy- 5-nonylacetophenone oxime in an hydrocarbon diluent, which forms water insoluble complexes with various metallic cations, such as copper. The initial organic contained 412.6 ppm of equivalent NH3 and the enriched gas from the top of the organic column was conducted with a hose to the second column and introduced at its bottom. In the second column 0.5 liters of aqueous were stirred with the enriched gas flow, having started with pure water. The final exhaust gas was evacuated to the atmosphere. The purpose of the batch test was to desorb ammonia from the organic and to absorb ammonia into the aqueous. The results of the batch test are shown on Table 1.

Table 1

DesorpEnriched

Item Units Organic Aqueous

tion gas gas

Flow L/h 0 150 0 150

Test volume L 0.5 0.5

Initial equiv. NH3 cone. ppm 412.6 0 0 1690

Final equiv. NH3 cone. ppm 27.1 1690 350.7 152

Equiv. CO2 cone. ppm -400

NH3 desorption from

% 93.4

organic

NH3 absorption in aqueous % 91 .0

Example 2

By using the same setup as in the previous example, as a desorption gas a 50/50 volume blend of air and CO2 gas was introduced at a flow of 2.5 L/min to the organic column at its bottom containing the same type of organic with 450.6 ppm of equivalent NH3. The results of the batch test are shown on Table 2.

Example 3

By using the same setup as in the previous examples, as desorption gas pure CO2 gas was introduced to the organic column at its bottom with a flow of 2.5 L/min, containing the same type of organic with 467.2 ppm of equivalent NH3. The results of the batch test are shown on Table 3.

Table 3

Enric

Desorp¬

Item Units Organic Aqueous hed

tion gas

gas

Flow L/h 0 150 0 150

Test volume L 0.5 0.5

Initial equiv. NH ₃ cone. ppm 467.2 0 0 1970

Final equiv. NH ₃ cone. ppm 18.8 1970 267.4 796

Equiv. C0 ₂ cone. ppm -1000000

NH ₃ desorption from

% 96.0

organic

NH ₃ absorption in

% 59.6

aqueous

Example 4

By using the same setup as before, both columns were filled with equivalent standard packing material. The copper/zinc loaded organic of 30 v/v% LIX 84-I in Shellsol D100 with 383.0 ppm of equivalent Nh was added at the top of the organic column at 3.6 L/h in countercurrent manner with a gas blend flow of 60 L/h and ~2 v/v% CO2 in air. The enriched gas from the top of the organic column was transferred with a hose to the bottom of the aqueous column, being introduced in countercurrent flow with 0.77 L/h of pure water. The results of the batch test are shown on Table 4.

Table 4

DesorpEnriche

Item Units Organic Aqueous

tion gas d gas

Flow L/h 3.6 60 0.77 60

Test volume L 3.6 0.77

Initial equiv. NH ₃

ppm 383.0 0 0 4.030 cone.

Final equiv. NH ₃

ppm 332.0 4.030 244.0 0 cone.

Equiv. C0 ₂ cone. ppm -20000

NH ₃ desorption

% 13.3

from organic

NH ₃ absorption in

% 100.0

aqueous

Example 5

By using the same setup of Example 4, the organic contained 552.6 ppm of equivalent NH3. The organic flow was kept at 3.6 L/h, changing the gas flow of ~2 v/v% CO2 in air to 288 L/hour and the aqueous flow to 0.85 L/h. The results of the batch test are shown on Table 5.

Table 5

DesorpEnriched

Item Units Organic Aqueous

tion gas gas

Flow L/h 3.6 288 0.85 288

Test volume L 3.6 0.85

Initial equiv. NH ₃ cone. ppm 552.6 0 0 3590

Final equiv. NH ₃ cone. ppm 334.4 3590 553.2 1440

Equiv. C0 ₂ cone. ppm -20000

NH ₃ desorption from

% 39.5

organic

NH ₃ absorption in

% 59.9

aqueous Example 6

Bench scale leach experiments where performed on samples sizes of 3 kg of zinc- and copper-containing material with particle sizes ranging between 0-4 mm. A rotating drum leaching reactor was used in which >90% of the copper and zinc was dissolved within 3 days

Example 7

Bench scale leach experiments where performed on samples sizes of 16.2 kg of zinc- and copper-containing material with particle sizes ranging between 0-12 mm. A column-type leaching reactor was used in which >80% of the copper and zinc dissolved within 10 days.

Example 8

Bench scale leach experiments where performed on samples sizes of 0.05 kg of zinc- and copper-containing material with particle sizes ranging between 0.25-0.5 mm and 0.5-1 mm. A stirred tank leaching reactor was used in which >90% of the copper and zinc dissolved within 2 hours.

Example 9

Tests were performed to produce zinc sulfate crystals out of zinc- and copper- containing material by making use of the described stripping and simultaneous crystallization process. Zinc sulfate heptahydrate, containing 0.1 w% of impurities, was successfully crystallized.