EA045218B1 - BATTERY RECYCLING METHOD - Google Patents

Description

Field of technology

The present invention relates to a method for recovering metals from raw materials containing such metals. More specifically, the present invention relates to the recovery of various metals contained in used lithium (lithium-ion) batteries.

State of the art

The following discussion of the prior art is provided solely to facilitate understanding of the present invention. Consideration does not constitute a confirmation or admission that any material referred to is or was part of the generally known knowledge at the priority date of the application.

The global use of rechargeable lithium-ion batteries has been growing rapidly in recent years and will continue to do so as markets for electric vehicles and mass energy storage develop. As demand for lithium-ion batteries increases, so does the demand for metal/metal oxide components used in batteries. The rapid increase in demand for some of these metals, such as cobalt, has called into question the stable supply of such resources. This led to a rapid increase in the cost of such metals.

There has been little interest in developing ways to extract and recycle various components of modern batteries. This is primarily due to the relatively small volume of lithium-ion batteries available for recycling and the relatively high cost of the typical pyro- and hydrometallurgical processes by which recovery is achieved. As demand for lithium-ion batteries continues to grow, so does the volume of used lithium-ion batteries that can be recycled. There is a need for low-cost and efficient recycling processes, especially for more complex metal/metal oxide components.

The composition of lithium-ion batteries has changed significantly recently. While some battery recycling processes have been developed, they have largely been limited to recovering specific metals from a specific battery type or raw material source. For example, previous batteries were predominantly lithium-cobalt, and the extraction methods focused on cobalt recovery. As demand for lithium has increased, extraction methods have shifted to extracting both cobalt and lithium. As battery technology further developed, cathodes included other metals such as manganese, nickel, aluminum, iron, and phosphorus. The methods used to extract lithium and cobalt are not suitable for extracting other metals and are not suitable for other battery chemistries.

Increasing use of lithium-ion batteries will increase the volume of used lithium-ion batteries available for recycling. However, the supply of used lithium-ion batteries will include many different types of batteries. The suitability of recovery methods for only one type of battery poses a significant challenge to the commercialization of such processes, as it will require sorting and separation. There is a need to develop a process to extract a range of metals from a range of different types of lithium-ion batteries.

It is with this objective in mind that the present invention has been developed.

The present invention is directed to overcome or at least mitigate one or more of the disadvantages of the prior art mentioned above, or to provide the consumer with a practical or commercial choice.

As used herein, unless the context otherwise requires, the word contain or variants such as contains or containing are to be understood to mean the inclusion of the specified integer or group of integers, but not the exclusion of any other integer or group of integers.

The essence of the invention

In accordance with a first aspect of the present invention, there is provided a method for recovering metals from a feed stream containing one or more valuable metals and lithium, the method comprising:

subjecting the raw material stream to sulfuric acid leaching to form a suspension containing a productive leaching solution of soluble metal salts and a solid residue;

separation of productive leach solution and solid residue;

subjecting the productive leach solution to one or more separate solvent extraction steps, each solvent extraction step removing one or more valuable metals from the productive leach solution, the remaining productive leach solution containing lithium; and recovering lithium from the productive leach solution.

The method of the present invention is particularly suitable for recovering at least a substantial portion of all valuable metals from used lithium-ion batteries, preferably in the form of high purity sulfates. This process is particularly robust because it allows lithium-ion battery chemistries to be used as a single or mixed source. This is achieved by acid leaching followed by successive steps of metal recovery using solvent extraction methods. This has proven to be particularly advantageous as the process can process a range of different battery chemistries, meaning that different battery types are not required to be sorted.

In preferred embodiments of the present invention, the feed stream contains one or more of copper, iron, manganese, aluminum, cobalt and nickel.

In one embodiment of the present invention, the method further includes the following step:

subjecting the feed stream to a pre-treatment process prior to the step of subjecting the feed stream to sulfuric acid leaching.

Preferably, the pre-treatment process includes one or more mechanical processing steps. More preferably, the mechanical processing steps include one or more of a crushing step and a grinding step.

Preferably, the pre-treatment process includes one or more size reduction steps. More preferably, one or more size reduction steps include a grinding step.

In one embodiment of the present invention, the pretreatment process includes one or more enrichment steps. Preferably, the one or more enrichment steps include one or more of an air classification step, a magnetic classification step, and a flotation step.

In one embodiment of the present invention, the steps of subjecting the feed stream to sulfuric acid leaching to form a slurry containing a productive leaching solution of soluble metal salts and a solid residue; and separation of the productive leach solution and the solid residue;

carried out in the leaching circuit.

Preferably, the leaching circuit comprises a leaching section, a thickening section and a filtration section.

In one embodiment of the present invention, the step of subjecting the feed stream to sulfuric acid leach to form a slurry containing a productive leach solution of soluble metal salts and a solid residue, more specifically involves subjecting the feed stream to sulfuric acid leach in one or more leach reactors. Preferably, the step includes subjecting the feed stream to sulfuric acid leaching in two or more leaching reactors. More preferably, the step includes subjecting the feed stream to sulfuric acid leaching in three or more leaching reactors. More preferably, the step includes subjecting the feed stream to sulfuric acid leaching in four or more leaching reactors. More preferably, the step includes subjecting the feed stream to sulfuric acid leaching in five or more leaching reactors.

Preferably, the step of subjecting the feed stream to sulfuric acid leaching is carried out at atmospheric pressure.

Preferably, the step of subjecting the feed stream to sulfuric acid leaching is carried out at an elevated temperature.

In one embodiment of the present invention, each solvent extraction step includes contacting a productive leach solution with an extractant to extract one or more metals to produce a saturated extractant containing one or more extracted metals. Preferably, each solvent extraction step further includes separating the saturated extractant from the productive leach solution. More preferably, each solvent extraction step further includes recovering the metal from the rich extractant.

In one embodiment of the present invention, one of the separate solvent extraction steps is designed to recover copper from the productive leach solution. Preferably, the productive leach solution is contacted with a copper extractant to produce a copper-depleted productive leach solution and a rich copper extractant. More preferably, the copper is recovered from the saturated copper extractant. Even more preferably, the copper is recovered as sulfate.

In one embodiment of the present invention, one of the separate solvent extraction steps is designed to recover aluminum from the productive leach solution. Preferably, the productive leach solution is contacted with an aluminum extractant to produce an aluminum-depleted productive leach solution and a saturated aluminum extractant. More preferably, the aluminum is recovered from the saturated aluminum extractant. In one embodiment of the present invention, aluminum is recovered from the productive leach solution after copper has been recovered.

In one embodiment of the present invention, contacting the productive leach solution with an aluminum extractant provides additional iron saturation from the productive leach solution to the saturated aluminum extractant. Preferably, the iron is extracted from

- 2 045218 saturated aluminum extractant.

In an alternative embodiment of the present invention, aluminum is recovered from the productive leach solution through a precipitation step. Preferably, if the aluminum is recovered through the precipitation step, the iron is also recovered through the precipitation step. Preferably, the aluminum and iron are recovered as a mixed hydroxide precipitate. More preferably, the precipitation step further includes separating the precipitated solids from the productive leach solution.

In one embodiment of the present invention, the precipitation step includes adding an oxidizing agent. Preferably, the oxidizing agent is selected from air, oxygen gas, oxygen-enriched gases and hydrogen peroxide.

In one embodiment of the present invention, the precipitation step includes increasing the pH of the productive leach solution to a level greater than 4.5. Preferably, the pH is increased by adding a base. Preferably the base is selected from ammonia, lime or limestone. In one embodiment of the present invention, the precipitation step includes two or more precipitation stages operating at increased pH. Preferably, the at least one precipitation stage operates at a pH between 3.2 and 4 and the at least one precipitation stage operates at a pH above 4.5. More preferably, an oxidizing agent is added to one or both precipitation steps.

In one embodiment of the present invention, one of the separate solvent extraction steps is designed to recover manganese from the productive leach solution. Preferably, the productive leach solution is contacted with a manganese extractant to produce a manganese-depleted productive leach solution and a rich manganese extractant. More preferably, the manganese is recovered from the saturated manganese extractant. Even more preferably, the manganese is recovered as sulfate. In one embodiment of the present invention, manganese is recovered from the productive leach solution after aluminum and/or iron have been recovered.

In one embodiment of the present invention, one of the separate solvent extraction steps is designed to recover cobalt from the productive leach solution. Preferably, the productive leach solution is contacted with a cobalt extractant to produce a cobalt-depleted productive leach solution and a rich cobalt extractant. More preferably, the cobalt is recovered from the saturated cobalt extractant. Even more preferably, the cobalt is recovered as sulfate. In one embodiment of the present invention, cobalt is recovered from the productive leach solution after manganese has been recovered.

In one embodiment of the present invention, one of the separate solvent extraction steps is designed to recover cobalt and manganese from the productive leach solution. The inventors envision that in this embodiment of the present invention, cobalt and manganese are simultaneously extracted from the productive leach solution using a single extractant. Preferably, the productive leach solution is contacted with a cobalt and manganese extractant to produce a cobalt and manganese depleted productive leach solution and a saturated cobalt and manganese extractant. More preferably, the cobalt and manganese are recovered from the saturated cobalt and manganese extractant. Even more preferably, the manganese and cobalt are recovered as sulfate. In one embodiment of the present invention, cobalt and manganese are recovered from the productive leach solution after aluminum and/or iron have been recovered.

In one embodiment of the present invention, one of the separate solvent extraction steps is designed to recover nickel from the productive leach solution. Preferably, the productive leach solution is contacted with a nickel extractant to produce a nickel-depleted productive leach solution and a rich nickel extractant. More preferably, the nickel is recovered from the saturated nickel extractant. Even more preferably, the nickel is recovered as sulfate. In one embodiment of the present invention, nickel is recovered from the productive leach solution after the manganese and/or cobalt have been recovered.

In one embodiment of the present invention, the individual solvent extraction steps are performed sequentially. Preferably, the individual solvent extraction steps are performed in the following sequence: copper, aluminum, manganese, cobalt and nickel. If manganese and cobalt are extracted together, the individual solvent extraction steps are performed in the following sequence: copper, aluminum, manganese/cobalt and nickel.

In one embodiment of the present invention, after one or more separate solvent extraction steps, the productive leach solution contains lithium.

In one embodiment of the present invention, recovery of lithium from a productive leach solution more specifically involves precipitation of a lithium-containing compound. More preferably, the lithium-containing compound is then removed from solution. In one embodiment, the lithium-containing compound is lithium carbonate. Preferably, the productive leach solution is contacted with ammonium carbonate to precipitate lithium carbonate. Alternatively, the productive leach solution is contacted with sodium carbonate to precipitate lithium carbonate.

In an alternative embodiment of the present invention, lithium-containing compounds are precipitated by crystallization. Preferably, lithium sulfate is crystallized from solution by

- 3 045218 evaporation.

In one embodiment of the present invention, recovering lithium from a productive leach solution more specifically includes recovering lithium by solvent extraction.

In one embodiment of the present invention, recovery of lithium from a productive leach solution more specifically includes recovery of lithium by ion exchange.

In one embodiment of the present invention, recovery of lithium from the productive leach solution more specifically involves the use of an electrolysis step. Preferably, the electrolysis step produces a LiOH solution.

In one embodiment of the present invention, after the lithium recovery step, the leach solution is subjected to an ammonia recovery step. Preferably, the ammonia recovery step includes a crystallization step of ammonium sulfate.

In accordance with another aspect of the present invention, there is provided a method for recovering metals from a feed stream containing one or more valuable metals and lithium, the method comprising:

subjecting the raw material stream to sulfuric acid leaching to form a suspension containing a productive leaching solution of soluble metal salts and a solid residue;

separation of productive leach solution and solid residue;

subjecting the productive leach solution to a solvent extraction step to extract copper from the productive leach solution;

subjecting the productive leach solution to a solvent extraction step to extract aluminum from the productive leach solution;

subjecting the productive leach solution to a solvent extraction step to extract manganese from the productive leach solution;

subjecting the productive leach solution to a solvent extraction step to extract cobalt from the productive leach solution;

subjecting the productive leach solution to a solvent extraction step to extract nickel from the productive leach solution; and recovering lithium from the productive leach solution.

Preferably, iron is also recovered in a solvent extraction step designed to recover aluminum from the productive leach solution.

Preferably, each solvent extraction step includes contacting the productive leach solution with an extractant to extract one or more metals to produce a saturated extractant containing one or more extracted metals. More preferably, each solvent extraction step further includes separating the saturated extractant from the productive leach solution. Even more preferably, each solvent extraction step further includes recovering the metal from the rich extractant.

In accordance with another aspect of the present invention, there is provided a method for recovering metals from a feed stream containing one or more valuable metals and lithium, the method comprising:

subjecting the raw material stream to sulfuric acid leaching to form a suspension containing a productive leaching solution of soluble metal salts and a solid residue;

separation of productive leach solution and solid residue;

subjecting the productive leach solution to a solvent extraction step to extract copper from the productive leach solution;

subjecting the productive leach solution to a solvent extraction step to extract aluminum from the productive leach solution;

subjecting the productive leach solution to a solvent extraction step to extract manganese and cobalt from the productive leach solution;

subjecting the productive leach solution to a solvent extraction step to extract nickel from the productive leach solution; and recovering lithium from the productive leach solution.

Preferably, iron is also recovered in a solvent extraction step designed to recover aluminum from the productive leach solution.

Preferably, each solvent extraction step includes contacting the productive leach solution with an extractant to extract one or more metals to produce a saturated extractant containing one or more extracted metals. More preferably, each solvent extraction step further includes separating the saturated extractant from the productive leach solution. Even more preferably, each solvent extraction step further includes recovering the metal from the rich extractant.

Preferably, the product of the solvent extraction step for removing manganese and cobalt from the productive leach solution is subjected to a process to separate the manganese and cobalt. More preferably, the product of a solvent extraction step for extracting manganese and cobalt from the productive leach solution is subjected to a solvent extraction step for extracting manganese. Even more preferably, the cobalt is recovered from the raffinate of the solvent extraction step to recover the manganese.

In accordance with another aspect of the present invention, there is provided a method for recovering metals from a feed stream containing one or more valuable metals and lithium, the method comprising:

subjecting the raw material stream to sulfuric acid leaching to form a suspension containing a productive leaching solution of soluble metal salts and a solid residue;

separation of productive leach solution and solid residue;

subjecting the productive leach solution to a solvent extraction step to extract copper from the productive leach solution;

subjecting the productive leach solution to a precipitation step designed to extract aluminum from the productive leach solution;

subjecting the productive leach solution to a solvent extraction step to extract manganese from the productive leach solution;

subjecting the productive leach solution to a solvent extraction step to extract cobalt from the productive leach solution;

subjecting the productive leach solution to a solvent extraction step to extract nickel from the productive leach solution; and recovering lithium from the productive leach solution.

Preferably, the precipitation step for removing aluminum from the productive leach solution also removes iron from the productive leach solution.

Preferably, each solvent extraction step includes contacting the productive leach solution with an extractant to extract one or more metals to produce a saturated extractant containing one or more extracted metals. More preferably, each solvent extraction step further includes separating the saturated extractant from the productive leach solution. Even more preferably, each solvent extraction step further includes recovering the metal from the rich extractant.

In accordance with another aspect of the present invention, there is provided a method for recovering metals from a feed stream containing one or more valuable metals and lithium, the method comprising:

subjecting the raw material stream to sulfuric acid leaching to form a suspension containing a productive leaching solution of soluble metal salts and a solid residue;

separation of productive leach solution and solid residue;

subjecting the productive leach solution to a solvent extraction step to extract copper from the productive leach solution;

subjecting the productive leach solution to a precipitation step designed to extract aluminum from the productive leach solution;

subjecting the productive leach solution to a solvent extraction step to extract manganese and cobalt from the productive leach solution;

subjecting the productive leach solution to a solvent extraction step to extract nickel from the productive leach solution; and recovering lithium from the productive leach solution.

Preferably, the precipitation step for removing aluminum from the productive leach solution also removes iron from the productive leach solution.

Preferably, each solvent extraction step includes contacting the productive leach solution with an extractant to extract one or more metals to produce a saturated extractant containing one or more extracted metals. More preferably, each solvent extraction step further includes separating the saturated extractant from the productive leach solution. Even more preferably, each solvent extraction step further includes recovering the metal from the rich extractant.

Preferably, the product of the solvent extraction step for removing manganese and cobalt from the productive leach solution is subjected to a process to separate the manganese and cobalt. More preferably, the product of a solvent extraction step for extracting manganese and cobalt from the productive leach solution is subjected to a solvent extraction step for extracting manganese. Even more preferably, the cobalt is recovered from the raffinate of the solvent extraction step to recover the manganese.

Brief description of graphic materials

The features of the present invention are further described more fully in the following description of several non-limiting embodiments thereof. The present description is added solely for the purpose of illustrating the present invention. It should not be construed as limiting the general content, disclosure, or description of the invention set forth above. The description will be given with links to the attached graphic materials, on which:

in fig. 1 shows a flow diagram of the method according to the present invention;

- 5 045218 in Fig. 2 shows a technological diagram of the leaching circuit;

in fig. Figure 3 shows a flow diagram of the copper solvent extraction circuit;

in fig. 4 shows a flow diagram of the aluminum solvent extraction circuit;

in fig. 4a shows a flow diagram of an alternative aluminum extraction circuit;

in fig. 5 shows a flow diagram of the cobalt/manganese extraction circuit;

in fig. 6 shows a flow diagram of the cobalt and manganese solvent extraction circuit;

in fig. 7 shows a flow diagram of the manganese solvent extraction circuit;

in fig. 7a shows a flow diagram of an alternative solvent extraction circuit for manganese and cobalt;

in fig. 8 shows a flow diagram of the nickel solvent extraction circuit;

in fig. 9 shows a flow diagram of the lithium extraction circuit;

in fig. 10 shows a flow diagram of the ammonium sulfate extraction circuit;

in fig. 11 is a graph showing the extraction distribution isotherm for copper;

in fig. 12 is a graph showing the crystallization of copper sulfate;

in fig. 13 is a graph showing the crystallization of copper sulfate;

in fig. 14 is a graph showing the effect of pH on metal deposition at 45°C;

in fig. 15 is a graph showing the effect of pH on metal deposition at 65°C;

in fig. 16 is a graph showing the effect of pH on metal deposition at 90°C;

in fig. 17 is a graph showing the effect of temperature on the separation of aluminum from cobalt;

in fig. 18 is a graph showing the effect of temperature on the separation of aluminum from nickel;

in fig. 19 is a graph showing the pH isotherm at 15% (v/v) D2EHPA at an O/A ratio of 1:1;

in fig. 20 is a graph showing the pH isotherm at 30% (v/v) Cyanex 272 at an O/A ratio of 3:1;

in fig. 21 is a graph showing the extraction distribution isotherm for cobalt at 30% (v/v) Cyanex 272 at pH 4.80;

in fig. 22 is a graph showing the extraction distribution isotherm for manganese at 30% (v/v) Cyanex 272 at pH 4.80;

in fig. 23 shows the pH isotherm at 10% (v/v) LIX84I at an O/A ratio of 1:1; and in fig. Figure 24 is a graph showing copper in the raffinate and extraction over cycle time.

Description of Embodiments

As described above, in one aspect of the present invention there is provided a method for recovering metals from a feed stream containing one or more valuable metals and lithium, the method comprising:

subjecting the raw material stream to sulfuric acid leaching to form a suspension containing a productive leaching solution of soluble metal salts and a solid residue;

separation of productive leach solution and solid residue;

subjecting the productive leach solution to one or more separate solvent extraction steps, each solvent extraction step comprising contacting the productive leach solution with an extractant to extract one or more metals, separating the rich extractant from the productive leach solution, and recovering the metal from the rich extractant; and recovering lithium from the productive leach solution.

In fig. 1 illustrates a method for recovering metals 10 in accordance with one embodiment of the present invention. In this embodiment, the feed stream 12 is subjected to a pre-treatment process 14 to prepare the feed stream for further processing. The resulting feed stream 16 is then fed to leach circuit 200 where it is contacted with sulfuric acid 18 to increase the solubility of the metal components. Undissolved solids 20 are removed from the leach solution to produce productive leach solution 22.

The productive leach solution 22 is sent to a copper solvent extraction circuit 300 where it is brought into contact with a copper extractant 26 . The copper extractant is saturated with copper, and the saturated extractant 28 is separated from the productive leach solution 22. The copper product is then recovered from the copper extractant.

The copper-depleted productive leach solution 30 is sent to an aluminum solvent extraction circuit 400 where it is brought into contact with an aluminum extractant 34 . The aluminum extractant is saturated with aluminum, and the saturated extractant 36 is separated from the productive leach solution 30. The aluminum product is then recovered from the aluminum extractant.

- 6 045218

The aluminum-depleted productive leach solution 38 is sent to the cobalt/manganese recovery circuit 500, where it is contacted with a cobalt and manganese extractant 42. The cobalt and manganese extractant is saturated with cobalt and manganese, and the saturated extractant 44 is separated from the productive leach solution 38. The cobalt product and manganese product are then recovered from the cobalt and manganese extractant.

The cobalt/manganese-depleted productive leach solution 48 is sent to a nickel solvent extraction circuit 800 where it is contacted with a nickel extractant 52 . The nickel extractant is saturated with nickel, and the saturated extractant 54 is separated from the productive leach solution 48. The nickel product is then recovered from the nickel extractant.

The resulting productive leach solution 56 is sent to lithium recovery circuit 900 where it is contacted with ammonium carbonate 58 to precipitate lithium carbonate. Product 60 in the form of lithium carbonate is recovered from solution. The resulting filtrate 62, which contains ammonia, is fed to ammonium sulfate recovery circuit 1000 to produce ammonium sulfate product 64.

A more detailed description of each of the above circuits is provided below.

Raw material.

In one embodiment of the present invention, the raw material comprises lithium ion batteries. The process of the present invention is designed to process a number of different types of batteries and recover a number of valuable metals. As used herein, unless the context otherwise requires, the term lithium-ion battery is to be understood to refer to any battery that contains lithium or lithium-containing compounds. It should also be understood that this term includes any components from such batteries, or waste from the production of such batteries.

In one embodiment, the feed stream contains battery cathode and anode material that contains lithium and one or more of cobalt, copper, manganese, nickel, aluminum and iron.

Pre-processing process.

The target valuable metals in lithium-ion batteries are found in the cathode and anode. Batteries available on the market typically contain a cathode and an anode housed in a housing. In some cases, there is also a plastic covering wrapped around the body. Since the body and coating do not contain any valuable metals, they are preferably removed before the step of subjecting the feed stream to sulfuric acid leaching.

In one embodiment of the present invention, the raw material is subjected to a mechanical processing step. Preferably, the mechanical processing step includes a crushing and/or grinding step. The inventors have discovered that by mechanically treating used lithium-ion batteries, it is possible to remove anode and cathode materials from lithium-ion battery cases.

In one embodiment of the present invention, the raw materials are subjected to a size reduction step. Preferably, the size reduction step includes a milling or grinding step. The inventors have discovered that particle size reduction is advantageous for subsequent acid leach steps.

In one embodiment of the present invention, the average particle size of the raw material is reduced to less than 5 mm. Preferably, the average particle size of the raw material is reduced to less than 4 mm. Preferably, the average particle size of the raw material is reduced to less than 3 mm. Preferably, the average particle size of the raw material is reduced to less than 2 mm. Preferably, the average particle size of the raw material is reduced to less than 1 mm. Preferably, the average particle size of the raw material is reduced to less than 900 μm. Preferably, the average particle size of the raw material is reduced to less than 800 μm. Preferably, the average particle size of the raw material is reduced to less than 700 μm. Preferably, the average particle size of the raw material is reduced to less than 600 μm. Preferably, the average particle size of the raw material is reduced to less than 500 μm.

In one embodiment, one or more pre-treatment steps are performed under an inert atmosphere.

In one embodiment of the present invention, the ground material is subjected to one or more enrichment steps. Preferably, the beneficiation step separates at least some of the non-value materials from the feed stream.

In one embodiment of the present invention, one or more enrichment steps include an air classification step. The inventors have found that air classification is useful for separating lightweight plastics from heavier valuable metals.

In one embodiment of the present invention, one or more enrichment steps include a magnetic separation step. The inventors have found that magnetic separation is useful in separating the magnetic housing material from the anode and cathode materials.

- 7 045218

In one embodiment, the size reduction step is performed before the enrichment step. In an alternative embodiment, the size reduction step is performed after the enrichment step.

Leach circuit.

As described above, the treated raw materials are sent to the leach circuit where they are subjected to sulfuric acid leaching to form a slurry containing a productive leaching solution of soluble metal salts and a solid residue. In fig. 2 shows a flow diagram of a leach circuit 200 in accordance with one embodiment of the present invention. In leach circuit 200, feed material 202 is mixed with water 204 and washed filtrate 206 to form a slurry with a target solids content. The target solids content depends on the quality of the raw material and is adjusted so that sufficient dilution water is present in the leaching outlet to maintain all soluble salts in solution.

The suspension is fed into leaching reactors 208 in series. Concentrated sulfuric acid 210 is added to the leach reactors 208 in sufficient excess to maintain the free acid concentration. At least a portion of any of the copper, cobalt, aluminum, iron, nickel, lithium, and manganese is extracted to form leach solution 212. Metals and their compounds react with sulfuric acid to produce their corresponding sulfate salts through a variety of potential reaction routes, with some of the main reactions listed below:

Cu (sol.) + 2H2SO4 (aq.) ^- CuSO ₄ (aq.) 3 SO _{2 (r)} + 2H ₂ O

2LiCoO ₂ (sol.) 3 H ₂ SO ₄ (aq.) 3 2SO ₂ ( _g .) ^- Li ₂ SO4 (aq.) 3 2COSO4 (aq.) 3 H ₂ ( _g. )

2LiMnNiCoO ₂ ( _P acT.) s~ 7H ₂ 8O ₄ (aq.) Εϊ ₂ 8Ο ₄ (βο ₄ η ₄ +2Νϊ8Ο ₄ ( _Β ο ₄ η ₄ s~ 2Mn8O4( _In sol.) +

2CoSO ₄₍ bo «h.) + 4Н ₂ О + ЗН ₂ ( _Г )

2A1 (sol.) 3 3H ₂ SO ₄ (aq.) ^- A1 ₂ (SO ₄ ) ₃ (aq.) + ZN ₂ ( _g )

2FC(pacr.) 3 3H ₂ SO4 (aq.) ^- Fe ₂ (SO ₄ ) ₃ (aq.) 3 ZN ₂ ( _g )

Ni(pacT.) + H ₂ SO ₄ (aq.) ^- N1SO4 (aq.) 3 H ₂ ( _g )

MP( _raS t ₄ 3 H ₂ SO ₄ (aq.) ^- MnSO ₄ (aq.) 3 H ₂ ( _g )

2LiOH(pacT.) 3 H ₂ SO ₄ (aq.) ^- Li ₂ SO ₄ (aq.) 3 2H ₂ O (g.)

If multiple leach reactors 108 are used, the leach reactors are arranged in series.

In one embodiment, a gas scrubbing system (not shown) is connected to the headspace of the leach reactors 208 to remove water vapor and other gases produced in the leach reactors 208.

Although the vat leach process is described above, it is contemplated that one skilled in the art could select from any leach circuit known in the art to achieve the same effect.

The resulting leach solution 212 is a suspension containing a productive leach solution and a solid residue. The leach solution 212 is sent to a thickener 214 to dewater the leach solution 212 slurry. The overhead product 216 from the thickener 214 contains no solids and is collected in a holding tank (not shown) for further processing. Underflow 218 from thickener 214 is released at a high solids content of approximately 65 wt.% solids to leach filter 220.

Underflow 218 is fed to leach filter 220 where it is further dewatered to produce a relatively dry cake having a moisture content of approximately 15%. The main filtrate 222 is collected and sent to a holding tank (not shown) for further processing. Fresh water 224 is used in the leach filter 220 to wash the cake and maximize recovery of the productive leach solution trapped in the filter cake. The washed filtrate 206 is returned to the beginning of the leach circuit 200, where it is mixed with the raw material 202. The filter cake is discharged from the leach filter 220 and supplied to a residue storage facility for disposal.

Operating conditions of the leaching circuit.

In one embodiment of the present invention, the raw material is mixed with water and/or washed filtrate until the target solids content is achieved. Preferably, the target solids content is in the range of 5 wt.% to 70 wt.% More preferably, the target solids content is in the range of 10 wt.% to 30 wt.% Even more preferably, the target solids content is approximately 20 wt.% The inventors have discovered that the preferred target solids content depends on the quality of the raw material. Generally,

- 8 045218 the higher the quality of the raw material, the lower the target solids content.

In one embodiment of the present invention, the concentration of sulfuric acid is at least 1 g/L. Preferably the concentration of sulfuric acid is at least 2 g/l. Preferably the concentration of sulfuric acid is at least 3 g/l. Preferably the concentration of sulfuric acid is at least 4 g/l. Preferably the concentration of sulfuric acid is at least 5 g/l. Preferably the concentration of sulfuric acid is at least 6 g/l. Preferably the concentration of sulfuric acid is at least 7 g/l. Preferably the concentration of sulfuric acid is at least 8 g/l. Preferably the concentration of sulfuric acid is at least 9 g/l. Preferably the concentration of sulfuric acid is at least 10 g/l. As one skilled in the art will appreciate, the concentration of sulfuric acid in the leach solution affects the reaction/dissolution kinetics. If it happens too slowly, the reaction will slow down to a point where it becomes impractical.

In one embodiment, the amount of sulfuric acid is sufficient to maintain a free acid concentration of at least 1 g/L. In one embodiment, the amount of sulfuric acid is sufficient to maintain a free acid concentration of at least 2 g/L. In one embodiment, the amount of sulfuric acid is sufficient to maintain a free acid concentration of at least 3 g/L. In one embodiment, the amount of sulfuric acid is sufficient to maintain a free acid concentration of at least 4 g/L. In one embodiment, the amount of sulfuric acid is sufficient to maintain a free acid concentration of at least 5 g/L. In one embodiment, the amount of sulfuric acid is sufficient to maintain a free acid concentration of at least 6 g/L. In one embodiment, the amount of sulfuric acid is sufficient to maintain a free acid concentration of at least 7 g/L. In one embodiment, the amount of sulfuric acid is sufficient to maintain a free acid concentration of at least 8 g/L. In one embodiment, the amount of sulfuric acid is sufficient to maintain a free acid concentration of at least 9 g/L. In one embodiment, the amount of sulfuric acid is sufficient to maintain a free acid concentration of at least 10 g/L.

Preferably, the step of subjecting the feed stream to sulfuric acid leaching to form a suspension containing a productive leaching solution of soluble metal salts and a solid residue is performed at atmospheric pressure. Methods for leaching materials at atmospheric pressure are well known to those skilled in the art and include cuvette leaching and vat leaching.

In preferred embodiments of the present invention, the step of subjecting the feed stream to sulfuric acid leaching to form a slurry containing a productive leaching solution of soluble metal salts and a solid residue occurs at elevated temperatures. In one embodiment, the step of subjecting the feed stream to sulfuric acid leaching to form a slurry containing a productive soluble metal salt leach solution and a solid residue is performed at a temperature of at least 40°C. In one embodiment, the step of subjecting the feed stream to sulfuric acid leaching to form a slurry containing a productive leaching solution of soluble metal salts and a solid residue is performed at a temperature of at least 45°C. In one embodiment, the step of subjecting the feed stream to sulfuric acid leaching to form a slurry containing a productive soluble metal salt leach solution and a solid residue is performed at a temperature of at least 50°C. In one embodiment, the step of subjecting the feed stream to sulfuric acid leaching to form a slurry containing a productive leaching solution of soluble metal salts and a solid residue is performed at a temperature of at least 55°C. In one embodiment, the step of subjecting the feed stream to sulfuric acid leaching to form a slurry containing a productive leaching solution of soluble metal salts and a solid residue is performed at a temperature of at least 60°C. In one embodiment, the step of subjecting the feed stream to sulfuric acid leaching to form a slurry containing a productive soluble metal salt leach solution and a solid residue is performed at a temperature of at least 65°C. In one embodiment, the step of subjecting the feed stream to sulfuric acid leaching to form a slurry containing a productive soluble metal salt leach solution and a solid residue is performed at a temperature of at least 70°C. In one embodiment, the step of subjecting the feed stream to sulfuric acid leaching to form a slurry containing a productive soluble metal salt leach solution and a solid residue is performed at a temperature of at least 75°C. In one embodiment, the step of subjecting the feed stream to sulfuric acid leaching to form a slurry containing a productive soluble metal salt leach solution and a solid residue is performed at a temperature of at least 80°C. In one embodiment, the step of exposing the feed stream to sulfuric acid

- 9 045218 leaching to form a suspension containing a productive leaching solution of soluble metal salts and a solid residue is carried out at a temperature of at least 85°C. In one embodiment, the step of subjecting the feed stream to sulfuric acid leaching to form a slurry containing a productive soluble metal salt leach solution and a solid residue is performed at a temperature of at least 90°C.

In one embodiment of the present invention, if leaching is performed at elevated temperature, the leaching reactors are preferably heated using steam.

In one embodiment of the present invention, the step of subjecting the feed stream to sulfuric acid leaching to form a slurry containing a productive leaching solution of soluble metal salts and a solid residue, more preferably comprising:

subjecting the feed stream to sulfuric acid leaching in the presence of an oxidizing agent to form a suspension containing a productive leaching solution of soluble metal salts and a solid residue.

As will be appreciated by one skilled in the art, copper metal does not dissolve readily and directly in sulfuric acid, resulting in slow and incomplete extraction. The inventors have found that providing an oxidizing agent during the sulfuric acid leaching step allows for almost complete leaching of copper. In one embodiment of the present invention, the oxidizing agent is hydrogen peroxide.

Solvent extraction.

After the step of separating the productive leach solution and the solid residue, the productive leach solution is fed to one or more separate solvent extraction steps to recover the valuable metals.

As used herein, unless the context otherwise requires, the term extractant is to be understood to refer to an organic compound dissolved in an organic solvent. An organic compound typically has an available proton that can be exchanged for a metal ion from an aqueous solution.

As used herein, unless the context otherwise requires, the term saturated extractant and its variants such as saturated organic matter are to be understood to refer to an extractant that is enriched with metal ions that have passed into the extractant from aqueous solution.

As used herein, unless the context otherwise requires, the term scrubbing is to be understood as referring to a rich extractant purification step in which at least some of the undesirable elements are removed.

As used herein, unless the context otherwise requires, the term stripping is to be understood as the step in which the metal of interest is transferred from the saturated extractant to the aqueous phase by the addition of a stripping solution, typically in the form of a dilute or concentrated acid or basic solution.

Solvent copper extraction circuit.

After the step of separating the productive leach solution and the solid residue, the productive leach solution is supplied to the copper solvent extraction circuit to recover the copper. In fig. 3 shows a flow diagram of a copper solvent extraction circuit 300 in accordance with one embodiment of the present invention. The productive leach solution 302 from the leach circuit 200 is sent to the copper solvent extraction circuit 300 . In the embodiment shown in FIG. 3, the copper solvent extraction circuit 300 includes an extraction stage 304, a scrubbing stage 306, and a stripping stage 308 for selectively extracting copper from the productive leach solution 302 to produce a stripped copper leach solution 310. The copper-free raffinate 312 is sent to a holding tank for further processing.

Before being sent to the extraction stage 304, the productive leach solution 302 is pumped through a countercurrent heat exchanger (not shown) to reduce the temperature of the productive leach solution 302.

The cooled productive leach solution 302 is sent to an extraction stage 304 .

In one embodiment of the present invention, the extraction stage includes one or more mixer-settlers for solvent extraction. More preferably, the extraction stage includes two or more mixer-settlers for solvent extraction. Even more preferably, the extraction stage comprises three or more mixer-settlers for solvent extraction.

In one embodiment, if the extraction phase contains two or more solvent extraction mixer-settlers, the two or more solvent extraction mixer-settlers are arranged in series.

In one embodiment, if the extraction stage contains two or more solvent extraction mixer-settlers arranged in series, the mixer-settlers operate in countercurrent.

- 10 045218

In the embodiment shown in FIG. 3, extraction stage 304 comprises three solvent extraction mixer and settlers arranged in series. The inventors envision that one or more mixer-settlers may be used. The productive leach solution 302 is sent to the first mixer-settler where it is contacted with a copper extractant 314 to selectively extract copper from the productive leach solution 302.

In a preferred embodiment of the present invention, the copper extractant is an organic extractant. More preferably, the copper extractant is an oxime. Even more preferably, the copper extractant is LIX84I.

The overall extraction reaction is presented in Equation 1 below:

2RH + CuSO ₄ = CuR ₂ + H ₂ SO ₄ (1)

The productive leach solution 302 and the copper extractant 314 are brought into countercurrent contact to maximize extraction efficiency. Ammonia solution 316 is dosed into each of the mixer-settlers to maintain the target pH by neutralizing the sulfuric acid through the reaction shown in Equation 2:

H2SO4 + 2NH ₃ - (NH4) ₂ SO ₄ (2)

Copper is selectively removed from the productive leach solution 302, and the copper-free raffinate 312 is sent to a holding tank for further processing.

The saturated extractant 318 is sent to the scrubbing stage 306. In scrubbing step 306, the rich copper extractant 318 is contacted with a portion of the wash solution 320 that contains copper ions. The copper in the wash solution 320 predominantly saturates the saturated extractant 318 and displaces impurity elements that saturate the saturated extractant 318 in the extraction step 304.

In one embodiment of the present invention, the scrubbing stage comprises one or more mixer-settlers.

In the embodiment shown in FIG. 3, the scrubbing stage 306 contains one mixer-settler.

The aqueous phase from the scrubbing stage 306 is sent back to the first mixer-settler for extraction. The saturated extractant 322 from the scrubbing stage 306 moves to the stripping stage 308. In stripping step 308, the saturated extractant 322 is contacted with the acidic stripping solution 324 to displace the majority of the copper ions in the organic phase (as opposed to the extraction reaction detailed in Equation 1) into the aqueous phase to produce a stripped copper leach solution 310.

In one embodiment, the stripping stage includes one or more settling mixers. Preferably, the stripping stage contains two or more mixer-settlers.

In one embodiment, if the stripping stage contains two or more mixer-settlers, the two or more mixer-settlers are arranged in series.

The organic phase 326 leaving the stripping stage 308 is returned to the extraction stage 304 where it is again saturated with copper. Thus, the organic phase 326 is kept in a closed loop in the copper solvent extraction circuit 300.

A portion of the stripped copper solution after leaching is returned for use as a wash solution in the scrubbing stage 306.

The distilled copper solution 310, after leaching from the distillation stage 306, is sent to the copper crystallization stage 328. In the copper crystallization step 328, the distilled copper solution 310 after leaching is contacted with sulfuric acid 330 to cause copper sulfate pentahydrate (CuSO ₄ .5H ₂ O) to crystallize as a solid in the salting out crystallization process. The copper crystallization step 328 produces a copper sulfate pentahydrate slurry 332 and a sulfuric leach solution.

In one embodiment of the present invention, the copper crystallization stage comprises a crystallization unit. Preferably, the crystallization step is performed in a jacketed reaction vessel or crystallizer with a cleanable surface.

The suspension 332 from the copper crystallization stage 328 is fed to the filtration stage 334. In a filtration stage 334, copper sulfate solids 336 are separated from the acidic solution 338 after leaching. Water 340 is used in the filtration stage 334 to flush out the filter cake and maximize the removal of trapped acidic solution 338 after leaching into the cake.

Preferably, the filtration stage includes a device for separating solid and liquid phases. More preferably, the solid-liquid separation device is a plate-frame filter.

In one embodiment, the acidic solution 338, after leaching, is returned back to the stripping stage 306. A portion of the copper filtrate is returned for use in scrubbing stage 306.

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Operating conditions of the copper solvent extraction circuit.

In one embodiment of the present invention, the productive leach solution is cooled before being sent to the extraction stage. In one embodiment, the productive leach solution is cooled to a temperature below 65°C. In one embodiment, the productive leach solution is cooled to a temperature below 60°C. In one embodiment, the productive leach solution is cooled to a temperature below 55°C. In one embodiment, the productive leach solution is cooled to a temperature below 50°C. In one embodiment, the productive leach solution is cooled to a temperature below 45°C. In one embodiment, the productive leach solution is cooled to a temperature below 40°C. In one embodiment, the productive leach solution is cooled to a temperature below 35°C. In one embodiment, the productive leach solution is cooled to a temperature below 30°C. Preferably, the productive leach solution is cooled to a temperature of at least below 50°C. The maximum temperature is determined by the flash point, and operation is advisable at a temperature at least 10°C below this temperature.

In one embodiment of the present invention, the copper extractant is diluted to the target concentration using a diluent. Preferably, the target concentration is in the range of 20 vol.% to 40 vol.% by volume. More preferably, the target concentration is approximately 30 vol%. As one skilled in the art will appreciate, dilution of the copper extractant is used to adjust the viscosity.

In embodiments of the present invention, if the copper extractant is diluted with a diluent, the diluent is Shellsol2046.

In a preferred embodiment of the present invention, the ratio of organic and aqueous phases in the extraction step is 2.15:1. As will be appreciated by one skilled in the art, the ratio of organic to aqueous phases in the extraction step depends on the copper content of the productive leach solution, as well as the copper saturation of the organic phase. One way to calculate the equilibrium concentration of an extractant is to use a material balance, i.e. it is equal to the difference between the total (analytical) concentration of the extractant and the sum of all solvated substances in the soluble phase.

Preferably, the target pH of the productive leach solution in the extraction stage is in the range of 1 to 3. More preferably, the target pH of the productive leach solution in the extraction stage is 2.

In one embodiment of the present invention, the ratio of saturated copper extractant to wash solution is in the range of 20:1 to 100:1 (organic to aqueous phase ratio) by volume. Preferably the ratio is in the range of 25:1 to 50:1 (organic to aqueous phase ratio) by volume. More preferably, the ratio is approximately 35:1 (organic to aqueous phase ratio) by volume.

In one embodiment of the present invention, the acidic stripping solution 324 contains sulfuric acid. Preferably, the concentration of sulfuric acid in the stripping solution is at least 100 g/l. More preferably, the concentration of sulfuric acid in the stripping solution is at least 110 g/L. More preferably, the concentration of sulfuric acid in the stripping solution is at least 120 g/L. More preferably, the concentration of sulfuric acid in the stripping solution is at least 130 g/L. More preferably, the concentration of sulfuric acid in the stripping solution is at least 140 g/L. More preferably, the concentration of sulfuric acid in the stripping solution is at least 150 g/L.

In one embodiment of the present invention, the concentration of sulfuric acid in the suspension leaving the crystallization stage is in the range of 150 to 350 g/l. Preferably, the concentration of sulfuric acid in the suspension leaving the crystallization stage is in the range from 160 to 300 g/l. Preferably, the concentration of sulfuric acid in the suspension leaving the crystallization stage is in the range from 170 to 275 g/l. Preferably, the concentration of sulfuric acid in the suspension leaving the crystallization stage is in the range from 180 to 260 g/l. Preferably, the concentration of sulfuric acid in the suspension leaving the crystallization stage is in the range from 190 to 250 g/l. Preferably, the concentration of sulfuric acid in the suspension leaving the crystallization stage is in the range from 200 to 340 g/l. Preferably, the concentration of sulfuric acid in the suspension leaving the crystallization stage is approximately 220 g/l.

In one embodiment of the present invention, the temperature of the crystallization step is controlled. Preferably, the temperature of the crystallization step is in the range from 10°C to 50°C. Preferably, the temperature of the crystallization step is in the range from 20°C to 40°C. Preferably, the temperature of the crystallization step is in the range from 25°C to 30°C. As one skilled in the art will appreciate, at lower temperatures copper sulfate is less soluble, which maximizes its recovery.

Aluminum solvent extraction circuit.

After the step of separating the productive leach solution and the solid residue, and preferably

- 12 045218 Finally, after the copper solvent extraction step, the productive leaching solution is fed into the aluminum solvent extraction circuit to extract aluminum. In fig. 4 shows a flow diagram of an aluminum solvent extraction circuit 400 in accordance with one embodiment of the present invention. The productive leach solution 402 from the copper solvent extraction circuit is sent to the aluminum solvent extraction circuit 400 . In the embodiment shown in FIG. 4, the aluminum solvent extraction circuit 400 includes an extraction stage 404, a scrubbing stage 406, and a stripping stage 408 for selectively recovering aluminum from the productive leach solution 402 to produce a stripped aluminum post-leaching solution 410. The aluminum-free raffinate 412 is sent to a holding tank for further processing.

In one embodiment of the present invention, the extraction stage includes one or more mixer-settlers for solvent extraction. More preferably, the extraction stage includes two or more mixer-settlers for solvent extraction. Even more preferably, the extraction stage comprises three or more mixer-settlers for solvent extraction.

In one embodiment, if the extraction phase contains two or more solvent extraction mixer-settlers, the two or more solvent extraction mixer-settlers are arranged in series.

In one embodiment, if the extraction stage contains two or more solvent extraction mixer-settlers arranged in series, the mixer-settlers operate in countercurrent.

In the embodiment shown in FIG. 4, extraction stage 404 contains three solvent extraction mixer and settlers arranged in series. The inventors envision that one or more mixer-settlers may be used. The productive leach solution 402 is sent to the first mixer-settler where it is contacted with an aluminum extractant 414 to selectively extract aluminum from the productive leach solution 402. Iron and any remaining copper ions also extract aluminum from the productive leach solution 402.

In a preferred embodiment of the present invention, the aluminum extractant is an organic extractant. More preferably, the aluminum extractant is selected from an organic phosphinic acid extractant, an organic phosphonic acid extractant, and an organic orthophosphoric acid extractant. Even more preferably, the aluminum extractant is selected from Cyanex272, Ionquest 801 (PC88A) and D2EHPA.

The overall extraction reaction is presented in Equation 3 below:

The productive leach solution 402 and the aluminum extractant 414 are contacted in countercurrent to maximize extraction efficiency. Ammonia solution 416 is dosed into each of the mixer-settlers to maintain the target pH by neutralizing the sulfuric acid through the reaction shown in Equation 2.

Aluminum, iron and copper ions are selectively removed from the productive leach solution 402, and the aluminum-free raffinate 412 is sent to a holding tank for further processing.

The saturated extractant 418 is sent to the scrubbing stage 406. In scrubbing step 406, the saturated aluminum extractant 418 is contacted with a portion of the wash solution 420 that contains aluminum ions. The aluminum ions in the wash solution 420 predominantly saturate the saturated extractant 418 and displace the impurity elements that saturated the saturated extractant 418 in the extraction step 404.

In one embodiment of the present invention, the scrubbing stage comprises one or more mixer-settlers.

In the embodiment shown in FIG. 4, the scrubbing stage 406 contains one mixer-settler.

The aqueous phase from the scrubbing stage 406 is sent back to the first mixer-settler for extraction. The saturated extractant 422 from the scrubbing stage 406 moves to the stripping stage 408. In stripping step 408, saturated extractant 422 is contacted with acidic stripping solution 424. The acidic stripper results in the displacement of most of the ions in the organic phase (as opposed to the extraction reaction detailed in Equation 3) into the aqueous phase to produce stripped aluminum solution 410 after leaching. A sufficient amount of water 426 is added to the stripping stage 406 to prevent crystallization of salts in the stripped aluminum solution 410 after leaching. After leaching, the distilled aluminum solution 410 is sent to the storage of aluminum products. A portion of the stripped aluminum solution 410 is returned after leaching for use as a wash solution in the scrubbing stage.

In one embodiment, the stripping stage 408 includes one or more mixer-13 045218 settlers. Preferably, the stripping stage 408 includes two or more mixer-settlers.

In one embodiment, if stripping stage 408 contains two or more mixer-settlers, the two or more mixer-settlers are arranged in series.

The organic phase 428 leaving the stripping stage 408 is returned to the extraction stage 404 where it is again saturated with aluminum. Thus, the organic phase 428 is kept in a closed loop in the aluminum solvent extraction circuit 400.

Operating conditions of the aluminum solvent extraction circuit.

In one embodiment of the present invention, the aluminum extractant is diluted to a target concentration using a diluent. Preferably, the target concentration is in the range of 20 vol.% to 40 vol.% by volume. More preferably, the target concentration is approximately 30 vol%.

In embodiments of the present invention, if the aluminum extractant is diluted with a diluent, the diluent is Shellsol2046.

In a preferred embodiment of the present invention, the ratio of organic and aqueous phases in the extraction step is 1.85:1. As will be understood by one skilled in the art, the ratio of organic and aqueous phases in the extraction stage depends on the aluminum content in the productive leach solution, as well as on the aluminum saturation of the organic phase. One way to calculate the equilibrium concentration of an extractant is to use a material balance, i.e. it is equal to the difference between the total (analytical) concentration of the extractant and the sum of all solvated substances in the soluble phase.

Preferably, the target pH of the productive leach solution in the extraction step is in the range of 2 to 3.5. More preferably, the target pH of the productive leach solution in the extraction step is 3.3.

In one embodiment of the present invention, the ratio of saturated aluminum extractant to wash solution is in the range of 20:1 to 100:1 (organic to aqueous phase ratio) by volume. Preferably the ratio is in the range of 30:1 to 70:1 (organic to aqueous phase ratio) by volume. More preferably, the ratio is approximately 50:1 (organic to aqueous phase ratio) by volume.

In one embodiment of the present invention, the saturated extractant is contacted with a sufficient amount of acidic stripping solution to achieve a target pH in the stripped aluminum solution after leaching. Preferably, the target pH is less than 2. More preferably, the target pH is less than 1.5. More preferably, the target pH is less than 1. The inventors have discovered that if the pH is greater than 2, the acid concentration is not sufficient to strip the metal from the extractant.

In one embodiment of the present invention, water is added to the stripping step to achieve a target concentration of aluminum in the stripped aluminum solution after leaching. Preferably, the target concentration is in the range of 50 g/L to 70 g/L.

Aluminum recovery circuit.

As described above, aluminum and iron can be recovered from the productive leach solution through a solvent extraction process. In fig. 4a shows an alternative embodiment of the present invention in which aluminum and iron are recovered from the productive leach solution in the aluminum precipitation loop 430. Reference numerals similar to those in FIG. 3 indicate similar elements. Similar to the description above, the productive leach solution is preferably supplied to the aluminum precipitation loop 430 after the copper solvent extraction step.

In the embodiment shown in FIG. 4a, the aluminum precipitation circuit 430 includes a primary precipitation stage 431, a primary thickening stage 432, a secondary precipitation stage 433, and a secondary thickening stage 434 for selectively precipitating and recovering aluminum and iron from the productive leach solution 402. The aluminum-free raffinate 412 is sent to a holding tank for further processing.

In the primary precipitation stage 431, the solution is contacted with an oxidizing agent 435 to oxidize the ferrous iron molecules in the solution to ferric ions. Suitable oxidizing agents are known to those skilled in the art and include air, oxygen gas, oxygen-enriched gases, and hydrogen peroxide. The pH of the solution is increased in the range of 3.2 to 4 by the addition of a base, such as ammonia 436. A number of other bases, such as lime or limestone, can be used to increase the pH of the solution. The main part of aluminum and iron particles in solution are deposited in the form of hydroxides in suspension 437.

The resulting suspension 437 is sent to the first thickening stage 432. The slurry 437 is concentrated and the overhead product 438 is sent to a secondary precipitation stage 433. The bottom product 439 from the first thickening stage 432 is sent to the filtration stage 440. The filter residue 441 is removed and sent for further processing or disposal. The filtrate 442 is sent to a secondary precipitation stage 433.

Preferably, the first thickening stage 431 uses a conventional thickening device.

- 14 045218

Other means of solid-liquid separation employed by those skilled in the art may be used.

In a secondary precipitation stage 433, the solution is contacted with an oxidizing agent 443 to oxidize the ferrous iron molecules in the solution to ferric ions. Suitable oxidizing agents are known to those skilled in the art and include air, oxygen gas, oxygen-enriched gases, and hydrogen peroxide. The pH of the solution is increased to above 4.5 by adding a 445 base such as ammonia. In one embodiment, the pH is increased to a level in the range of 4.5-6.0. A number of other bases, such as lime or limestone, can be used to increase the pH of the solution. It has been found that the use of oxidizing agent 443 in both stages of precipitation is advantageous because the rate of oxidation of ferrous iron to ferric iron is low at the low pH of the first stage and much higher at the higher pH of the second stage. The remaining aluminum and iron particles in solution are precipitated as hydroxides in suspension 446.

The resulting slurry 446 is sent to a secondary thickening stage 434. The slurry 446 is concentrated and the overflow 412, which is substantially free of aluminum and iron, is sent to a holding tank for further processing. The thickener underflow 447 may be returned to the first precipitation stage 431 to minimize losses. Alternatively or additionally, the thickener underflow 447 may be sent to a leach loop or a separate acid dissolution stage to recover co-precipitated cobalt, nickel and lithium.

Preferably, the secondary thickening stage 434 uses a conventional thickening device.

The inventors have also discovered that the primary precipitation stage and the secondary precipitation stage can be replaced by a single precipitation stage at a pH in the range of 4.5-6 in the presence of an oxidizing agent. However, the disadvantage of this approach is the large losses of cobalt.

Operating conditions of the aluminum recovery circuit.

In one embodiment of the present invention, the primary precipitation step is carried out at a temperature above 50°C. Preferably, the primary precipitation step is carried out at a temperature above 55°C. More preferably, the primary deposition step is carried out at a temperature above 60°C. Even more preferably, the primary precipitation step is carried out at a temperature above 65°C.

In one embodiment of the present invention, the secondary precipitation step is carried out at a temperature above 50°C. Preferably, the secondary precipitation step is carried out at a temperature above 55°C. More preferably, the secondary precipitation step is carried out at a temperature above 60°C. Even more preferably, the secondary precipitation step is carried out at a temperature above 65°C.

In one embodiment of the present invention, the Eh (redox potential) of the solution at at least one of the deposition steps is at least 400 mV Ag/AgCl.

Cobalt/manganese recovery circuit.

After the step of separating the productive leach solution and the solid residue, and preferably after the copper solvent extraction step and the aluminum solvent extraction step, the productive leach solution is fed to the cobalt/manganese recovery circuit to recover cobalt and manganese. In fig. 5 shows a flow diagram of a cobalt/manganese recovery circuit 500 in accordance with one embodiment of the present invention. The productive leach solution 502 from the aluminum solvent extraction circuit 400 is sent to the cobalt/manganese recovery circuit 500 .

In the embodiment shown in FIG. 5, the cobalt/manganese recovery circuit 500 includes a cobalt and manganese solvent extraction circuit 504 for selectively removing cobalt/manganese from the productive leach solution 502. Cobalt- and manganese-free raffinate 506 is sent to a holding tank for further processing. The resulting cobalt and manganese stream 507 is sent for further processing. The cobalt/manganese recovery circuit 500 further includes a manganese solvent extraction circuit 508 for separating cobalt and manganese. The resulting cobalt-containing stream 510 is sent to a cobalt crystallization stage 512 to recover cobalt product 514. The resulting manganese-containing stream 516 is sent to a manganese crystallization stage 518 to recover manganese product 520.

Solvent extraction circuit for cobalt and manganese.

A detailed view of the cobalt and manganese solvent extraction circuit 504 is shown in FIG. 6, with reference numerals similar to those in FIG. 5 indicate similar elements. The cobalt and manganese solvent extraction circuit 504 includes an extraction stage 602, a scrubbing stage 604, and a stripping stage 606 for selectively extracting cobalt/manganese from the productive leach solution 502 to produce a stripped cobalt and manganese leach solution 608. Cobalt- and manganese-free raffinate 506 is sent to a holding tank for further processing.

In one embodiment of the present invention, the extraction stage includes one or more mixer-settlers for solvent extraction. More preferably, the extraction stage contains two or more mixer-settlers for solvent extraction. Even more preferably, the extraction stage comprises three or more mixer-settlers for solvent extraction.

In one embodiment, if the extraction phase contains two or more solvent extraction mixer-settlers, the two or more solvent extraction mixer-settlers are arranged in series.

In one embodiment, if the extraction stage contains two or more solvent extraction mixer-settlers arranged in series, the mixer-settlers operate in countercurrent.

In the embodiment shown in FIG. 6, extraction stage 602 contains four solvent extraction mixer-settlers arranged in series. The inventors envision that one or more mixer-settlers may be used. The productive leach solution 502 is sent to the first mixer-settler where it is contacted with a cobalt-manganese extractant 610 to selectively extract cobalt/manganese from the productive leach solution 502.

In a preferred embodiment of the present invention, the cobalt and manganese extractant is selected from an organic phosphinic acid extractant, an organic phosphonic acid extractant, and an organic orthophosphoric acid extractant. Even more preferably, the cobalt and manganese extractant is selected from Cyanex272, Ionquest 801 (PC88A) and D2EHPA.

The overall extraction reaction is presented in Equation 4 below:

2RH + XSO ₄ = XR ₂ + H ₂ SO ₄ (where X = Co ⁺² , Mn ² ) (4)

The productive leach solution 502 and the cobalt and manganese extractant 610 are contacted in countercurrent to maximize extraction efficiency. Ammonia solution 612 is dosed into each of the mixer-settlers to maintain the target pH by neutralizing the sulfuric acid through the reaction shown in Equation 2.

Cobalt and manganese are selectively removed from the productive leach solution 502, and the cobalt- and manganese-free raffinate 506 is sent to a holding tank for further processing.

The saturated extractant 614 is sent to the scrubbing stage 604. In scrubbing step 604, the saturated extractant 614 is brought into contact with a portion of the rinse solution 616. The cobalt/manganese in the rinse solution 616 advantageously saturates the saturated extractant 614 and displaces some of the impurity elements that saturated the saturated extractant 614 in the extraction step 602.

In one embodiment of the present invention, the scrubbing stage comprises one or more mixer-settlers.

In the embodiment shown in FIG. 6, the scrubbing stage 604 contains one mixer-settler.

The aqueous phase from the scrubbing stage 604 is sent back to the first mixer-settler for extraction. The saturated extractant 618 from the scrubbing stage 604 moves to the stripping stage 606. In stripping step 606, saturated extractant 618 is contacted with acidic stripping solution 620 to displace the majority of the cobalt/manganese ions in the organic phase (as opposed to the extraction reaction detailed in Equation 1) into the aqueous phase to produce stripped cobalt solution 608 and manganese after leaching. A sufficient amount of water 622 is added to the stripping stage 606 to prevent crystallization of salts in the stripped cobalt and manganese solution 608 after leaching.

The distilled solution 608 of cobalt and manganese after leaching is sent to the manganese solvent extraction circuit 508.

In one embodiment, the stripping stage includes one or more settling mixers. Preferably, the stripping stage contains two or more mixer-settlers.

In one embodiment, if the stripping stage contains two or more mixer-settlers, the two or more mixer-settlers are arranged in series.

The organic phase 624 leaving the stripping stage 606 is returned to the extraction stage 602 where it is again saturated with cobalt/manganese. Thus, the organic phase 624 is kept in a closed loop in the cobalt/manganese solvent extraction loop 504. The distilled solution 608 of cobalt and manganese after leaching is sent to the manganese solvent extraction circuit 508.

Operating conditions of the cobalt/manganese solvent extraction circuit.

In one embodiment of the present invention, the cobalt and manganese extractant is diluted to the target concentration using a diluent. Preferably, the target concentration is in the range of 15 vol.% to 35 vol.%. More preferably, the target concentration is approximately 30 vol%. As one skilled in the art will appreciate, dilution of the cobalt and manganese extractant is used to adjust the viscosity.

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In embodiments of the present invention, if the cobalt and manganese extractant is diluted with a diluent, the diluent is Shellsol2046.

In a preferred embodiment of the present invention, the ratio of organic and aqueous phases in the extraction step is 3.1:1. As will be appreciated by one skilled in the art, the ratio of organic to aqueous phases in the extraction step depends on the cobalt/manganese content of the productive leach solution, as well as the cobalt/manganese saturation of the organic phase. One way to calculate the equilibrium concentration of an extractant is to use a material balance, i.e. it is equal to the difference between the total (analytical) concentration of the extractant and the sum of all solvated substances in the soluble phase.

Preferably, the target pH of the productive leach solution in the extraction stage is in the range of 4.2 to 5. More preferably, the target pH of the productive leach solution in the extraction stage is 4.8.

In one embodiment of the present invention, the ratio of the saturated cobalt-manganese extractant to the wash solution is in the range of 20:1 to 100:1 (ratio of organic to aqueous phases) by volume. Preferably the ratio is in the range of 30:1 to 70:1 (organic to aqueous phase ratio) by volume. More preferably, the ratio is approximately 50:1 (organic to aqueous phase ratio) by volume.

In one embodiment of the present invention, the acidic stripping solution contains sulfuric acid.

In one embodiment of the present invention, the saturated extractant is contacted with a sufficient amount of acidic stripping solution to achieve the target pH in the stripped cobalt-manganese leach solution. Preferably the target pH is less than 3.

In one embodiment of the present invention, water is added to the stripper stage to achieve a target cobalt concentration in the stripped cobalt-manganese solution after leaching. Preferably, the target concentration is in the range of 80 g/L to 140 g/L.

Solvent manganese extraction circuit.

A detailed view of the manganese extraction circuit 508 is shown in FIG. 7, with reference numerals similar to those in FIG. 5 and 6 indicate similar elements. The manganese extraction circuit 508 includes a pre-saturation stage 720, an extraction stage 722, a scrubbing stage 724, and a stripping stage 726 for selectively extracting manganese from the stripped cobalt and post-manganese solution 608 to produce a stripped post-leaching solution 730 of manganese. finate 738 cobalt.

The stripped cobalt and manganese solution 608, after leaching, is sent to an extraction stage 722 where it is contacted with the pre-saturated organic phase 734 from the pre-saturation stage 720. It is contemplated that the presaturated organic phase 734 may be supplemented or replaced by a fresh stream of manganese extractant (not shown).

In a pre-saturation step 720, the stripped organic phase 736 from the manganese stripper step 726 (described below) is contacted with the cobalt filtrate 740 from the cobalt filter 742 (described below). It is contemplated that the manganese extraction circuit 508 may be operated without the pre-saturation stage 720. In such an embodiment, fresh manganese extractant is introduced into extraction step 722. It has been found that the use of a pre-saturation stage 720 eliminates the need to introduce ammonia into the manganese solvent extraction loop.

In one embodiment of the present invention, the extraction stage includes one or more mixer-settlers for solvent extraction. More preferably, the extraction stage includes two or more mixer-settlers for solvent extraction. Even more preferably, the extraction stage comprises three or more mixer-settlers for solvent extraction.

In one embodiment, if the extraction phase contains two or more solvent extraction mixer-settlers, the two or more solvent extraction mixer-settlers are arranged in series.

In one embodiment, if the extraction stage contains two or more solvent extraction mixer-settlers arranged in series, the mixer-settlers operate in countercurrent.

In the embodiment shown in FIG. 7, the pre-saturation stage 720 contains three mixer-settlers arranged in series. The inventors envision that one or more mixer-settlers may be used. The cobalt filtrate 740 is sent to the first mixer-settler where it is contacted with a stripped organic phase 736 to selectively extract cobalt ions from the cobalt filtrate 740.

In a preferred embodiment of the present invention, the stripped organic phase 736 from the manganese stripper stage 726 contains an orthophosphoric acid organic extractant. Even more preferably, the distilled organic phase 736 is di(2-ethylhexyl)orthophosphoric acid (D2EHPA).

- 17 045218

The overall extraction reaction is presented in Equation 5 below:

2RH + XSO ₄ - XR2 + H2SO4 (5)

The cobalt filtrate 740 and the distilled organic phase 736 are contacted in countercurrent to maximize extraction efficiency. Ammonia solution 534 is metered into each of the mixer-settlers at a fixed flow rate to maintain the target pH by neutralizing the sulfuric acid through the reaction shown in Equation 2.

The flow rate of cobalt filtrate 740 to the presaturation stage 720 is adjusted to achieve the target cobalt saturation in the stripped organic phase 736. Preferably, the target cobalt saturation is such that the molar amount of saturated cobalt ions is substantially equivalent to the molar amount of manganese ions in the stripped solution 608 cobalt and manganese after leaching. The preferred ratio of cobalt to manganese is in the range of 1:1 to 1.1:1. Thus, the target saturation depends on the manganese content of the stripped cobalt-manganese solution 608 after leaching. If insufficient cobalt saturation occurs, ammonia must be introduced into the extraction stage to extract all the manganese.

The presaturated raffinate 510 still contains cobalt and thus can be sent to the extraction stage 602 of the cobalt and manganese solvent extraction loop 504.

The presaturated organic phase 734 is sent to an extraction stage 722 where it is contacted with the stripped cobalt and manganese leaching solution 608.

In one embodiment of the present invention, the extraction stage includes one or more mixer-settlers for solvent extraction. More preferably, the extraction stage includes two or more mixer-settlers for solvent extraction. Even more preferably, the extraction stage comprises three or more mixer-settlers for solvent extraction.

In one embodiment, if the extraction phase contains two or more solvent extraction mixer-settlers, the two or more solvent extraction mixer-settlers are arranged in series.

In one embodiment, if the extraction stage contains two or more solvent extraction mixer-settlers arranged in series, the mixer-settlers operate in countercurrent. In the embodiment shown in FIG. 7, extraction stage 722 contains three solvent extraction mixer-settlers arranged in series. The inventors envision that one or more settling mixers may be used. The stripped cobalt and manganese leach solution 608 is sent to a first mixer-settler where it is contacted with a presaturated organic phase 734 to selectively extract manganese from the stripped cobalt and manganese leach solution 608. Manganese ions preferentially saturate the presaturated organic phase 734, displacing cobalt ions according to the reaction presented in equation 6:

CoR ₂ + MnSO ₄ = MnR ₂ + CoSO ₄ (6)

The stripped cobalt and manganese leach solution 608 and the pre-saturated organic phase 734 are contacted in countercurrent to maximize extraction efficiency. Ammonia solution 744 is dosed into each of the mixer-settlers to maintain the target pH by neutralizing the sulfuric acid through the reaction shown in Equation 2.

Manganese is selectively recovered from the stripped cobalt-manganese solution 608 after leaching to produce a rich organic phase 746. The manganese-free raffinate 510 is sent to a cobalt crystallization stage 512.

The saturated organic phase 746 is sent to a scrubbing stage 724. In scrubbing step 724, the manganese organic phase 736 is contacted with a portion of the wash solution 478. The manganese in the wash solution 478 predominantly saturates the saturated organic phase 746 and displaces the cobalt that saturates the saturated organic phase 746, according to the reaction given in Equation 6.

In one embodiment of the present invention, the scrubbing stage comprises one or more mixer-settlers.

In the embodiment shown in FIG. 7, the scrubbing stage 724 contains one mixer-settler.

The aqueous phase from the scrubbing stage 724 is sent back to the first mixer-settler for extraction. The saturated extractant 750 from the scrubbing stage 724 moves to the stripping stage 726. In a stripping step 726, the saturated extractant 750 is contacted with an acidic stripping solution 752 to displace the majority of the manganese ions in the organic phase (as opposed to the extraction reaction detailed in Equation 4) into the aqueous phase to produce a stripped manganese leach solution 730. A sufficient amount of water 754 is added to the stripping stage 726 to prevent crystallization of salts in the stripped manganese solution 730 after leaching.

In one embodiment, the stripping stage includes one or more settling mixers. Preferably, the stripping stage contains two or more mixer-settlers.

In one embodiment, if the stripping stage contains two or more mixer-settlers, the two or more mixer-settlers are arranged in series.

The organic phase 736 leaving the stripping stage 726 is returned to the pre-saturation stage 720. Thus, the organic phase 736 is kept in a closed loop in the manganese solvent extraction loop 704. The distilled manganese solution 730 after leaching is sent to the manganese crystallization stage 518.

The distilled manganese solution 730, after leaching from the distillation stage 726, is sent to the manganese crystallization stage 518. In the manganese crystallization step 518, water is removed from the stripped manganese solution 730 after leaching to allow crystallization of MnSO ₄ .H ₂ O in slurry 756.

In one embodiment of the present invention, the manganese crystallization stage comprises a crystallization unit. Preferably, the crystallization unit is an evaporative crystallizer.

The slurry 756 from the manganese crystallization stage 518 is fed to the filtration stage 758. In filtration stage 758, manganese sulfate solids 520 are separated from the filtrate. Water 760 is used in the filtration stage 758 to flush away filter cake and maximize manganese recovery. The filtrate is returned back to the manganese crystallization stage 518 to maximize manganese recovery. A portion of the filtrate may also be returned for use in the manganese scrubbing stage 724.

The cobalt raffinate 510 from the extraction stage 722 is sent to the cobalt crystallization stage 512. In cobalt crystallization step 512, water is removed from cobalt raffinate 510 to allow crystallization of CoSO ₄ .7H ₂ O in slurry 762.

In one embodiment of the present invention, the cobalt crystallization stage 512 includes a crystallization unit. Preferably, the crystallization unit is an evaporative crystallizer.

The suspension 762 from the cobalt crystallization stage 512 is fed to the filtration stage 742. In filtration stage 742, cobalt sulfate solids 514 are separated from the filtrate. Water 764 is used in the filtration stage 742 to flush away filter cake and maximize cobalt recovery. The filtrate is returned back to the cobalt crystallization stage 512 to maximize cobalt recovery. A portion of the filtrate is returned for use in the pre-saturation stage 720.

Operating conditions of the manganese solvent extraction circuit.

In one embodiment of the present invention, the manganese extractant is diluted to a target concentration using a diluent. Preferably, the target concentration is in the range of 15 vol.% to 35 vol.%. More preferably, the target concentration is approximately 30 vol%. As will be appreciated by one skilled in the art, dilution of the manganese extractant is used to adjust the viscosity.

In embodiments of the present invention, if the manganese extractant is diluted with a diluent, the diluent is Shellsol2046.

Preferably, the target pH of the pre-saturation stage is in the range of 4.5 to 12. More preferably, the target pH of the pre-saturation stage is 4.8.

Preferably, the cobalt saturation in the presaturated organic phase is at least 10 g/l. Preferably, the cobalt saturation in the presaturated organic phase is at least 11 g/l. Preferably, the cobalt saturation in the presaturated organic phase is at least 12 g/l. Preferably, the cobalt saturation in the presaturated organic phase is at least 13 g/l. Preferably, the cobalt saturation in the presaturated organic phase is at least 13.5 g/L. Preferably, the cobalt saturation in the presaturated organic phase is approximately 13.8 g/L.

In one embodiment of the present invention, the presaturated organic phase is diluted to the target concentration using a diluent. Preferably, the target concentration is in the range of 15 vol.% to 35 vol.%. More preferably, the target concentration is approximately 30 vol%. As will be appreciated by one skilled in the art, dilution of the manganese extractant is used to adjust the viscosity.

In embodiments of the present invention, if the presaturated organic phase is diluted with a diluent, the diluent is Shellsol2046.

In a preferred embodiment of the present invention, the ratio of organic and aqueous phases in the extraction step is 1.65:1. As will be appreciated by one skilled in the art, the ratio of organic to aqueous phases in the extraction step depends on the manganese content of the productive leach solution, as well as the manganese saturation of the organic phase. One way

- 19 045218 calculation of the equilibrium concentration of the extractant is the use of material balance, i.e.

it is equal to the difference between the total (analytical) concentration of the extractant and the sum of all solvated substances in the soluble phase.

In one embodiment of the present invention, the ratio of saturated manganese extractant to wash solution is in the range of 20:1 to 100:1 (ratio of organic to aqueous phases) by volume. Preferably the ratio is in the range of 30:1 to 70:1 (organic to aqueous phase ratio) by volume. More preferably, the ratio is approximately 50:1 (organic to aqueous phase ratio) by volume.

In one embodiment of the present invention, the acidic stripping solution contains sulfuric acid.

In one embodiment of the present invention, the saturated extractant is contacted with a sufficient amount of acidic stripping solution to achieve the target pH in the stripped manganese solution after leaching. Preferably the target pH is less than 3.

In one embodiment of the present invention, water is added to the stripping stage to achieve a target concentration of manganese in the stripped manganese solution after leaching. Preferably, the target concentration is in the range of 100 g/L to 220 g/L.

Sequential solvent extraction circuits for manganese and cobalt.

As described above, cobalt and manganese can be co-extracted from the productive leach solution in a single solvent extraction step. In fig. 7a shows an alternative embodiment of the present invention in which manganese and cobalt are sequentially removed from the productive leach solution in separate solvent extraction steps. Reference numerals similar to those in FIG. 5 and 6 indicate similar elements. In this embodiment, the productive leach solution 502 from the aluminum solvent extraction circuit 400 is sent to the manganese recovery circuit 766 to extract manganese from the productive leach solution 502. The manganese-free raffinate 767 is sent to a cobalt recovery loop 768 to recover cobalt from the productive leach solution. Cobalt- and manganese-free raffinate 506 is sent to a holding tank for further processing.

The manganese solvent extraction circuit 766 includes an extraction stage 769, a scrubbing stage 770, and a stripping stage 771 for selectively extracting manganese from the productive leach solution 502 to produce a stripped manganese leach solution 772. Manganese-free raffinate 767 is sent to a holding tank for further processing.

In one embodiment of the present invention, extraction stage 769 includes one or more mixer-settlers for solvent extraction. More preferably, the extraction stage 769 includes two or more mixer-settlers for solvent extraction. Even more preferably, the extraction stage 769 includes three or more mixer-settlers for solvent extraction.

In one embodiment, if the extraction phase 769 contains two or more solvent extraction mixer-settlers, the two or more solvent extraction mixer-settlers are arranged in series.

In one embodiment, if extraction stage 769 includes two or more solvent extraction mixer-settlers arranged in series, the settler mixers operate in countercurrent.

In the embodiment shown in FIG. 7a, extraction stage 769 contains five solvent extraction mixer-settlers arranged in series. The inventors envision that one or more mixer-settlers may be used. The productive leach solution 502 is sent to the first mixer-settler where it is contacted with a manganese extractant 773 to selectively extract manganese from the productive leach solution 502.

In a preferred embodiment of the present invention, the manganese extractant 773 is an organic orthophosphoric acid extractant. Even more preferably, the manganese extractant is di-(2-ethylhexyl)orthophosphoric acid (D2EHPA).

The overall extraction reaction is presented in Equation 5a below:

2RH + MnSO ₄ = MnR ₂ + H ₂ SO ₄ (5a)

The productive leach solution 502 and the manganese extractant 773 are contacted in countercurrent to maximize extraction efficiency. Ammonia solution 774 is dosed into each of the mixer-settlers to maintain the target pH by neutralizing the sulfuric acid through the reaction shown in Equation 2.

Manganese is selectively removed from the productive leach solution 502 and the manganese-free raffinate 767 is sent to a holding tank for further processing.

The saturated extractant 775 is sent to the scrubbing stage 770. In the scrubbing step 770, the saturated extractant 775 is brought into contact with a portion of the wash solution 776. The manganese in the wash solution 776 predominantly saturates the saturated extractant 775 and displaces some of the impurity elements that saturate the saturated extractant 775 in the extraction step 769.

In one embodiment of the present invention, the scrubbing stage comprises one or more mixer-settlers.

In the embodiment shown in FIG. 7a, scrubbing stage 770 includes one mixer-settler.

The aqueous phase from the scrubbing stage 770 is sent back to the first mixer-settler for extraction. The saturated extractant 777 from the scrubbing stage 770 moves to the stripping stage 771. In a stripping step 771, the saturated extractant 777 is contacted with an acidic stripping solution 778 to displace the majority of the manganese ions in the organic phase (as opposed to the extraction reaction detailed in Equation 5) into the aqueous phase to produce a stripped manganese leach solution 772. A sufficient amount of water 779 is added to the stripping stage 771 to prevent crystallization of salts in the stripped manganese solution 772 after leaching.

The distilled manganese solution 772 after leaching is sent to the manganese crystallization stage 780.

In one embodiment, stripping stage 771 includes one or more mixer-settlers. Preferably, the stripping stage 771 includes two or more mixer-settlers.

In one embodiment, if stripping stage 771 contains two or more mixer-settlers, the two or more mixer-settlers are arranged in series.

The organic phase 773 leaving the stripping stage 771 is returned to the extraction stage 769 where it is again saturated with manganese. Thus, the organic phase 773 is kept in a closed loop in the manganese solvent extraction loop 766.

The distilled manganese solution 772, after leaching from the distillation stage 771, is sent to the manganese crystallization stage 780. In the manganese crystallization step 780, water is removed from the stripped manganese solution 772 after leaching to allow crystallization of the MnSO ₄ .H ₂ O in the slurry 782.

In one embodiment of the present invention, the manganese crystallization stage comprises a crystallization unit. Preferably, the crystallization unit is an evaporative crystallizer.

The suspension 782 from the manganese crystallization stage 780 is fed to the filtration stage 783. In a filtration stage 783, manganese sulfate solids 784 are separated from the filtrate. Water 785 is used in the filtration stage 783 to flush away filter cake and maximize manganese recovery. The filtrate is returned back to the stripping stage 771 or the manganese crystallization stage 780 to supplement the water used and maximize manganese recovery. A portion of the leachate may also be returned for use in the manganese scrubbing stage 769.

The cobalt solvent extraction circuit 768 includes an extraction stage 785, a scrubbing stage 786, and a stripping stage 787 for selectively extracting cobalt from the manganese-free raffinate 767 to produce a stripped cobalt solution 788 after leaching. The cobalt-free raffinate 506 is sent to a holding tank for further processing.

In one embodiment of the present invention, extraction stage 785 includes one or more mixer-settlers for solvent extraction. More preferably, extraction stage 785 includes two or more mixer-settlers for solvent extraction. Even more preferably, extraction stage 785 includes three or more mixer-settlers for solvent extraction.

In one embodiment, if the extraction phase 785 contains two or more solvent extraction mixer-settlers, the two or more solvent extraction mixer-settlers are arranged in series.

In one embodiment, if extraction stage 785 contains two or more solvent extraction mixer-settlers arranged in series, the settler-mixers operate in countercurrent.

In the embodiment shown in FIG. 7a, extraction stage 785 contains four solvent extraction mixer-settlers arranged in series. The inventors envision that one or more mixer-settlers may be used. The manganese-free raffinate 767 is sent to a first mixer-settler where it is contacted with a cobalt extractant 789 to selectively extract cobalt from the manganese-free raffinate 767.

In a preferred embodiment of the present invention, the cobalt extractant 789 is an organic extractant. More preferably, the cobalt extractant 789 is an organic phosphinic acid extractant. Even more preferably, the cobalt 789 extractant is Cyanex272.

The overall extraction reaction is presented in Equation 6a below:

2RH + CoSO ₄ = CoR ₃ + H ₂ SO ₄ (6a)

Manganese-free raffinate 767 and cobalt extractant 789 are brought into contact in countercurrent

- 21 045218 to maximize extraction efficiency. Ammonia solution 790 is dosed into each of the mixer-settlers to maintain the target pH by neutralizing the sulfuric acid through the reaction shown in Equation 2.

Cobalt is selectively recovered from the manganese-free raffinate 767, and the cobalt-free raffinate 767 is sent to a holding tank for further processing.

The saturated extractant 791 is sent to the scrubbing stage 786. In scrubbing step 786, the saturated extractant 791 is contacted with a portion of the wash solution 792. The rinse solution 792 displaces some of the impurity elements that were saturated with the saturated extractant 791 in the extraction step 785.

In one embodiment of the present invention, the scrubbing stage comprises one or more mixer-settlers.

In the embodiment shown in FIG. 7a, scrubbing stage 786 includes one mixer-settler.

The aqueous phase from the scrubbing stage 786 is sent back to the first mixer-settler for extraction. The saturated extractant 793 from the scrubbing stage 786 moves to the stripping stage 787. In a stripping step 787, the saturated extractant 793 is contacted with an acidic stripping solution 794 to displace the majority of the cobalt ions in the organic phase (as opposed to the extraction reaction detailed in Equation 3) into the aqueous phase to produce a stripped cobalt leach solution 788. A sufficient amount of water 795 is added to the stripping stage 787 to prevent crystallization of salts in the stripped cobalt solution 788 after leaching.

The distilled cobalt solution 788 after leaching is sent to the cobalt crystallization stage 796.

In one embodiment, stripping stage 787 includes one or more mixer-settlers. Preferably, the stripping stage 787 includes two or more mixer-settlers.

In one embodiment, if stripping stage 787 contains two or more mixer-settlers, the two or more mixer-settlers are arranged in series.

The organic phase 789 leaving the stripping stage 787 is returned to the extraction stage 785 where it is again saturated with cobalt. Thus, the organic phase 789 is kept in a closed loop in the cobalt solvent extraction loop 768.

The distilled cobalt solution 788, after leaching from the distillation stage 787, is sent to the cobalt crystallization stage 796. In the cobalt crystallization step 796, water is removed from the stripped cobalt solution 788 after leaching to allow crystallization of CoSO ₄ .7H ₂ O in slurry 798.

In one embodiment of the present invention, the cobalt crystallization stage 796 includes a crystallization unit. Preferably, the crystallization unit is an evaporative crystallizer.

The suspension 798 from the cobalt crystallization stage 796 is fed to the filtration stage 799. In a filtration step 799, cobalt sulfate solids 7701 are separated from the filtrate. Water is applied to the filtration stage 799 to wash away filter cake and maximize cobalt recovery. The filtrate is returned back to the cobalt crystallization stage 796 to maximize cobalt recovery. A portion of the filtrate may also be returned for use in the cobalt scrubbing stage 785 or returned upstream of the cobalt solvent extraction loop.

Operating conditions of the manganese solvent extraction circuit.

In one embodiment of the present invention, the manganese extractant is diluted to a target concentration using a diluent. Preferably, the target concentration is in the range of 15 vol.% to 35 vol.%. More preferably, the target concentration is approximately 30 vol%. As will be appreciated by one skilled in the art, dilution of the manganese extractant is used to adjust the viscosity.

In embodiments of the present invention, if the manganese extractant is diluted with a diluent, the diluent is Shellsol2046. Alternatively, other diluents known to those skilled in the art may be used, for example Exsol D80.

Preferably, the target pH of the productive leach solution in the extraction step is in the range of 3 to 4.5. More preferably, the target pH of the productive leach solution in the extraction step is 3.5.

In a preferred embodiment of the present invention, the ratio of organic and aqueous phases in the extraction step is 1.65:1. As will be appreciated by one skilled in the art, the ratio of organic to aqueous phases in the extraction step depends on the manganese content of the productive leach solution, as well as the manganese saturation of the organic phase. One way to calculate the equilibrium concentration of an extractant is to use a material balance, i.e. it is equal to the difference between the total (analytical) concentration of the extractant and the sum of all solvated substances in the soluble phase.

- 22 045218

In one embodiment of the present invention, the ratio of saturated manganese extractant to wash solution is in the range of 20:1 to 100:1 (ratio of organic to aqueous phases) by volume. Preferably the ratio is in the range of 30:1 to 70:1 (organic to aqueous phase ratio) by volume. More preferably, the ratio is approximately 50:1 (organic to aqueous phase ratio) by volume.

In one embodiment of the present invention, the acidic stripping solution contains sulfuric acid.

In one embodiment of the present invention, the saturated extractant is contacted with a sufficient amount of acidic stripping solution to achieve the target pH in the stripped manganese solution after leaching. Preferably the target pH is less than 3.

In one embodiment of the present invention, water is added to the stripping stage to achieve a target concentration of manganese in the stripped manganese solution after leaching. Preferably, the target concentration is in the range of 100 g/L to 220 g/L.

Operating conditions of the cobalt solvent extraction circuit.

In one embodiment of the present invention, the cobalt extractant is diluted to a target concentration using a diluent. Preferably, the target concentration is in the range of 15 vol.% to 35 vol.%. More preferably, the target concentration is approximately 30 vol%. As one skilled in the art will appreciate, dilution of the cobalt extractant is used to adjust the viscosity.

In embodiments of the present invention, if the cobalt extractant is diluted with a diluent, the diluent is Shellsol2046.

In a preferred embodiment of the present invention, the ratio of organic and aqueous phases in the extraction step is 3.1:1. As will be appreciated by one skilled in the art, the ratio of organic to aqueous phases in the extraction step depends on the cobalt/manganese content of the productive leach solution, as well as the cobalt/manganese saturation of the organic phase. One way to calculate the equilibrium concentration of an extractant is to use a material balance, i.e. it is equal to the difference between the total (analytical) concentration of the extractant and the sum of all solvated substances in the soluble phase.

Preferably, the target pH of the productive leach solution in the extraction stage is in the range of 4.2 to 5. More preferably, the target pH of the productive leach solution in the extraction stage is 4.8.

In one embodiment of the present invention, the ratio of saturated cobalt extractant to wash solution is in the range of 20:1 to 100:1 (organic to aqueous phase ratio) by volume. Preferably the ratio is in the range of 30:1 to 70:1 (organic to aqueous phase ratio) by volume. More preferably, the ratio is approximately 50:1 (organic to aqueous phase ratio) by volume.

In one embodiment of the present invention, the acidic stripping solution contains sulfuric acid.

In one embodiment of the present invention, the saturated extractant is contacted with a sufficient amount of acidic stripping solution to achieve the target pH in the stripped cobalt solution after leaching. Preferably the target pH is less than 3.

In one embodiment of the present invention, water is added to the stripping stage to achieve a target cobalt concentration in the stripped cobalt solution after leaching. Preferably, the target concentration is in the range of 80 g/L to 140 g/L.

Solvent nickel extraction circuit.

After the step of separating the productive leach solution from the solid residue, and preferably after the copper, aluminum and cobalt/manganese extraction steps, the productive leach solution is fed to a nickel solvent extraction circuit to recover the nickel. In fig. 8 shows a flow diagram of a nickel solvent extraction circuit 800 in accordance with one embodiment of the present invention. The productive leach solution 802 from the cobalt/manganese solvent extraction circuit is sent to the nickel solvent extraction circuit 800 . In the embodiment shown in FIG. 8, the nickel solvent extraction circuit 800 includes an extraction stage 804, a scrubbing stage 806, and a stripping stage 808 for selectively extracting nickel from the productive leach solution 802 to produce a stripped nickel leach solution 810. Nickel-free raffinate 812, substantially free of other metals at this stage, is sent to a holding tank for further processing.

In one embodiment of the present invention, the extraction stage includes one or more mixer-settlers for solvent extraction. More preferably, the extraction stage comprises two or more solvent extraction mixer-settlers arranged in series.

In one embodiment, if the extraction stage contains two or more solvent extraction mixer-settlers arranged in series, the mixer-settlers operate in countercurrent.

In the embodiment shown in FIG. 8, the extraction stage 804 contains three solvent extraction mixers-23 045218 settlers arranged in series. The inventors envision that one or more mixer-settlers may be used. Productive solution

802 leaching is sent to the first mixer-settler, where it is brought into contact with the extractant

814 nickel for selective extraction of nickel from the productive leaching solution 802.

In a preferred embodiment of the present invention, the nickel extractant is an organic extractant. More preferably, the nickel extractant is an oxime. Even more preferably, the nickel extractant is LIX84I. Alternatively, the nickel extractant is a carboxylic acid such as Versatic 10.

The overall extraction reaction is presented in Equation 7 below:

2RH + NiSO ₄ = NiR ₂ + H ₂ SO ₄ (7)

The productive leach solution 802 and the nickel extractant 814 are contacted in countercurrent to maximize extraction efficiency. Ammonia solution 816 is dosed into each of the mixer-settlers to maintain the target pH by neutralizing the sulfuric acid through the reaction shown in Equation 2.

Nickel is selectively removed from the productive leach solution 802, and the nickel-free raffinate 812 is sent to a holding tank for further processing.

The saturated extractant 818 is sent to the scrubbing stage 806. In scrubbing step 806, the rich nickel extractant 818 is contacted with a portion of the wash solution 820 that contains nickel ions. Nickel ions in the wash solution 820 preferentially saturate the saturated extractant 818 and displace the impurity elements that saturate the saturated extractant 818 in the extraction step 804.

In one embodiment of the present invention, the scrubbing stage comprises one or more mixer-settlers.

In the embodiment shown in FIG. 8, the scrubber cleaning stage contains one mixer-settler.

The aqueous phase from the scrubbing stage 806 is sent back to the first mixer-settler for extraction. The saturated extractant 822 from the scrubbing stage 806 moves to the stripping stage 808. In stripping step 808, the saturated extractant 822 is contacted with an acidic stripping solution 824 to displace the majority of the nickel ions in the organic phase (as opposed to the extraction reaction detailed in Equation 7) into the aqueous phase to produce a stripped nickel leach solution 810. A sufficient amount of water 825 is added to the stripping stage 808 to prevent crystallization of salts in the stripped nickel solution 810 after leaching.

In one embodiment, the stripping stage 808 includes one or more mixer settlers. Preferably, the stripping stage 808 includes two or more mixer-settlers.

In one embodiment, if stripping stage 808 contains two or more mixer-settlers, the two or more mixer-settlers are arranged in series.

The organic phase 826 leaving the stripping stage 808 is returned to the extraction stage 804 where it is again saturated with nickel. Thus, the organic phase 826 is kept in a closed loop in the nickel solvent extraction loop 800.

The stripped nickel solution 810, after leaching from the stripping stage 808, is sent to the nickel crystallization stage 828.

In the nickel crystallization step 828, water is removed from the stripped nickel solution 810 after leaching to allow NiSO ₄ .6H ₂ O to crystallize in slurry 830.

In one embodiment of the present invention, the nickel crystallization stage 828 includes a crystallization unit. Preferably, the crystallization unit is an evaporative crystallizer.

The suspension 830 from the nickel crystallization stage 828 is fed to the filtration stage 832. In a filtration stage 832, nickel sulfate solids 834 are separated from the filtrate. Water 840 is used in the filtration stage 832 to flush away filter cake and maximize nickel recovery. The filtrate is returned back to the nickel crystallization stage 828 to maximize nickel recovery. A portion of the filtrate is returned for use in scrubbing stage 806. The nickel sulfate solids 834 are transported to the nickel sulfate product storage area.

Operating conditions of the nickel solvent extraction circuit.

In one embodiment of the present invention, the nickel extractant is diluted to a target concentration using a diluent. Preferably, the target concentration is in the range of 5% by volume to 35% by volume. More preferably, the target concentration is approximately 10 vol%. As one skilled in the art will appreciate, dilution of the nickel extractant is used to adjust the viscosity.

In embodiments of the present invention, if the nickel extractant is diluted with a diluent, the diluent is Shellsol2046.

- 24 045218

In a preferred embodiment of the present invention, the ratio of organic and aqueous phases in the extraction step is 3.55:1. As will be appreciated by one skilled in the art, the ratio of organic to aqueous phases in the extraction step depends on the nickel content of the productive leach solution, as well as the nickel saturation of the organic phase. One way to calculate the equilibrium concentration of an extractant is to use a material balance, i.e. it is equal to the difference between the total (analytical) concentration of the extractant and the sum of all solvated substances in the soluble phase.

Preferably, the target pH of the productive leach solution in the extraction step is at least 6.4. More preferably, the target pH of the productive leach solution in the extraction step is in the range of 5 to 7.

In one embodiment of the present invention, the ratio of saturated nickel extractant to wash solution is in the range of 20:1 to 100:1 (organic to aqueous phase ratio) by volume. Preferably the ratio is in the range of 30:1 to 70:1 (organic to aqueous phase ratio) by volume. More preferably, the ratio is approximately 50:1 (organic to aqueous phase ratio) by volume.

In one embodiment of the present invention, the acidic stripping solution contains sulfuric acid.

In one embodiment of the present invention, the saturated extractant is contacted with a sufficient amount of acidic stripping solution to achieve the target pH in the stripped manganese solution after leaching. Preferably the target pH is less than 4.

In one embodiment of the present invention, water is added to the stripping stage to achieve a target nickel concentration in the stripped manganese solution after leaching. Preferably, the target concentration is 50 g/L and 120 g/L.

Lithium extraction circuit.

After the step of separating the productive leach solution from the solid residue, and preferably after the copper, aluminum, cobalt/manganese and nickel extraction steps, the productive leach solution is fed to a lithium recovery loop to recover the lithium. In fig. 9 shows a flow diagram of a lithium solvent extraction circuit 900 in accordance with one embodiment of the present invention. The productive leach solution 902 from the nickel solvent extraction circuit is sent to the lithium solvent extraction circuit 900. In the embodiment shown in FIG. 9, lithium solvent extraction circuit 900 includes an evaporation stage 904, a precipitation stage 906, and a filtration stage 908 for selectively extracting lithium from the productive leach solution 902 to produce solid lithium 910. The lithium-free filtrate 912, at this point substantially free of metals, is sent into a holding tank for further processing.

The productive leach solution 902 is sent to an evaporation stage 904 where water is removed from the productive leach solution 902 to produce a concentrated lithium leach solution 914. The purpose of water removal is to maximize the lithium concentration in the concentrated lithium leach solution 914 to a level just below the solubility limit of the salts in the concentrated lithium leach solution 914.

In one embodiment of the present invention, the evaporation stage 904 includes an evaporation unit. Preferably, the evaporation unit is a falling film evaporator. Energy is introduced through the use of steam in an indirect heat exchanger. Evaporation is carried out at high temperature and moderate vacuum to promote efficient water removal.

The concentrated lithium solution 914, after leaching, is sent to a precipitation stage 906 where it is contacted with ammonium carbonate 916 to precipitate lithium carbonate into a lithium carbonate slurry 918. Precipitation of lithium carbonate occurs according to the reaction in equation 8.

L12SO4 (aq.) + (aq.) — L12CO3 (sol.) + (NH4)2SO4 (aq.) (8)

Preferably, the precipitation step comprises one or more reaction vessels.

Preferably, the lithium carbonate slurry 918 is maintained at an elevated temperature. The inventors have discovered that by performing the deposition step at elevated temperatures, deposition efficiency is maximized by exploiting the inverse temperature-solubility profile of lithium carbonate.

The lithium carbonate slurry 918 is supplied to a filtration stage 908 to separate solid lithium carbonate 910 from the lithium carbonate slurry 918.

Preferably, the filtration stage includes a device for separating solid and liquid phases. More preferably, the solid-liquid separation device is a centrifuge.

Fresh water 920 is applied to the filtration stage 908 to flush out the cake and maximize the removal of trapped solution from the cake leaching. The lithium-free filtrate 912 containing ammonia is sent to a holding tank for further processing.

A precipitation stage 906 and a filtration stage 908 are used to selectively recover lithium from the productive leach solution 902 to produce solid lithium 910.

- 25 045218 lithium filtrate 912, at this stage essentially free of metals, is sent to a holding tank for further processing.

Operating conditions of the lithium extraction circuit.

In a preferred embodiment of the present invention, the lithium concentration in the concentrated lithium solution after leaching is set to 300 g/L Li ₂ SO ₄ .

Preferably, the lithium carbonate slurry is maintained at a temperature of at least 90°C.

Alternative means of lithium extraction.

It is contemplated that other means for recovering lithium may be used. Suitable means available to those skilled in the art include precipitation of other lithium-containing compounds, solvent extraction, ion exchange and electrolysis.

Solvent extraction of lithium.

If solvent extraction is used to recover lithium, the productive leach solution may be contacted with a lithium extractant to recover the lithium. The lithium extraction circuit is provided to include an extraction stage, a scrubbing stage, and a stripping stage for selectively extracting lithium from the productive leach solution to produce a stripped lithium leach solution.

In one embodiment of the present invention, the productive leach solution is subjected to an impurity removal step prior to the lithium solvent extraction step. It is contemplated that the productive leach solution may contain impurities such as Ca, Mg, K, Na, and possibly other alkali metals and alkaline earth metals. These impurities are preferably removed to prevent them from saturating the lithium extractant. In one embodiment of the present invention, the productive leach solution is subjected to an ion exchange step to remove multivalent ions from the productive leach solution. A number of suitable ion exchange resins are known to those skilled in the art, including Dowex G26 resin. Alternatively, these impurities may be precipitated from the productive leach solution.

It is contemplated that a number of commercially available extractants are available to extract lithium from the productive leach solution. In a preferred embodiment of the present invention, the lithium extractant is an organic extractant. In one embodiment, the lithium extractant contains an organic phosphorus-based extractant. Preferably, the lithium extractant is selected from Cyanex 936, Cyanex 923 and Cyanex 926. In one embodiment, the lithium extractant contains an oxime. Preferably, the oxime is selected from LIX 54, Mextral 54-100 and LIX 55. In one embodiment, the lithium extractant comprises a mixture of two or more organic phosphorus-based extractants and oxime-based extractants.

The ammonia solution is dosed into each of the mixer-settlers to maintain the target pH. Preferably, the target pH is greater than 9. Even more preferably, the target pH is between 9 and 11.

Preferably, the stripping step involves contacting the saturated organic extractant with an acidic stripping solution. In one embodiment of the present invention, the acidic stripping solution contains sulfuric acid. In an alternative embodiment of the present invention, the acidic stripping solution is hydrochloric acid. The use of hydrochloric acid results in a distilled lithium chloride solution after leaching. It is contemplated that LiCl may be crystallized from this distilled solution after leaching.

Other known processes for recovering lithium from the leach solution through solvent extraction processes known to those skilled in the art may be used without departing from the scope of the present invention.

Ion exchange involving lithium ions.

If ion exchange is used to recover lithium from the productive leach solution, the method includes the step of contacting the productive leach solution with an ion exchange resin to selectively extract lithium from the productive leach solution. The method further includes contacting the saturated ion exchange resin with the stripping solution to recover lithium from the stripped lithium solution.

Lithium electrolysis.

If lithium is recovered from the leach solution by electrolysis, the method is contemplated to include the step of sending the productive leach solution to an electrolytic cell to recover the lithium. Preferably, the productive leach solution is subjected to an impurity removal step to remove dissolved metallic species. These metallic species include, for example, Ca, Mg and other transition metal ions.

In one embodiment of the present invention, lithium is recovered in the electrolyzer in the form of lithium hydroxide. Preferably, the electrolyzer is an electrodialyzer (membrane electrolyzer) or a two-chamber electrolyzer. In one embodiment, the pH of the productive leach solution is maintained in the range of 9.5 to 12.5 in the electrolyzer.

In one embodiment of the present invention, lithium is recovered in the electrolyzer as metal

- 26 045218 lithium. Preferably, the electrolyser is a molten salt electrolyser. More preferably, the electrolyzer matrix is a chloride salt.

Other electrolytic processes known to those skilled in the art may be used to recover lithium or lithium products from the leach solution without departing from the scope of the present invention.

Precipitation of lithium-containing compounds.

It is contemplated that the lithium products may be precipitated and separated from the productive leach solution by the addition of a precipitating agent.

In one embodiment of the present invention, the precipitating agent is a carbonate, for example sodium carbonate. As one skilled in the art will appreciate, the addition of carbonate results in the precipitation of lithium carbonate. The recovered sludge can also be sent to a treatment loop to remove impurities. In one embodiment of the present invention, solid lithium phosphate may be precipitated from the leach solution. Preferably, phosphate material is added to the leach solution to precipitate lithium phosphate. More preferably, the phosphate material is lithium phosphate.

Other means for precipitating and recovering lithium products known to those skilled in the art may be used without departing from the scope of the present invention.

Ammonium sulfate recovery circuit.

In certain embodiments of the present invention, separating lithium from the productive leach solution produces a filtrate containing ammonia. This ammonia is recovered as ammonium sulfate in the ammonium sulfate recovery loop. In fig. 10 shows a flow diagram of an ammonium sulfate recovery circuit 1000 in accordance with one embodiment of the present invention. The filtrate 912 from the lithium recovery circuit is sent to the ammonium sulfate recovery circuit 1000 . In the embodiment shown in FIG. 10, the ammonium sulfate recovery circuit 1000 includes a crystallization stage 1004 and a filtration stage 1006 for selectively recovering ammonium sulfate from the productive leach solution 902 to produce solid ammonium sulfate 1008.

The filtrate 902 is sent to the crystallization stage 1004. In crystallization step 1004, water is removed from filtrate 912 to allow ammonia to crystallize as ammonium sulfate (NH ₄ ) ₂ SO ₄ in slurry 1012. Water removal is designed to maximize solid recovery of ammonium sulfate. Any compounds still remaining in the lithium filtrate (ie NiSO ₄ , Li ₂ CO _{3 ,} etc.) co-crystallize as impurities.

In one embodiment of the present invention, the ammonium sulfate crystallization stage 1004 includes a crystallization unit. Preferably, the crystallization unit is an evaporative crystallizer. Even more preferably, the evaporative crystallizer is a forced circulation evaporative crystallizer.

The suspension 1012 from the crystallization stage 1004 is fed to the filtration stage 1006. In filtration stage 1006, precipitated ammonium sulfate solids are separated from filtrate 1014. Water 1016 is used in filtration stage 1006 to wash away filter cake and maximize ammonium sulfate recovery. The filtrate 1014 is returned back to the crystallization stage 1004 to maximize ammonium sulfate recovery. The recovered solid ammonium sulfate 1008 is discharged from the filter press and supplied to the ammonium sulfate product storage area.

Example 1.

The lithium-ion laptop battery was disassembled, and the composition of materials is shown in Table. 1.

Table 1

Battery composition

SAMPLE Weight (kg) Content (%) Total 7,468 100.00% Share Share of plastic waste 0.788 10.56% Battery cell share 6,459 86.48% Share of electrical circuits 0.221 2.96%

An analysis of the chemical composition of the battery was carried out, the results of which are given in table. 2.

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table 2

Chemical analysis

SAMPLE _Li2O CoO _{Fe2O3 _} A1 ₂ O _{3 SiO2} TY ₂ MP S R _SnO2 ppm ppm ppm ppm ppm ppm ppm ppm ppm ppm Composition limit 56405 268790 240620 101390 550 120 605 970 2250 5.0 SAMPLE Ta ₂ O _{5 Nb2Os Na2O} PHO SaO MgO K ₂ O NiO C&O Rb ppm ppm ppm ppm ppm ppm ppm ppm ppm ppm Composition limit <0.5 5 <5 96 1255 910 <5 19771 124766 <0.5

Example 2.

Three separate leaches were applied to the battery samples of Example 1 to analyze the effect that different leaching conditions had on the dissolution of metals in the battery.

The battery sample was leached with H ₂ SO ₄ with SO ₂ gas sprayed through the leach vessel. The parameters for leaching are shown in table. 3.

Table 3

Leaching parameters

Size: 10 mm P100 Initial weight: 57.1 G Pulp Density: 20.0 % (v/v) H2SO4 concentration: thirty % (v/v) Leaching time: 2.0 watch Temperature leaching: 100 °C Speed splash (SO2) 0.02 l/min

The leaching results are shown in table. 4 and table. 5.

Table 4

Leaching results

Solid phase Weight Li Co Fe A1 Ni Si G ppm ppm ppm ppm ppm ppm Remainder after leaching 74.74 7269 147122 635 10795 5210 70767 Solution Volume Li Co Fe A1 Ni Si ml mg/l mg/l mg/l mg/l mg/l mg/l Solution after leaching 24.52957 19460 27730 thirty 27570 2760 18080 Washed Part 249.1459 3382 12098 24 4732 585 3793 Calculated limit 57.14 32608.36 257092.6 948.1149 46588.33 10550.36 116864.4

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Table 5

Leaching Distribution

Solid phase Weight Li Co Fe A1 Ni Si G diets diets diets diets diets diets Residue after leaching 74.74 0.29158 1 0.74851 6 0.87604 3 0.30308 1 0.64592 7 0.79206 6 Solution Volume Li Co Fe A1 Ni Si ml diets diets diets diets diets diets Solution after leaching 24.5295 7 0.25619 1 0.04630 0.01358 3 0.25404 4 0.11230 3 0.06641 5 Washed 249.145 0.45222 0.20518 0.11037 0.44287 0.24177 0.14151 Part 9 9 1 3 5 9 Calculated limit 57.14 1 1 1 1 1 1

The battery sample was leached with H ₂ SO ₄ . The parameters for leaching are shown in table. 6.

Table 6

Leaching parameters

Size: 10 mm P100 Initial weight: 47.6 G Pulp Density: 12.5 % (v/v) H2SO4 concentration: 20 % (v/v) Leaching time: 4.0 watch Leaching temperature: 100 °C

The leaching results are shown in table. 7 and table. 8.

Table 7 _________________________Leaching results_______________________

Solid phase Weight Li Co Fe A1 Ni Si G ppm ppm ppm ppm ppm ppm Remainder after leaching 13.42 40 499 40 885 125 551 Solution Volume Li Co Fe A1 Ni Si ml mg/l mg/l mg/l mg/l mg/l mg/l Solution after leaching 195.0 6152 47095 141 7662 1292 18149 T=120 min S/S 15.2 6062 46574 137 7562 1279.00 17740.00 Washed Part 366.7 463 3625 eleven 576 99.00 1413.00 Calculated limit 47.55 30750 236126 718 38532 6506 91154.54

- 29 045218

Leaching Distribution

Table 8

Solid phase Weight Li Co Fe A1 Ni Si G diets diets diets diets diets diets Remainder after leaching 13.4 0.04% 0.06% 1.57% 0.65% 0.54% 0.17% Solution Volume Li Co Fe A1 Ni Si ml diets diets diets diets diets diets Solution after leaching T=120 min S/S Washed part 195.0 15.2 366.7 82.07% 6.29% 11.61% 81.81% 6.29% 11.84% 80.53% 6.08% 11.81% 81.57% 6.26% 11.53% 81.45% 6.27% 11.73% 81.67% 6.21% 11.95% Calculated limit 47.55 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

The battery sample was leached with H2SO4, with 1 molar hydrogen peroxide added dropwise until the ORP (oxidation-reduction potential) stabilized. The parameters for leaching are shown in table. 9.

Table 9

Leaching parameters

Size: 10 mm P100 Initial weight: 51.6 G Pulp Density: 20.0 %(v/v) H2SO4 concentration: thirty %(v/v) Leaching time: 2.0 watch Temperature leaching: 80 °C Spray rate (SO2) 51.6 G Supported ORP: >300 mV (Ag/AgCl) Total added IM H2O2: 157.71 G

The leaching results are shown in table. 10 and table. eleven.

Table 10

Leaching results

Solid phase Weight Li Co Fe A1 Ni Si G ppm ppm ppm ppm ppm ppm Remainder after leaching 17.3 4037 65674 1045 12605 1271 66480 Solution Volume Li Co Fe A1 Ni Si ml mg/l mg/l mg/l mg/l mg/l mg/l Solution after leaching 253.0 5704 42328 84 8381 1362 16790 Washed 340.2 427 3670 6 652 109.00 1616.00

- 30 045218

Part Calculated limit (ppm) 51.55 32163 253970 802 49660 7829 115366.24

Table 11

Leaching Distribution

Solid phase Weight Li Co Fe A1 Ni Si G diets diets diets diets diets diets Residue after leaching 17.3 4.21% 8.68% 43.70% 8.52% 5.45% 19.34% Solution Volume Li Co Fe A1 Ni Si ml diets diets diets diets diets diets Solution after leaching 253.0 87.03% 81.78% 51.36% 82.82% 85.36% 71.42% Washed Part 340.2 8.76% 9.54% 4.93% 8.67% 9.19% 9.25% Calculated limit 51.55 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 253.0 87.03% 81.78% 51.36% 82.82% 85.36% 71.42%

Example 3.

Copper extraction.

A series of tests were carried out to determine the conditions necessary to extract copper from the productive leach solution. The primary purpose of solvent copper extraction was to maximize copper recovery to produce a raffinate leach solution with low copper content (<10 ppm Cu).

Due to the high concentration of copper in the solution, a commercially available high concentration extractant was used as the organic phase. 30% (v/v) LIX84I in Exxsol D80 was selected. The characteristics of LIX reagents are well known in the art, in particular the stripping characteristics.

A distribution isotherm test was performed on the leach solution obtained from a pilot leach plant to determine the number of copper extraction stages. The solution after leaching contained 28.3 g/l Co, 6.24 g/l Cu, 5.09 g/l Ni, 6.56 g/l Al, 1.42 g/l Fe, 4.14 g/l Li and 3.39 g/l Mn. Distribution isotherm testing was carried out at pH 2.0 and 40°C to prevent crystallization of ammonia alum. The distribution isotherm and McCabe-Thiele diagram are presented in Fig. eleven.

The results show that >99% Cu can be extracted in two steps using an O/A ratio of 1.1:1 at pH 2.0. The raffinates contained <10 ppm Cu. The results show excellent separation of copper from other metals in the artificial solution after leaching. A minority of the cobalt is co-extracted, which can be problematic because cobalt is not effectively stripped from the saturated solvent under copper stripping conditions. Less co-extraction of cobalt is expected at lower pH. Stripping copper from the saturated organic phase is a common process and can be expected to be accomplished by those skilled in the art.

Copper sulfate crystallization tests were carried out to determine the initial conditions required for copper sulfate recovery. In the first test, an artificial solution containing 75 g/l Cu (as copper sulfate) and 140 g/l H2SO4 was used. In the second test, an artificial solution containing 75 g/l Cu (as copper sulfate) and 125 g/l H2SO4 was used. The test solutions were cooled from 40°C to 20°C after leaching and then sulfuric acid was added to account for the sulfate contained in the copper sulfate. The results are presented in Fig. 12 and 13 respectively.

For both cases, the copper concentration was reduced by 25-30 g/L Cu, with copper crystallization being more dependent on temperature than on acid concentration. The copper sulfate product was recovered by filtering and washing batches of the product. Product batches were dried at 50°C. The product contained 25.5% Cu, 27 ppm Fe, 25 ppm Al and <3 ppm Co, <6 ppm Ni and <20 ppm Li.

- 31 045218

Example 4.

Three precipitation tests were carried out to determine the conditions necessary to precipitate aluminum and iron from solution while minimizing co-precipitation of cobalt, nickel and lithium. Raffinate copper recovered from the bulk copper solvent extraction test was used for both tests. The leaching solution contained 25.2 g/L Co, 4.38 g/L Ni, 7.70 g/L Al, 1.27 g/L Fe, 3.58 g/L Li and 3.08 g/L m. Tests included an initial addition of hydrogen peroxide to increase Eh to >600 mV for iron oxidation and a 28% (v/v) addition of ammonia to adjust the pH. Tests were carried out at different temperatures: 45°C, 65°C and 90°C. The results are presented in Fig. 14, 15 and 16 respectively.

The results show that >85% iron and >90% aluminum are effectively precipitated from solution at pH >4.50. Co-precipitation of cobalt, nickel and lithium increases with increasing pH. At 45°C, cobalt coprecipitation is <5% at pH less than 3.75 and >10% at pH greater than 4.75.

The effect of temperature on the separation of aluminum from cobalt and nickel is presented in Fig. 17 and 18 respectively.

The results show that aluminum precipitation occurs at significantly lower pH at 65°C and 90°C compared to 45°C. At 45°C the pH for high aluminum precipitation must be greater than 4.50, at which cobalt co-precipitation is high (>8%). In comparison, at 65°C and 90°C, aluminum deposition efficiency was 95% at pH 4.0, while cobalt co-precipitation was <2%. Similar results were obtained for nickel. Nickel co-precipitation was 11% at pH 4.00 at both 65°C and 90°C. Nickel losses were higher (>20%) when targeting high aluminum deposition efficiency at a lower temperature.

Data indicate that iron and aluminum deposition can be achieved with minimal cobalt loss by using a countercurrent deposition circuit. This may include primary and secondary deposition circuits, including the following:

the primary neutralization (PN) circuit treats the raffinate obtained by solvent extraction of copper at pH 3.2-4.0 at 65-90°C to precipitate most of the iron and aluminum by adding hydrogen peroxide or air and ammonia solution;

the PN output is concentrated, filtered and washed, and the residue is disposed of; and the solution after leaching from PN is transferred to the secondary neutralization stage (SN). The SN circuit treats the PN leaching solution at pH 4.5-6.0 at 45-65°C to precipitate remaining aluminum and iron by adding hydrogen peroxide or air and ammonia.

The SN outlet is thickened and the thickener overhead is transferred to the cobalt recovery loop. The underflow from the thickener is pumped upstream to any leach loop, a small separate leach loop, or fed to the PN loop to recover co-precipitated cobalt, nickel and lithium.

Example 5.

A pH isotherm was performed on the post-precipitation leachate to evaluate the selectivity of D2EHPA for manganese/cobalt separation. The pH isotherm was obtained using an organic phase containing 15% D2EHPA in Exxsol D80 at an O/A ratio of 1:1. The isotherm is shown in Fig. 19.

The pH isotherm shows that manganese was extracted more selectively over cobalt using D2EHPA. At pH 3.5, manganese extraction is 75% and cobalt extraction is 12%.

A pH isotherm test was performed on the post-sedimentation leachate to evaluate the selectivity of Cyanex 272 for manganese and cobalt/nickel separation. The pH isotherm was obtained using an organic phase containing 30% Cyanex 272 in Exxsol D80 at an O/A ratio of 3:1. The isotherm is shown in Fig. 20.

The pH isotherm shows that the extraction of manganese and cobalt was carried out more selectively relative to nickel and magnesium using Cyanex 272. At pH 5.0, the extraction of manganese and cobalt is 92% and 78%, respectively, while the extraction of magnesium and nickel is only 19% and < 1% respectively.

A distribution isotherm test was performed on the post-sedimentation leachate to estimate the number of steps required to extract manganese and cobalt from solution. The test was carried out using 30% Cyanex 272 in Exxsol D80 at pH 4.80. Isotherms for cobalt and manganese are presented in Fig. 21 and 22 respectively.

The isotherms are steep and show high extraction of cobalt and manganese at 3-4 stages.

Example 6.

A pH isotherm was performed on the raffinate obtained by solvent extraction of cobalt to evaluate the selectivity of LIX84I for nickel/lithium separation. The pH isotherm was obtained using an organic phase containing 10% LIX84I in Exxsol D80 at an O/A ratio of 1:1 at 55°C. The isotherm is shown in Fig. 23.

The pH isotherm shows that nickel extraction was more selective relative to cobalt and lithium using LIX84I. At pH 5.0, nickel extraction is 84%, and cobalt extraction is only 10%. Lithium extraction is <1%. Excellent separation of nickel from lithium is expected with countercurrent solvent extraction.

Example 7.

A test leaching cycle was carried out in which 980 kg of raw material was continuously processed for 116 hours. The test setup consisted of six (6) stirred leach tanks operating at atmospheric pressure and 80°C. The leach feed, pulped, was fed to leach tank 1 (L1) via a gravity feeder, and 96% sulfuric acid and water were pumped into leach tank 1 with a fixed flow control to maintain the target solids percentage of 12.8%, and an acid addition rate of 1100 kg/t.

The extraction of cobalt, nickel and lithium during the cycle was high, reaching 99.8, 99.5 and 99.5%, respectively. Copper extraction was much lower at 80.9%. This was due to low Eh in the circuit during the first 72 hours of operation. Eh was consistently <100 mV during this period. During this period, leach residue quality averaged 10.5% Cu and copper extraction averaged 73%.

Starting 72 hours before the end of the cycle, hydrogen peroxide was added to Leach Reactor 2 at an average flow rate of 63 g/min. This led to a significant increase in copper extraction. During the remaining operating period, copper extraction exceeded 97.5%. A total of 5.8 m3 of pooled leach solution containing 27.4 g/L Co, 4.9 g/L Ni, 8.4 g/L Cu and 4.0 g/L Li was recovered.

Example 8.

A copper solvent extraction and crystallization test cycle was conducted in which approximately 2000 L of leach solution removed from the leach test facility was processed for 136 hours. Copper recovery involved solvent extraction of copper from the productive leach solution and crystallization of copper sulfate from the saturated distilled leach solution. The copper solvent extraction and crystallization circuits were operated in a closed loop in which the copper crystallization filtrate was used as the stripped solution after leaching.

The copper solvent extraction test setup consisted of 3 extraction stages, 1 washing stage and 2 stripping stages. Ammonia solution was used as pH in the extraction stages and sulfuric acid was used in the scrubbing and stripping stages.

The concentration of copper in the productive leach solution and SX copper raffinate over the cycle duration is presented in FIG. 24. Copper extraction averaged 99.6% over the course of the cycle and was consistently >99.5% after 25 hours of operation and for the remainder of the operation. Raffinate copper concentrations averaged 12 mg/L during this period.

The copper sulfate crystallization circuit consisted of an intermediate tank for the saturated distilled solution after leaching, two crystallization tanks operating at 30°C, and a vacuum filter. 96% sulfuric acid was dosed into the first crystallization tank to achieve a concentration of 150 g/l H2SO4 in the crystallization filtrate. The crystallization tanks were indirectly cooled with chilled water to reach a temperature of 30°C. The suspension from the second crystallization tank was filtered and washed. The combined filtrate and washed filtrate were stored in stripper tanks operating in the operating/standby mode. This solution, after leaching, was pumped into the stripping circuit by solvent extraction of copper.

The average copper concentration in the copper SX stripper feedstock and the saturated stripper leach solutions was 50.0 g/L Cu and 67.1 g/L Cu, respectively. Some 48.6 kg of copper sulfate was obtained with an average quality of 27.4% Cu. The main impurity was iron (138 ppm). The copper recovered is relatively consistent with the copper present in the original leach solution subjected to solvent copper extraction.

Example 9.

An iron and aluminum deposition test cycle was conducted in which approximately 1700 d of solvent copper raffinate recovered from the CuSX test facility was processed for 90 hours. The circuit contained two stages of deposition, partial neutralization (PN) and secondary neutralization (SN).

Iron and aluminum deposition in the PN circuit consistently exceeded 80% for most of the cycle. The filtrate contained an average of 23.8 g/L Co, 4.13 g/L Ni, 2.93 g/L Mn, 1.21 g/L Al, and 55 mg/L Fe. Concentrations of cobalt, manganese and nickel were lower than in the feed, mainly due to dilution. The cobalt and nickel contained in the PN residue were not recovered and were estimated to be approximately 1.2% contained cobalt and 4.0% contained nickel. These losses can be minimized by operating at a lower pH in the PN circuit.

Iron and aluminum deposition in the SN circuit consistently exceeded 98% for most of the cycle. The filtrate contained an average of 18.1 g/L Co, 2.70 g/L Ni, 2.84 g/L Mn, 17 mg/L Al, and <1 mg/L Fe. The SN residue was recovered by filtering, washing and drying product batches.

-

Claims (19)

The results show significant cobalt and nickel losses in the SN residue of approximately 23% and 35%, respectively. This causes a significantly lower concentration of cobalt and nickel in the filtrate after SN compared to the SN feed. Cobalt and nickel can be recovered from these residues by returning the residue upstream from the PN loop or to a separate leach loop. Example 10. A cobalt solvent extraction cycle was carried out in which approximately 500 L of the post-precipitation filtrate taken from the precipitation test facility was treated for 45 hours. The circuit consisted of four extraction stages, two scrubbing stages and two distillation stages. The initial solution after leaching contained 19.9 g/L Co and 2.93 g/L Mn. The raffinate contained 18.3 mg/L Co and 2.93 mg/L Mn, resulting in 99.9% Co and 99.97% Mn recovery over the cycle. The saturated distilled solution after leaching contained 84.0 g/L Co, 9.06 g/L Mn and only 22 mg/L Mg and 6 mg/L Ni. Example 11. A manganese solvent extraction cycle was carried out in which approximately 120 L of artificial saturated distillate leach solution containing 81 g/L Co and 13 g/L Mn were treated for 48 hours. The circuit consisted of four pre-saturation stages, six extraction stages, three scrubbing stages and two distillation stages. Once equilibrated, the extraction raffinate contained 84.7 g/L Co and only 5 mg/L Mn, and the saturated distillate leach solution contained 39.5 g/L Mn and only 24 mg/L Co. These results show excellent separation of cobalt from manganese. Example 12. A nickel solvent extraction cycle was conducted in which approximately 1000 L of SX cobalt filtrate recovered from the cobalt solvent extraction test facility was treated for 94 hours. The circuit consisted of four extraction stages, two scrubbing stages and three distillation stages. The initial solution after leaching contained 2.06 g/L Ni, 2.48 g/L Li, 50 mg/L Co and 65 mg/L Mn. The raffinate contained 104 mg/L Ni, 31 mg/L Co, 2.35 g/L Li and 62 mg/L Mg, indicating 95% Ni extraction for the cycle, but consistently achieved >99% after 24 hours of operation over the duration cycle. The saturated distilled solution after leaching contained 51.7 g/L Ni and only 1 mg/L Mg, 7 mg/L Co and <2 mg/L Li. The results show excellent separation of nickel from cobalt, lithium and magnesium in the leaching solution. Those skilled in the art will appreciate that the invention described herein is subject to changes and modifications other than those specifically described. The invention is intended to include all such variations and modifications. The invention also includes the steps, features, compositions and compounds mentioned or indicated in the specification, individually or collectively, and any possible combinations or any two or more steps or features. Each document, reference, patent application or patent cited in this text is expressly incorporated herein in its entirety by reference, meaning that it should be read and treated as part of this text. A document, reference, patent application or patent cited in this text is not repeated in this text solely for reasons of brevity. Any manufacturer's instructions, descriptions, product specifications and process sheets for any products mentioned herein or in any document incorporated herein by reference are hereby incorporated herein by reference and may be used in the practice of the present invention. The scope of the present invention is not limited to any of the specific embodiments described herein. These embodiments are for illustration purposes only. Functionally equivalent products, compositions and methods are clearly within the scope of the invention as described herein. The present invention described herein may include one or more ranges of values (eg, size, offset and field strength, etc.). A range of values will mean all values within the range, including values that define the range and values approximating the range that lead to the same or substantially the same result as values immediately approximating the value that defines the boundary range. CLAIM 1. A method of extracting metals from a feed stream containing one or more valuable metals and lithium, the method comprising subjecting the feed stream to sulfuric acid leaching to form a suspension containing a productive leach solution of soluble metal salts and a solid residue; separation of productive leach solution and solid residue; subjecting the productive leach solution to a precipitation step to extract iron and - 34 045218 aluminum, wherein the precipitation step includes the addition of ammonia to increase the pH of the productive leach solution to a level above 4.5; subjecting the productive leach solution to one or more separate solvent extraction steps, wherein each solvent extraction step extracts one or more valuable metals from the productive leach solution, and wherein each solvent extraction step includes dosing the productive leach solution with ammonia to maintain a target pH, and the remaining productive leach solution the leach solution contains lithium; extracting lithium from the productive leach solution; After the lithium extraction step, the leach solution is subjected to an ammonium sulfate extraction step. 2. The method of claim 1, wherein the feed stream contains at least lithium and one or more metals selected from copper, iron, aluminum, manganese, cobalt and nickel. 3. The method according to claim 1 or 2, characterized in that the method further includes the following step: subjecting the raw material stream to a pre-treatment process prior to the step of subjecting the raw material stream to sulfuric acid leaching. 4. The method according to claim 3, characterized in that the pre-treatment process includes one or more mechanical processing steps. 5. The method according to claim 3 or 4, characterized in that the mechanical processing steps include one or more of a crushing step and a grinding step. 6. Method according to any one of claims 3 to 5, characterized in that the pre-treatment process includes one or more size reduction steps. 7. Method according to any one of claims 3-6, characterized in that the pre-treatment process includes one or more enrichment steps. 8. The method of claim 7, wherein the one or more enrichment steps comprise one or more of an air classification step, a magnetic classification step, and a flotation step. 9. The method according to any of the previous paragraphs, characterized in that the step of subjecting the raw material stream to sulfuric acid leaching is carried out at atmospheric pressure. 10. The method according to any of the previous paragraphs, characterized in that the step of subjecting the raw material stream to sulfuric acid leaching is carried out at an elevated temperature. 11. The method according to any of the previous paragraphs, characterized in that the step of subjecting the feed stream to sulfuric acid leaching is carried out in one or more leaching reactors. 12. The method of any one of the preceding claims, wherein each solvent extraction step comprises contacting a productive leach solution with an extractant to extract one or more metals to produce a saturated extractant containing one or more extracted metals; separating the saturated extractant from the productive leaching solution; extraction of metal from a saturated extractant. 13. A method according to any of the preceding claims, characterized in that one of the separate solvent extraction steps is designed to extract manganese from the productive leach solution. 14. A method according to any of the preceding claims, characterized in that one of the separate solvent extraction steps is designed to extract cobalt from the productive leach solution. 15. A method according to any one of claims 1 to 12, characterized in that one of the separate solvent extraction steps is designed to extract both cobalt and manganese from the productive leach solution. 16. A method according to any of the preceding claims, characterized in that one of the separate solvent extraction steps is designed to extract nickel from the productive leach solution. 17. Method according to any of the previous paragraphs, characterized in that the individual solvent extraction steps are performed sequentially. 18. The method according to claim 17, characterized in that individual solvent extraction steps are performed to extract metals in the following sequence: copper, manganese, cobalt, nickel. 19. The method according to claim 17, characterized in that manganese and cobalt are extracted together, and separate solvent extraction steps are performed to extract metals in the following sequence: copper, manganese and cobalt, nickel. -