Method for resource utilization of laterite-nickel ore

Abstract

A method for resource utilization of laterite-nickel ore, belonging to the technical field of the treatment of laterite-nickel ore. The method comprises: performing high-pressure acid leaching and solid-liquid separation on laterite-nickel ore to obtain a leachate and a leaching residue; adjusting the pH value of the leachate to 2.0-5.0 and performing solid-liquid separation to obtain an adjusted solution and an adjustment residue; performing ion exchange on the adjusted solution by means of an ion exchange resin to obtain an adsorbent resin having adsorbed nickel ions from the adjusted solution and the adsorption residual liquid; and using an acid to desorb the adsorbent resin and performing ion exchange membrane treatment on a nickel-containing desorption acid solution obtained after desorption, so as to separate a nickel salt from the acid to obtain a nickel salt solution to be purified. The method is simple, has a low cost, and can stably and reliably recover a metal such as nickel.

Description

A method for resource utilization of laterite nickel ore Technical Field

[0001] The present disclosure relates to the technical field of laterite nickel ore processing, and in particular to a method for resource utilization of laterite nickel ore. Background Art

[0002] Nickel is a very important metal with excellent physical and chemical properties, making it widely used in many fields. With the development of human society, the demand for nickel resources will continue to increase. Nickel resources mainly come from nickel-containing ores, including nickel sulfide ore and laterite nickel ore. Nickel sulfide ore reserves account for approximately 30%, while laterite nickel ore reserves account for 70%. Currently, nickel products on the market are mainly derived from nickel sulfide ore, which is limited by the relatively insufficient reserves of nickel sulfide ore. Laterite nickel ore has the advantages of abundant reserves and easy mining. Therefore, fully developing and utilizing the nickel contained in laterite nickel ore resources is of great practical significance.

[0003] Currently, the mainstream wet process for extracting nickel from laterite nickel ore to make nickel sulfate products has problems such as low nickel recovery rate, long process flow, high alkali consumption, and high production cost.

[0004] In view of this, the present disclosure is proposed.

[0005] Summary of the Invention

[0006] The purpose of the present disclosure is to provide a method for resource utilization of laterite nickel ore to solve or improve the above technical problems.

[0007] The present disclosure can be implemented as follows:

[0008] The present disclosure provides a method for resource utilization of laterite nickel ore, comprising the following steps: subjecting the laterite nickel ore to high-pressure acid leaching, solid-liquid separation, and obtaining a leachate and a leach residue; adjusting the pH value of the leachate to 2.0-5.0, and solid-liquid separation, to obtain a pH-adjusted liquid and a pH-adjusted residue; subjecting the pH-adjusted liquid to ion exchange with an ion exchange resin to obtain an adsorption resin that adsorbs nickel ions in the pH-adjusted liquid and a remaining adsorption residual liquid; desorbing the adsorption resin with acid, and subjecting the obtained nickel-containing desorbed acid solution to ion exchange membrane treatment to separate nickel salts from the acid, thereby obtaining a nickel salt solution to be purified.

[0009] In an optional embodiment, the method further includes: subjecting the nickel salt solution to be purified to a first extraction to remove calcium, manganese, copper and zinc in the nickel salt solution to obtain a first raffinate; subjecting the first raffinate to a second extraction to remove magnesium in the first raffinate to obtain a purified nickel salt solution; and deoiling, evaporating and crystallizing the purified nickel salt solution to obtain nickel salt.

[0010] In an alternative embodiment, the pH of the leachate is adjusted by adding additional nickel laterite ore to the leachate.

[0011] In an optional embodiment, the high pressure acid leaching includes at least one of the following features:

[0012] Feature 1: The acid used includes at least one of sulfuric acid, hydrochloric acid and nitric acid;

[0013] Feature 2: The amount of acid used is 8 to 12 times the theoretical amount of nickel and cobalt in the laterite nickel ore reacting with sulfuric acid;

[0014] Feature 3: The pressure is 3MPa-6MPa; and / or the temperature is 220°C-280°C; and / or the time is 2h-6h.

[0015] In an alternative embodiment, the ion exchange with the ion exchange resin is performed in a continuous ion exchange device.

[0016] In an optional embodiment, the continuous ion exchange device includes an adsorption zone, a water washing zone, a desorption zone, and a backwash zone, and the ion exchange resin for ion exchange includes at least one of the following characteristics:

[0017] Feature 1: Ion exchange resins include hydrogen ion exchange resins and sodium ion exchange resins;

[0018] Feature 2: The resin loading in each resin column is 70vt%-85vt%;

[0019] Feature 3: The flow rate in the adsorption zone is 6BV / h-12BV / h;

[0020] Feature 4: The acid used in the desorption zone includes at least one of sulfuric acid, hydrochloric acid and nitric acid, the concentration of the acid used in the desorption zone is 3 mol / L-6 mol / L, and the flow rate in the desorption zone is 2 BV / h-5 BV / h;

[0021] Feature 5: The detergent used in the water washing zone and the backwash zone is water, and the flow rate of the detergent in the water washing zone and the backwash zone is independently 3BV / h-6BV / h.

[0022] In an optional embodiment, the ion exchange membrane process includes at least one of the following features:

[0023] Feature 1: The ion exchange membrane includes at least one of an anion exchange membrane, a cation exchange membrane and an amphoteric ion exchange membrane;

[0024] Feature 2: The temperature of ion exchange membrane treatment is 20℃-60℃;

[0025] Feature 3: The dialysis coefficient of ion exchange membrane treatment is 1.0×10 -7m / s-30×10 -7 m / s.

[0026] In an optional embodiment, the extractant used for the first extraction includes P204, and / or the extractant used for the second extraction includes C272; and / or the degreasing agent used for degreasing includes activated carbon.

[0027] In an optional embodiment, the method further includes returning the adjusted slag to the high-pressure acid leaching process.

[0028] In an optional embodiment, the further step includes returning the acid separated by the ion exchange membrane treatment to the high-pressure acid leaching process.

[0029] In an optional embodiment, the method further comprises: mixing the adsorption residual liquid with an oxidant to make Fe 2+ Oxidized to Fe 3+ Then mix with alkaline reagent and separate solid and liquid to remove Al in the adsorption residual liquid. 3+ 、Fe 3+ Sc 3+ and Cr 3+ , and obtain the impurity-removed liquid and impurity-removed residue.

[0030] In an optional embodiment, the oxidant includes at least one of hydrogen peroxide, manganese dioxide and sodium persulfide.

[0031] In an optional embodiment, the alkaline agent includes at least one of calcium carbonate, magnesium oxide, magnesium hydroxide, calcium hydroxide, sodium hydroxide, sodium carbonate and sodium sulfide.

[0032] In an alternative embodiment, the Al 3+ 、Fe 3+ Sc 3+ and Cr 3+ It is carried out at a pH of 4.0-7.0.

[0033] In an optional embodiment, the method further comprises: extracting and stripping the impurity-removed liquid to obtain a manganese-containing solution, a zinc-containing solution, and a cobalt-containing extracted liquid;

[0034] The manganese-containing solution and the zinc-containing solution are degreased, evaporated and crystallized to obtain manganese salt and zinc salt respectively.

[0035] In an optional embodiment, the extractant used for extracting the impurity-removed liquid includes P204, and the stripping agent used for stripping includes sulfuric acid.

[0036] In an optional embodiment, the method further comprises: extracting and stripping the cobalt-containing extracted liquid to obtain a cobalt-containing solution, and deoiling, evaporating and crystallizing the cobalt-containing solution to obtain a cobalt salt.

[0037] In an optional embodiment, the extractant used to extract the extracted liquid includes P507, and the stripping agent used for stripping includes sulfuric acid.

[0038] The beneficial effects of the present disclosure include:

[0039] The method for resource utilization of laterite nickel ore provided by the present disclosure is simple, low-cost, and can stably and reliably recover nickel and other metals. For example, the present disclosure uses ion exchange resin to highly selectively separate nickel from most impurities in the laterite nickel ore leachate, and after analysis, a preliminarily purified nickel salt solution can be obtained. In combination with an ion exchange membrane, the nickel salt in the nickel-containing analytical acid solution can be separated from the acid, and the nickel concentration in the nickel-containing analytical acid solution can be enriched to a higher level, which is conducive to improving the separation effect and purity of nickel. In addition, the ion exchange membrane process relies on the concentration difference within the medium as the driving force for separation, so no additional energy consumption is required, and the power cost is low.

[0040] By combining the ion exchange resin adsorption method with the ion exchange membrane method, high-purity and high-recovery nickel can be recovered from laterite nickel ore, which has the advantages of simple process, low production cost, stability and reliability. BRIEF DESCRIPTION OF THE DRAWINGS

[0041] In order to more clearly illustrate the technical solutions of the embodiments of the present disclosure, the following briefly introduces the drawings required for use in the embodiments. It should be understood that the following drawings only illustrate certain embodiments of the present disclosure and therefore should not be regarded as limiting the scope. For ordinary technicians in this field, other relevant drawings can be obtained based on these drawings without creative work.

[0042] FIG1 is a process flow chart of the method for resource utilization of laterite nickel ore provided in Example 1 of the present disclosure. DETAILED DESCRIPTION

[0043] To make the purpose, technical solutions, and advantages of the embodiments of the present disclosure more clear, the technical solutions in the embodiments of the present disclosure are described clearly and completely below. Where specific conditions are not specified in the embodiments, conventional conditions or conditions recommended by the manufacturer were followed. Reagents or instruments used where the manufacturer is not specified are conventional products that can be purchased commercially.

[0044] The method for resource utilization of laterite nickel ore provided by the present disclosure is described in detail below.

[0045] The present disclosure provides a method for resource utilization of laterite nickel ore, which may include the following steps: subjecting the laterite nickel ore to high-pressure acid leaching, solid-liquid separation, and obtaining a leachate and a leach residue; adjusting the pH value of the leachate to 2.0-5.0, and solid-liquid separation to obtain a conditioned liquid and a conditioned residue; subjecting the conditioned liquid to ion exchange with an ion exchange resin to obtain an adsorption resin that adsorbs nickel ions in the conditioned liquid and a remaining adsorption residual liquid; desorbing the adsorption resin with acid, and subjecting the obtained nickel-containing desorbed acid solution to ion exchange membrane treatment to separate nickel salts from the acid, thereby obtaining a nickel salt solution to be purified.

[0046] In some embodiments, the laterite nickel ore may be ground prior to high pressure acid leaching to obtain a laterite nickel ore powder, thereby improving the effect of high pressure acid leaching. The laterite nickel ore powder is then mixed with water to prepare a slurry.

[0047] For reference, high-pressure acid leaching can be carried out in a high-pressure reactor, and the corresponding pressure can be 3 MPa-6 MPa, such as 3 MPa, 4 MPa, 5 MPa, or 6 MPa. The corresponding temperature can be 220° C.-280° C., such as 220° C., 230° C., 240° C., 250° C., 260° C., 270° C., or 280° C., and the corresponding time can be 2 hours-6 hours, such as 2 hours, 3 hours, 4 hours, 5 hours, or 6 hours.

[0048] The acid used in the high pressure acid leaching process may include, for example, at least one of sulfuric acid, hydrochloric acid, and nitric acid. In some embodiments, the acid used in the high pressure acid leaching process is sulfuric acid with a concentration of about 98%.

[0049] The amount of acid used can be 8 to 12 times the theoretical amount of nickel and cobalt in the laterite nickel ore reacting with sulfuric acid, such as 8 times, 9 times, 10 times, 11 times or 12 times.

[0050] The high-pressure acid leaching treatment can leach out elements such as nickel, iron, manganese, zinc and cobalt from the laterite nickel ore, and the leachate and leach residue can be separated by solid-liquid separation (such as centrifugation or filtration).

[0051] In the present disclosure, the pH value of the leachate is adjusted by adding additional laterite nickel ore to the leachate. In this way, the addition of an alkaline reagent for neutralization and adjustment can be avoided, which can greatly reduce the processing cost.

[0052] For reference, the pH value of the leachate can be adjusted to 2.0, 2.5, 3.0, 3.5, 4.0, 4.5 or 5.0, or any other value or interval within the range of 2.0-5.0.

[0053] The adjusted slag obtained after adjustment can be further returned to the high-pressure leaching process.

[0054] In the present disclosure, ion exchange of the adjusting liquid through the ion exchange resin is carried out in a continuous ion exchange device, so that the nickel ions in the adjusting liquid are ion-exchanged with the resin. The nickel ions in the adjusting liquid are adsorbed by the ion exchange resin, which is conducive to separating nickel from most impurities.

[0055] The above-mentioned continuous ion exchange equipment can be purchased and used directly. It consists of an adsorption zone, a water washing zone, a desorption zone, and a backwash zone. During operation, the adjustment solution first enters the adsorption zone for ion exchange. After the nickel ions have been adsorbed, the resin column moves to the water washing zone for water washing. After the water washing, the resin column moves to the desorption zone for desorption. After desorption, the resin column moves to the backwash zone for backwashing. After backwashing, the resin column enters the next round of adsorption, water washing, desorption, and backwashing.

[0056] For reference, the ion exchange resin may include a hydrogen ion exchange resin and a sodium ion exchange resin to effectively adsorb nickel ions while substantially not adsorbing other metals or impurities.

[0057] The continuous ion exchange device may be provided with 24 resin columns by way of example. In other embodiments, the number of resin columns may be adjusted as needed.

[0058] The resin loading in each resin column may be 70 vt% to 85 vt%, such as 70 vt%, 75 vt%, 80 vt% or 85 vt%, or any other value within the range of 70 vt% to 85 vt%. "vt%" may be understood as the volume percentage of the resin column.

[0059] If the resin filling amount is too low, it will be detrimental to the utilization of the resin column equipment space; if the resin filling amount is too high, the volume of the ion exchange lipid material will expand during adsorption and decomposition, so a certain buffer space is required. Too high a filling amount will cause damage to the resin material.

[0060] The flow rate in the adsorption zone can be 6 BV / h to 12 BV / h, such as 6 BV / h, 7 BV / h, 8 BV / h, 9 BV / h, 10 BV / h, 11 BV / h, or 12 BV / h, or any other value within the range of 6 BV / h to 12 BV / h, wherein "BV" refers to the volume of the resin material packed in the resin column.

[0061] If the flow rate in the adsorption zone is too low, the production efficiency will be reduced; if the flow rate in the adsorption zone is too high, it will easily lead to insufficient adsorption.

[0062] The acid used in the desorption zone may be an inorganic acid, such as at least one of sulfuric acid, hydrochloric acid, and nitric acid. The concentration of the acid used in the desorption zone may be 3 mol / L to 6 mol / L, such as 3 mol / L, 3.5 mol / L, 4 mol / L, 4.5 mol / L, 5 mol / L, 5.5 mol / L, or 6 mol / L, or any other value within the range of 3 mol / L to 6 mol / L.

[0063] The flow rate in the desorption zone can be 2BV / h-5BV / h, such as 2BV / h, 2.5BV / h, 3BV / h, 3.5BV / h, 4BV / h, 4.5BV / h or 5BV / h, or any other value within the range of 2BV / h-5BV / h.

[0064] The above analytical conditions can achieve a high analytical efficiency for nickel ions adsorbed in the resin.

[0065] The detergent used in the water washing zone and the backwash zone is water. The flow rate of the detergent in the water washing zone and the backwash zone can independently be 3BV / h-6BV / h, such as 3BV / h, 3.5BV / h, 4BV / h, 4.5BV / h, 5BV / h, 5.5BV / h or 6BV / h, etc., or it can be any other value within the range of 3BV / h-6BV / h.

[0066] The nickel-containing analytical acid solution obtained after the analytical process is subjected to ion exchange membrane treatment. In the present disclosure, the ion exchange membrane used for the ion exchange membrane treatment may include at least one of an anion exchange membrane, a cation exchange membrane, and an amphoteric ion exchange membrane.

[0067] The temperature of the ion exchange membrane treatment can be 20°C-60°C, such as 20°C, 25°C, 30°C, 35°C, 40°C, 45°C, 50°C, 55°C or 60°C, or any other value within the range of 20°C-60°C.

[0068] If the treatment temperature of the ion exchange membrane is higher than 60°C, the service life of the ion exchange membrane will be affected.

[0069] The dialysis coefficient of ion exchange membrane treatment can be 1.0×10 -7 m / s-30×10 -7 m / s, such as 1×10 -7 m / s, 5×10 -7 m / s, 10×10 -7 m / s, 15×10 -7 m / s, 20×10 -7 m / s, 25×10 -7 m / s or 30×10 -7 m / s, etc., or 1.0×10 -7m / s-30×10 -7 Any other value within the range of m / s.

[0070] The ion exchange membrane with this dialysis coefficient is more conducive to the separation of nickel ions and acid than ion exchange membranes with other dialysis coefficients.

[0071] Accordingly, the ion exchange membrane treatment time can be 30 min-60 min.

[0072] The nickel salt and the acid in the nickel-containing analytical acid solution can be separated by the ion exchange membrane treatment, and the separated acid can be returned to the high-pressure acid leaching process as needed.

[0073] As mentioned above, using an ion exchange membrane to treat the nickel-containing analytical acid solution obtained from the resin system can separate the acid from the nickel-containing analytical acid solution, yielding nickel salts and acid. This allows the nickel concentration in the nickel-containing analytical acid solution to be enriched to a very high level. Furthermore, the ion exchange membrane process relies on concentration differences within the medium as the driving force for separation, thus requiring no additional energy consumption and resulting in low power costs.

[0074] The present disclosure adopts a combination of ion exchange resin adsorption method and ion exchange membrane method, which can process a low-nickel solution containing many impurities and high concentration to obtain a high-nickel solution containing few impurities and low concentration, which is beneficial for subsequent further processing.

[0075] The nickel salt solution to be purified obtained by the ion exchange membrane treatment can be subjected to a first extraction to remove calcium, manganese, copper and zinc in the nickel salt solution to obtain a first raffinate.

[0076] For reference, the extractant used in the first extraction may, for example but not by way of limitation, include P204. The concentration of the extractant used in the first extraction may be 10%-40%, and the volume ratio of the extractant used in the first extraction to the nickel salt solution to be purified may be (1:3)-(1:6). The temperature of the first extraction may be 30°C-60°C, and the pH value may be 2.0-5.0.

[0077] The first raffinate can be further subjected to a second extraction to remove magnesium in the first raffinate to obtain a purified nickel salt solution.

[0078] The extractant used in the second extraction may illustratively, but not exclusively, include C272. The concentration of the extractant used in the second extraction may be 10% to 30%, and the volume ratio of the extractant used in the second extraction to the first raffinate may be (1:3) to (1:6). The temperature of the second extraction may be 30° C. to 60° C., and the pH may be 3.0 to 5.0.

[0079] Furthermore, the purified nickel salt solution is subjected to deoiling, evaporation and crystallization to obtain high-purity nickel salt.

[0080] The degreasing agent used for degreasing may include, by way of example but not limitation, activated carbon. The activated carbon may be added in a liquid-to-solid ratio of (1 mL:0.01 g) to (1 mL:0.05 g). The evaporation and crystallization temperature of the nickel salt solution may be 40° C. to 80° C.

[0081] Continuing from the above, the present invention utilizes ion exchange resin to highly selectively separate nickel from most impurities in laterite nickel ore leachate, and obtains a preliminarily purified nickel-containing analytical acid solution after analysis. After acid separation and extraction to remove impurities through an ion exchange membrane, a purified nickel salt solution is obtained. After degreasing and evaporation and crystallization, nickel salts (such as nickel sulfate) that meet battery-grade standards can be prepared.

[0082] In the present disclosure, the adsorption residual liquid obtained by ion exchange of the above ion exchange resin can be mixed with an oxidant to make the Fe 2+ Oxidized to Fe 3+ Then mix with alkaline reagent and separate solid and liquid to remove Al in the adsorption residual liquid. 3+ 、Fe 3+ Sc 3+ 、Cu 2+ and Cr 3+ , and obtain the impurity-removed liquid and impurity-removed residue.

[0083] For reference, the oxidant may, by way of example but not limitation, include at least one of hydrogen peroxide, manganese dioxide, and sodium persulfide. The concentration of the oxidant may be 20 wt% to 50 wt%, and the amount of the oxidant used relative to the adsorption residual solution may be 1.0 to 1.5 times the theoretical amount of the reaction. The reaction time between the adsorption residual solution and the oxidant may be 30 to 60 minutes.

[0084] The alkaline agent may illustratively but not limitatively include at least one of calcium carbonate, magnesium oxide, magnesium hydroxide, calcium hydroxide, sodium hydroxide, sodium carbonate, and sodium sulfide.

[0085] The above removal of Al 3+ 、Fe 3+ Sc 3+ and Cr 3+ The reaction can be carried out at a pH of 4.0-7.0 (e.g., 4.0, 4.5, 5.0, 5.5, 6.0, 6.5, or 7.0, etc.) The pH range can be controlled by adjusting the amount of the alkaline reagent.

[0086] The impurity-removed liquid can also be subjected to extraction and back extraction to obtain a manganese-containing solution, a zinc-containing solution and a cobalt-containing extracted liquid.

[0087] For reference, the extractant used to extract the impurity-removed liquid can, for example but not by way of limitation, include P204. The concentration of the extractant can be 10% to 40%, and the volume ratio of the extractant to the feed liquid can be (1:3) to (1:6). The temperature during the extraction process can be 30°C to 60°C, and the pH can be 2.0 to 5.0.

[0088] The stripping agent used for stripping may, for example but not limited to, sulfuric acid. The concentration of the stripping agent may be 0.5 mol / L to 4 mol / L, and the amount used for stripping may be 1.0 to 1.3 times the theoretical amount of metal ion exchange. The stripping process temperature may be 30° C. to 60° C., and the pH may be 1.0 to 5.0.

[0089] The manganese-containing solution and the zinc-containing solution can be degreased and evaporated to obtain manganese salts and zinc salts, respectively. The degreaser used in the degreasing process can include, by way of example but not limitation, activated carbon. The activated carbon can be added in a liquid-to-solid ratio of (1 mL:0.005 g) to (1 mL:0.03 g). The manganese-containing solution and / or the zinc-containing solution are transferred to a rotary evaporator for evaporation and crystallization at a temperature of 40°C to 80°C.

[0090] Similarly, the cobalt-containing extracted liquid can be subjected to extraction and back extraction to obtain a cobalt-containing solution, and the cobalt-containing solution can be deoiled, evaporated and crystallized to obtain a cobalt salt.

[0091] For reference, the extractant used in the cobalt-containing post-extraction solution may, for example but not exclusively, include P507. The concentration of the extractant may be 10% to 40%, and the volume ratio of the extractant to the feed solution may be (1:3) to (1:6). The extraction process temperature may be 30°C to 60°C, and the pH may be 2.0 to 5.0.

[0092] The stripping agent used for stripping may, for example but not limited to, include sulfuric acid. The concentration of the stripping agent may be 1.0 mol / L to 3.0 mol / L, and the amount used for stripping may be 1.0 to 1.3 times the theoretical amount of metal ion exchange. The stripping process temperature may be 30° C. to 60° C.

[0093] The degreasing agent used for degreasing may include, for example but not limited to, activated carbon. The amount of activated carbon used may be (1 mL: 0.01 g) to (1 mL: 0.05 g). The evaporation crystallization temperature of the cobalt-containing solution may be 40° C. to 80° C.

[0094] As mentioned above, the present invention utilizes the high selectivity of ion exchange resin for nickel to selectively separate nickel and cobalt from laterite nickel ore leachate, and then further treats them to obtain nickel salt, cobalt salt, manganese salt and zinc salt products with higher purity.

[0095] The present disclosure provides a method for effectively recovering nickel, cobalt, manganese, and zinc from laterite nickel ore, with advantages such as simplicity, low production cost, and stability and reliability. Furthermore, the regeneration capability of ion exchange resins and the recyclability of ion exchange membranes significantly reduce recovery costs.

[0096] The features and performance of the present disclosure are further described in detail below with reference to the embodiments.

[0097] Example 1

[0098] This embodiment provides a method for resource utilization of laterite nickel ore, referring to FIG1 , comprising the following steps:

[0099] S1: Laterite nickel ore is ground to obtain a powder with an average particle size of 50 μm. The powder is mixed with water at a solid-to-liquid ratio of 1 g:2 mL to obtain a slurry. 98% sulfuric acid (10 times the theoretical amount required for nickel and cobalt reaction) is added to the slurry. The slurry is then transferred to an autoclave for high-pressure leaching (at a pressure of 4.5 MPa, a reaction temperature of 260°C, and a reaction time of 4 hours). After the reaction is complete, the slurry is filtered to obtain a leachate and a leached residue.

[0100] The leachate is then added with additional laterite nickel ore powder to adjust the pH of the leachate to 5.0, and filtered to obtain an adjusted solution and adjusted residue. The adjusted residue is returned to the autoclave for leaching reaction, and the adjusted solution is subjected to the next step of treatment.

[0101] Calculated by mass percentage, the laterite nickel ore contains Ni 1.79wt%, Co 0.036wt%, Mn 0.28wt%, Fe 26.4wt%, Mg 11.7wt%, Al 1.65wt%, Zn 0.86wt%, Cr 0.29wt%, and Sc 0.0027wt%.

[0102] S2: 100 L of the adjusted liquid obtained in S1 was pumped into the adsorption zone of a continuous ion exchange device (TPD-24II type equipment produced by Sanda Membrane Environmental Technology Co., Ltd.) using four peristaltic pumps at a rate of 8 BV / h, each pumping at a rate of 8 BV / h. The continuous ion exchange device has a total of 24 columns, and the resin column dimensions are 5 cm in diameter × 12 cm in height. The resin columns in the adsorption zone are 4 in parallel and 4 in series, the water washing zone is 2 columns in series, the desorption zone is 4 columns in series, and the backwash zone is 2 columns in series. Each column is filled with 80% of the volume of LX-92 hydrogen-type ion exchange resin; pure water is pumped into the water washing zone using a peristaltic pump at a rate of 4 BV / h. The desorption acid is sulfuric acid with a concentration of 4 mol / L, which is pumped into the desorption zone using a peristaltic pump at a rate of 2 BV / h. Pure water is pumped into the backwash zone using a peristaltic pump at a rate of 5 BV / h. The resin column switching time is 40 min. The adsorption residual liquid after the nickel ions are adsorbed and the product obtained by desorption are respectively processed in the subsequent corresponding steps. The continuous ion exchange system includes an adsorption zone, a water washing zone, a desorption zone and a backwash zone. During operation, the adjustment liquid first enters the adsorption zone for ion exchange. After adsorbing nickel ions, the resin column switches to the water washing zone for water washing. After water washing, the resin column switches to the desorption zone for desorption. After desorption, the resin column switches to the backwash zone for backwashing. After backwashing, the resin column enters the next round of adsorption, water washing, desorption and backwashing processes.

[0103] S3: The adsorption residual liquid obtained in S2 was first added with 27.5wt% hydrogen peroxide with a concentration of 1.2 times the theoretical amount of ferrous iron to react with the reaction for 40 minutes to make Fe 2+ Oxidized to Fe 3+ Then, magnesium oxide was added to adjust the pH of the solution to 6.5, and the Al in the solution was 3+ 、Fe 3+ Sc 3+ 、Cu 2+ and Cr 3+ The ions are precipitated to remove impurities, and after the pH is stabilized, the solid and liquid are separated to obtain the impurity-removed liquid and impurity-removed residue.

[0104] The impurity-removed liquid was first subjected to P204 extraction (referred to as "extraction three", the P204 concentration was 20%, the extractant to liquid volume ratio was 1:4, the extraction temperature was 45°C, and the pH value was 3.0), and sulfuric acid back extraction (the concentration of sulfuric acid was 2 mol / L, the amount used was 1.1 times the theoretical amount for metal ion exchange, the back extraction temperature was 40°C, and the pH value was 3.0) to obtain manganese sulfate solution and zinc sulfate solution, followed by activated carbon deoiling (activated carbon was added at a liquid-to-solid ratio of 1 mL:0.01 g) and evaporation crystallization (60°C) to obtain manganese sulfate and zinc sulfate. The residual liquid after extracting manganese and zinc is then extracted with P507 (referred to as "extraction four", the concentration of P507 is 30%, the volume ratio of extractant to liquid is 1:3, the extraction temperature is 40°C, and the pH value is 4.5), and back-extraction with sulfuric acid (the concentration of sulfuric acid is 2 mol / L, the amount used is 1.1 times the theoretical amount for exchange with cobalt metal ions, and the back-extraction temperature is 45°C) to obtain a cobalt sulfate solution, which is then deoiled with activated carbon (activated carbon is added at a liquid-to-solid ratio of 1 mL:0.02 g) and evaporated and crystallized (temperature is 75°C) to obtain cobalt sulfate.

[0105] S4: The nickel-containing analytical acid solution obtained in S2 is treated with an anion exchange membrane with a membrane permeability coefficient of 15×10 -7 m / s, the solution temperature is controlled at 40°C during treatment, and the treatment time is 40 min, so that the nickel salt and sulfuric acid in the nickel-containing analytical acid solution are separated to obtain the nickel salt solution to be purified and the dilute sulfuric acid solution.

[0106] The dilute sulfuric acid solution is returned to the high-pressure leaching process of S1. The nickel salt solution to be purified is first subjected to P204 extraction (referred to as "extraction 1", the concentration of P204 is 20%, the volume ratio of the extractant to the liquid is 1:4, the extraction temperature is 45°C, and the pH value is 3.0) to remove a small amount of calcium, manganese, copper and zinc impurities in the solution, and sulfuric acid stripping (the concentration of sulfuric acid is 2 mol / L, the amount used is 1.1 times the theoretical amount of metal ion exchange, and the stripping temperature is 40°C) to obtain the first raffinate; the first raffinate is then subjected to C272 Extraction (abbreviated as "extraction 2", the concentration of C272 is 15%, the volume ratio of extractant to liquid is 1:3, the extraction temperature is 40°C, and the pH value is 5.0) is used to remove a small amount of magnesium in the solution, and sulfuric acid back extraction (the concentration of sulfuric acid is 2 mol / L, the amount used is 1.1 times the theoretical amount for metal ion exchange, and the back extraction temperature is 40°C) is used to obtain a purified nickel sulfate solution, which is then deoiled with activated carbon (activated carbon is added at a liquid-to-solid ratio of 1 mL:0.01 g) and evaporated and crystallized (temperature is 70°C) to obtain nickel sulfate.

[0107] Example 2

[0108] This embodiment provides a method for resource utilization of laterite nickel ore, comprising the following steps:

[0109] S1: Laterite nickel ore is ground to obtain a powder with an average particle size of 60 μm. The powder is mixed with water at a solid-liquid ratio of 1 g:1.5 mL to obtain a slurry. 98% sulfuric acid (12 times the theoretical amount required for nickel and cobalt reaction) is added, and the slurry is transferred to an autoclave for high-pressure acid leaching (pressure 4.8 MPa, reaction temperature 270°C, reaction time 6 hours). After the reaction is complete, the slurry is filtered to obtain a leachate and leached residue.

[0110] The leachate is then added with additional laterite nickel ore powder to adjust the pH of the leachate to 3.0, and filtered to obtain an adjusted solution and adjusted residue. The adjusted residue is returned to the autoclave for leaching reaction, and the adjusted solution is subjected to the next step of treatment.

[0111] Calculated by mass percentage, the laterite nickel ore contains Ni 1.65wt%, Co 0.056wt%, Mn 0.68wt%, Fe 26.4wt%, Mg 11.7wt%, Al 1.65wt%, Zn 0.183wt%, Cr 0.29wt%, and Sc 0.0023wt%.

[0112] S2: 150 L of the adjusted liquid obtained in S1 was pumped into the adsorption zone of a continuous ion exchange device (TPD-24II type equipment produced by Sanda Membrane Environmental Technology Co., Ltd.) using three peristaltic pumps at a rate of 11 BV / h, each. The continuous ion exchange device has a total of 24 columns, and the resin column dimensions are 5 cm in diameter × 12 cm in height. The adsorption zone resin columns are 3 in parallel and 5 in series, the water washing zone is 2 columns in series, the desorption zone is 5 columns in series, and the backwash zone is 2 columns in series. Each column is filled with 80% of the volume of LX-92 hydrogen-type ion exchange resin; pure water is pumped into the water washing zone using a peristaltic pump at a rate of 5.5 BV / h. The desorption acid is sulfuric acid with a concentration of 5 mol / L, which is pumped into the desorption zone using a peristaltic pump at a rate of 2 BV / h. Pure water is pumped into the backwash zone using a peristaltic pump at a rate of 6 BV / h. The resin column switching time is 30 min. The adsorption residual liquid after the nickel ions are adsorbed and the product obtained by desorption are respectively processed in the subsequent corresponding steps. The continuous ion exchange system includes an adsorption zone, a water washing zone, a desorption zone and a backwash zone. During operation, the adjustment liquid first enters the adsorption zone for ion exchange. After adsorbing nickel ions, the resin column switches to the water washing zone for water washing. After water washing, the resin column switches to the desorption zone for desorption. After desorption, the resin column switches to the backwash zone for backwashing. After backwashing, the resin column enters the next round of adsorption, water washing, desorption and backwashing processes.

[0113] S3: The adsorption residual liquid obtained in S2 was first added with manganese dioxide with a theoretical reaction volume of 1.1 times that of ferrous iron to react for 30 minutes to make Fe 2+ Oxidized to Fe 3+ Then, magnesium hydroxide was added to adjust the pH of the solution to 6.0. 3+、Fe 3+ Sc 3+ 、Cu 2+ and Cr 3+ The ions are precipitated to remove impurities, and after the pH is stabilized, the solid and liquid are separated to obtain the impurity-removed liquid and impurity-removed residue.

[0114] The impurity-removed liquid was first subjected to P204 extraction (P204 concentration was 30%, the volume ratio of extractant to liquid was 1:5, the extraction temperature was 45°C, and the pH was 3.0), and sulfuric acid stripping (sulfuric acid concentration was 2 mol / L, the amount used was 1.1 times the theoretical amount for metal ion exchange, the stripping temperature was 40°C, and the pH was 5.0) to obtain manganese sulfate solution and zinc sulfate solution. Subsequently, activated carbon deoiling (activated carbon was added at a liquid-to-solid ratio of 1 mL:0.02 g) and evaporation crystallization (temperature was 75°C) were performed to obtain manganese sulfate and zinc sulfate. The residual liquid after extracting manganese and zinc is then extracted with P507 (the concentration of P507 is 30%, the volume ratio of extractant to liquid is 1:3, the extraction temperature is 45°C, and the pH value is 4.5), and stripped with sulfuric acid (the concentration of sulfuric acid is 2.0 mol / L, the amount used is 1.1 times the theoretical amount for cobalt metal ion exchange, and the stripping temperature is 45°C) to obtain a cobalt sulfate solution, which is then deoiled with activated carbon (activated carbon is added at a liquid-to-solid ratio of 1 mL:0.01 g) and evaporated and crystallized (at a temperature of 80°C) to obtain cobalt sulfate.

[0115] S4: The nickel-containing analytical acid solution obtained in S2 is treated with an anion exchange membrane with a membrane permeability coefficient of 20×10 -7 m / s, the solution temperature is controlled at 40°C during treatment, and the treatment time is 50 min, so that the nickel salt and sulfuric acid in the nickel-containing analytical acid solution are separated to obtain the nickel salt solution to be purified and the dilute sulfuric acid solution.

[0116] The dilute sulfuric acid solution is returned to the high-pressure leaching process of S1. The nickel salt solution to be purified is first extracted with P204 (the concentration of P204 is 20%, the volume ratio of the extractant to the liquid is 1:4, the extraction temperature is 45°C, and the pH value is 3.5) to remove a small amount of calcium, manganese, copper and zinc impurities in the solution, and then stripped with sulfuric acid (the concentration of sulfuric acid is 2 mol / L, the amount is 1.1 times the theoretical amount of metal ion exchange, and the stripping temperature is 40°C) to obtain a first raffinate. The first raffinate is then subjected to C2 The solution was subjected to C272 extraction (the concentration of C272 was 15%, the volume ratio of the extractant to the feed liquid was 1:3, the extraction temperature was 40°C, and the pH value was 4.5) to remove a small amount of magnesium in the solution, and sulfuric acid back-extraction was performed (the concentration of sulfuric acid was 2 mol / L, the amount used was 1.1 times the theoretical amount for metal ion exchange, and the back-extraction temperature was 40°C) to obtain a purified nickel sulfate solution, which was then subjected to activated carbon deoiling (activated carbon was added at a liquid-to-solid ratio of 1 mL:0.01 g) and evaporation crystallization (temperature was 70°C) to obtain nickel sulfate.

[0117] Example 3

[0118] This embodiment provides a method for resource utilization of laterite nickel ore, comprising the following steps:

[0119] S1: Laterite nickel ore is ground to obtain a powder with an average particle size of 50 μm. The powder is mixed with water at a solid-to-liquid ratio of 1 g:1.5 mL to obtain a slurry. 98% sulfuric acid (12 times the theoretical amount required for nickel and cobalt reaction) is added, and the slurry is transferred to an autoclave for high-pressure acid leaching (pressure 5.0 MPa, reaction temperature 280°C, reaction time 2 hours). After the reaction is complete, the slurry is filtered to obtain a leachate and leaching residue.

[0120] The leachate is then added with additional laterite nickel ore powder to adjust the pH of the leachate to 5.0, and filtered to obtain an adjusted solution and adjusted residue. The adjusted residue is returned to the autoclave for leaching reaction, and the adjusted solution is subjected to the next step of treatment.

[0121] Calculated by mass percentage, the laterite nickel ore contains Ni 1.79wt%, Co 0.036wt%, Mn 0.28wt%, Fe 26.4wt%, Mg 11.7wt%, Al 1.65wt%, Zn 0.187wt%, Cr 0.29wt%, and Sc 0.0027wt%.

[0122] S2: 80 L of the adjusted solution obtained in S1 was pumped into the adsorption zone of a continuous ion exchange device (TPD-24II type equipment produced by Sanda Membrane Environmental Technology Co., Ltd.) using three peristaltic pumps, each at a rate of 8.5 BV / h. The continuous ion exchange device had a total of 24 columns, with resin column dimensions of 5 cm diameter × height × 12 cm. The adsorption zone resin columns were 3 in parallel and 4 in series, the water washing zone resin columns were 4 in series, the desorption zone resin columns were 4 in series, and the backwash zone resin columns were 4 in series. Each column was filled with 80% of the volume of LX-92 hydrogen-type ion exchange resin. Pure water was pumped into the water washing zone using a peristaltic pump at a rate of 4 BV / h. The desorption acid was sulfuric acid with a concentration of 4.5 mol / L, which was pumped into the desorption zone using a peristaltic pump at a rate of 2.5 BV / h. Pure water was pumped into the backwash zone using a peristaltic pump at a rate of 5.5 BV / h. The resin column switching time was 35 minutes. The adsorption residual liquid after the nickel ions were adsorbed and the product obtained by desorption were respectively processed in the subsequent corresponding steps. The continuous ion exchange system includes an adsorption zone, a water washing zone, a desorption zone and a backwash zone. During operation, the adjustment liquid first enters the adsorption zone for ion exchange. After adsorbing nickel ions, the resin column switches to the water washing zone for water washing. After water washing, the resin column switches to the desorption zone for desorption. After desorption, the resin column switches to the backwash zone for backwashing. After backwashing, the resin column enters the next round of adsorption, water washing, desorption and backwashing processes.

[0123] S3: The adsorption residual liquid obtained in S2 was first added with 27.5wt% hydrogen peroxide with a concentration of 1.2 times the theoretical amount of ferrous iron to react for 50 minutes to make Fe 2+ Oxidized to Fe 3+ Then, sodium carbonate was added to adjust the pH of the solution to 6.0, so that the Al 3+ 、Fe 3+ Sc 3+ 、Cu 2+ and Cr 3+ The ions are precipitated to remove impurities, and after the pH is stabilized, the solid and liquid are separated to obtain the impurity-removed liquid and impurity-removed residue.

[0124] The impurity-removed liquid was first subjected to P204 extraction (P204 concentration was 30%, the volume ratio of extractant to liquid was 1:5, the extraction temperature was 40°C, and the pH was 3.0), and sulfuric acid stripping (sulfuric acid concentration was 2.0 mol / L, the amount used was 1.1 times the theoretical amount for metal ion exchange, the stripping temperature was 40°C, and the pH was 3.0) to obtain manganese sulfate solution and zinc sulfate solution. Subsequently, activated carbon deoiling (activated carbon was added at a liquid-to-solid ratio of 1 mL:0.01 g) and evaporation crystallization (temperature was 75°C) were performed to obtain manganese sulfate and zinc sulfate. The residual liquid after extracting manganese and zinc is then extracted with P507 (the concentration of P507 is 20%, the volume ratio of extractant to liquid is 1:4, the extraction temperature is 45°C, and the pH value is 4.5), and stripped with sulfuric acid (the concentration of sulfuric acid is 2.0 mol / L, the amount used is 1.1 times the theoretical amount for cobalt metal ion exchange, and the stripping temperature is 45°C) to obtain a cobalt sulfate solution. The solution is then deoiled with activated carbon (activated carbon is added at a liquid-to-solid ratio of 1 mL:0.01 g) and evaporated and crystallized (at a temperature of 80°C) to obtain cobalt sulfate.

[0125] S4: The nickel-containing analytical acid solution obtained in S2 is treated with an anion exchange membrane with a membrane permeability coefficient of 10×10 -7 m / s, the solution temperature is controlled at 35°C during treatment, and the treatment time is 60 min, so that the nickel salt and sulfuric acid in the nickel-containing analytical acid solution are separated to obtain the nickel salt solution to be purified and the dilute sulfuric acid solution.

[0126] The dilute sulfuric acid solution is returned to the high-pressure leaching process of S1. The nickel salt solution to be purified is first extracted with P204 (the concentration of P204 is 20%, the volume ratio of the extractant to the liquid is 1:4, the extraction temperature is 45°C, and the pH value is 3.5) to remove a small amount of calcium, manganese, copper and zinc impurities in the solution, and then stripped with sulfuric acid (the concentration of sulfuric acid is 2 mol / L, the amount is 1.1 times the theoretical amount of metal ion exchange, and the stripping temperature is 40°C) to obtain a first raffinate. The first raffinate is then subjected to C2 The solution was subjected to C272 extraction (the concentration of C272 was 20%, the volume ratio of the extractant to the feed liquid was 1:4, the extraction temperature was 50°C, and the pH value was 4.0) to remove a small amount of magnesium in the solution, and sulfuric acid back-extraction was performed (the concentration of sulfuric acid was 2 mol / L, the amount used was 1.1 times the theoretical amount for metal ion exchange, and the back-extraction temperature was 40°C) to obtain a purified nickel sulfate solution, which was then subjected to activated carbon deoiling (activated carbon was added at a liquid-to-solid ratio of 1 mL:0.05 g) and evaporation crystallization (temperature was 65°C) to obtain nickel sulfate.

[0127] Comparative Example 1

[0128] The difference between this comparative example and Example 1 is:

[0129] S4: Nickel carbonate is added to the nickel-containing analytical acid solution obtained in S2 to neutralize the residual acid in the nickel-containing analytical acid solution, thereby obtaining a nickel salt solution to be purified and a slag to be adjusted. The nickel salt solution to be purified is first subjected to P204 extraction to remove a small amount of calcium, manganese, copper, and zinc impurities in the solution, as in Example 1, to obtain a first raffinate. The first raffinate is then subjected to C272 extraction to remove a small amount of magnesium in the solution, thereby obtaining a purified nickel sulfate solution. The solution is then subjected to activated carbon degreasing and evaporation and crystallization to obtain nickel sulfate.

[0130] Comparative Example 2

[0131] This comparative example is the mainstream wet process for extracting nickel from laterite nickel ore to produce nickel sulfate products on the current market, specifically:

[0132] The adjusted solution obtained in S1 of Example 1 was first added with 27.5wt% hydrogen peroxide with a concentration of 1.2 times the theoretical amount of ferrous iron to react and react for 40 minutes, then alkali was added to adjust the concentration to 5.0 to remove iron and aluminum. After solid-liquid separation, the obtained iron and aluminum-removed solution was added with sodium hydroxide to adjust the solution pH to 8.0 to precipitate nickel and cobalt. After solid-liquid separation, MHP precipitate residue and nickel and cobalt precipitation solution were obtained. The MHP precipitate residue was added with pure water at a solid-liquid ratio of 1:1.5g / mL to make a pulp, and then 98% sulfuric acid was added to dissolve the reaction slurry until the pH value stabilized to 1.5-2.0, followed by solid-liquid separation to obtain a leachate and an insoluble residue. The leachate was first added with iron powder with a theoretical reaction amount of 1.1 times that of copper to react for 60 minutes to remove copper, and then 27.5% hydrogen peroxide with a theoretical reaction amount of 1.2 times that of ferrous iron was added to react for 50 minutes to oxidize ferrous iron. The reacted solution was added with calcium carbonate to adjust the solution pH to 5.0 to remove iron and aluminum. After solid-liquid separation, a liquid after iron and aluminum removal and iron-aluminum slag were obtained. After iron and aluminum removal, the liquid is first extracted with P204 extractant to remove manganese, calcium, zinc, and copper impurities. Acid stripping is then performed to obtain a mixed solution of calcium sulfate, manganese sulfate, and copper sulfate. This mixed solution is then extracted with C272 extractant and stripped with sulfuric acid to obtain a mixed solution of manganese sulfate and copper sulfate. Manganese powder is then added to remove copper to obtain a manganese sulfate solution. The manganese sulfate solution is degreased and evaporated to crystallize to obtain manganese sulfate. The raffinate after calcium, manganese, copper, and zinc extraction is then extracted with P507 extractant to extract cobalt. After stripping, a cobalt sulfate solution is obtained. After degreasing, cobalt sulfate is evaporated and crystallized to obtain cobalt sulfate. The raffinate after cobalt extraction is then extracted with C272 extractant to extract magnesium, obtaining a purified nickel sulfate solution. This solution is then degreased with activated carbon and evaporated to obtain nickel sulfate.

[0133] Test example

[0134] The nickel sulfate and cobalt sulfate obtained in Examples 1-3 and Comparative Examples 1-2 were used for comparison. The element data in the Examples and Comparative Examples were tested by ICP-AES equipment.

[0135] The following Tables 1 and 2 are performance comparison results of nickel sulfate and cobalt sulfate products prepared in Examples and Comparative Examples. The specific data were obtained by ICP-AES testing.

[0136] Table 1 Comparison of nickel sulfate product performance

[0137] As can be seen from Table 1, there is no significant difference in quality between the nickel sulfate prepared in the example and the nickel sulfate prepared in the comparative example, and both meet the battery-grade nickel sulfate standard.

[0138] Table 2 Comparison of cobalt sulfate product performance

[0139] As can be seen from Table 2, there is no significant difference in quality between the cobalt sulfate prepared in the example and the nickel sulfate prepared in the comparative example, and both meet the battery-grade cobalt sulfate standard.

[0140] Furthermore, taking Example 1 and Comparative Examples 1-2 as examples, the costs of extracting 1 ton of nickel from laterite nickel ore raw materials to prepare nickel sulfate were compared, and the results are shown in Table 3.

[0141] Table 3 Cost comparison results

[0142] As can be seen from Table 3, in Example 1, due to the advantages of the ion exchange resin having a regeneration function and the ion exchange membrane being recyclable, the total cost is 19,753 yuan / ton. In Comparative Example 1, since nickel carbonate is used to neutralize the residual acid in the analytical solution, nickel carbonate is synthesized using nickel sulfate and soda ash, which increases the cost of raw and auxiliary materials. Compared with Example 1, the cost increases by 2,467 yuan / ton.

[0143] In Comparative Example 2, sodium hydroxide is directly added to the laterite nickel ore leachate to precipitate nickel and cobalt. Since impurities inevitably precipitate, the sodium hydroxide consumption is large, so the cost of raw materials is the highest. Compared with Example 1, the total cost increases by 5,679 yuan / ton.

[0144] Therefore, the method provided by the present disclosure can greatly reduce production costs while effectively obtaining battery-grade nickel sulfate and cobalt sulfate. Industrial Applicability

[0145] The method for resource recovery of laterite nickel ore provided by the present disclosure is simple to operate, can effectively recover substances such as nickel, cobalt, manganese and zinc from laterite nickel ore, and has the advantages of simple process, low production cost, stability and reliability.

Claims

1. A method for the resource utilization of laterite nickel ore, characterized in that, it comprises the following steps: Subject the laterite nickel ore to high-pressure acid leaching, followed by solid-liquid separation to obtain a leaching solution and a leaching residue; adjust the pH value of the leaching solution to 2.0 - 5.0, and then perform solid-liquid separation to obtain an adjusted solution and an adjusted residue; pass the adjusted solution through an ion exchange resin for ion exchange to obtain an adsorption resin that adsorbs nickel ions in the adjusted solution and the remaining adsorption residual solution; use acid to desorb the adsorption resin, and perform ion exchange membrane treatment on the nickel-containing desorption acid solution obtained after desorption to separate nickel salts from the acid, thereby obtaining a nickel salt solution to be purified.

2. The method according to claim 1, characterized in that, it further comprises: Perform a first extraction on the nickel salt solution to be purified to remove calcium, manganese, copper, and zinc in the nickel salt solution, thereby obtaining a first raffinate; Perform a second extraction on the first raffinate to remove magnesium in the first raffinate, thereby obtaining a purified nickel salt solution; Perform degreasing and evaporation crystallization on the purified nickel salt solution to obtain nickel salts.

3. The method according to claim 1 or 2, characterized in that, The pH value of the leaching solution is adjusted by adding additional laterite nickel ore to the leaching solution.

4. The method according to any one of claims 1 - 3, characterized in that, The high-pressure acid leaching includes at least one of the following characteristics: Characteristic one: The acid used includes at least one of sulfuric acid, hydrochloric acid, and nitric acid; Characteristic two: The dosage of the acid used is 8 - 12 times the theoretical amount of the reaction of nickel and cobalt in the laterite nickel ore with sulfuric acid; Characteristic three: The pressure is 3 MPa - 6 MPa; and / or, the temperature is 220 °C - 280 °C; and / or, the time is 2 h - 6 h.

5. The method according to any one of claims 1 - 4, characterized in that, The ion exchange of the ion exchange resin is carried out in a continuous ion exchange device.

6. The method according to claim 5, characterized in that, The continuous ion exchange device includes an adsorption zone, a water washing zone, a desorption zone, and a backwashing zone. The ion exchange of the ion exchange resin includes at least one of the following characteristics: Characteristic one: The ion exchange resin includes a hydrogen ion exchange resin and a sodium ion exchange resin; Characteristic two: The filling amount of the resin in each resin column is 70 vt% - 85 vt%; Characteristic three: The flow rate in the adsorption zone is 6 BV / h - 12 BV / h; Characteristic four: The acid used in the desorption zone includes at least one of sulfuric acid, hydrochloric acid, and nitric acid. The concentration of the acid used in the desorption zone is 3 mol / L - 6 mol / L, and the flow rate in the desorption zone is 2 BV / h - 5 BV / h; Characteristic five: The detergents used in the water washing zone and the backwashing zone are water, and the flow rates of the detergents in the water washing zone and the backwashing zone are independently 3 BV / h - 6 BV / h.

7. The method according to any one of claims 1 - 6, characterized in that, The ion exchange membrane treatment includes at least one of the following characteristics: Characteristic one: The ion exchange membrane includes at least one of an anion exchange membrane, a cation exchange membrane, and an amphoteric ion exchange membrane; Feature 2: The temperature for the ion exchange membrane treatment is 20°C - 60°C; Feature Three: The dialysis coefficient treated by the ion exchange membrane is 1.0×10 -7 m / s - 30×10 -7 m / s.

8. The method according to any one of claims 1 - 7, wherein, the extractant used in the first extraction includes P204, and / or, the extractant used in the second extraction includes C272; and / or, the defoamer used for oil removal includes activated carbon.

9. The method according to any one of claims 1 - 8, wherein, it further includes: returning the adjusted slag to the high-pressure acid leaching process.

10. The method according to any one of claims 1 - 9, wherein, it further includes: returning the acid separated by the ion exchange membrane treatment to the high-pressure acid leaching process.

11. The method according to any one of claims 1 - 10, wherein, it further includes: Mix the adsorption residual liquid with an oxidant to oxidize Fe in the adsorption residual liquid 2+ to Fe 3+ ; then mix it with an alkaline reagent, and perform solid-liquid separation to remove Al 3+ , Fe 3+ , Sc 3+ , Cu 2+ and Cr 3+ in the adsorption residual liquid to obtain a decontaminated liquid and decontamination slag.

12. The method according to claim 11, wherein, the oxidant includes at least one of hydrogen peroxide, manganese dioxide, and sodium persulfide.

13. The method according to claim 11, wherein, the alkaline reagent includes at least one of calcium carbonate, magnesium oxide, magnesium hydroxide, calcium hydroxide, sodium hydroxide, sodium carbonate, and sodium sulfide.

14. The method according to claim 13, wherein, Removal of Al 3+ , Fe 3+ , Sc 3+ and Cr 3+ is carried out under the condition that the pH is 4.0 - 7.

0.

15. The method according to any one of claims 11 - 14, wherein, it further includes: performing extraction and back-extraction on the post-impurity removal solution to obtain a manganese-containing solution, a zinc-containing solution, and an extracted solution containing cobalt; performing oil removal and evaporation crystallization on the manganese-containing solution and the zinc-containing solution to obtain manganese salts and zinc salts respectively.

16. The method according to claim 15, wherein, the extractant used for extracting the post-impurity removal solution includes P204, and the back-extraction agent used for back-extraction includes sulfuric acid.

17. The method according to claim 15 or 16, wherein, it further includes: performing extraction and back-extraction on the extracted solution containing cobalt to obtain a cobalt-containing solution, and performing oil removal and evaporation crystallization on the cobalt-containing solution to obtain cobalt salts.

18. The method according to claim 17, wherein, the extractant used for extracting the extracted solution includes P507, and the back-extraction agent used for back-extraction includes sulfuric acid.

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