Integrated wet method for low-waste high-valued recovery of lithium iron phosphate battery mixed powder

Abstract

The invention discloses an integrated wet method for low-waste and high-valued recovery of lithium iron phosphate battery mixed powder, and belongs to the technical field of waste lithium battery recovery. According to the method, a ferric chloride-chlorine coupled multi-stage circulating leaching technology is adopted, the waste lithium iron phosphate positive and negative electrode mixed powder is leached, and leaching liquid and leaching residues enriched with Li, Al, Fe and Cu are synchronously obtained. Reducing the leachate by iron powder to recover copper, performing precipitation treatment by sodium hydroxide, and separating to obtain iron-aluminum-containing precipitate and lithium-sodium-containing mother liquor; carrying out alkali dissolution on the iron-containing aluminum precipitate to obtain a sodium metaaluminate solution and a ferric hydroxide precipitate; the ferric hydroxide is recycled after being dissolved in acid; leaching residues are subjected to acid dissolution to achieve separation of iron phosphate and graphite, iron phosphate conversion liquid is subjected to multi-stage counter-current extraction and reverse extraction, and ferric chloride and phosphoric acid are recycled respectively. According to the process, consumption of fresh chemicals and waste emission are remarkably reduced, the comprehensive recovery rate of metal is increased, and good environmental benefits and economic benefits are achieved.

Description

Technical Field

[0001] This invention relates to the field of efficient recycling technology for waste lithium iron phosphate batteries, specifically to an integrated wet process for recycling mixed powder of lithium iron phosphate batteries with low waste and high value. Background Technology

[0002] With the rapid development of the new energy vehicle and energy storage industries, lithium iron phosphate (LiFePO4, LFP) batteries have been widely used in the power and energy storage fields due to their advantages such as good thermal stability, low cost, and long cycle life. In recent years, as the first batch of LFP batteries installed in vehicles enters their retirement period, the demand for waste battery treatment and resource recycling has grown rapidly. LFP batteries are rich in strategic resources such as lithium, iron, and phosphorus. If they are not effectively recycled, it will not only waste resources but also potentially cause environmental pollution. Therefore, the development of efficient and green resource recycling technologies is particularly urgent.

[0003] Currently, the mainstream recycling technology is the wet recycling process, which is safe and has low energy consumption. It mainly recovers lithium through leaching and impurity removal, and the main components prepared are battery-grade lithium carbonate and iron phosphate. CN116553510B discloses a method for recycling and regenerating waste lithium iron phosphate powder. This method involves lithium extraction through acid leaching and the preparation of iron phosphate using high-temperature roasting and hydrothermal phosphorus replenishment. While this method can achieve closed-loop regeneration of LFP cathode materials, it involves high energy consumption, large carbon emissions, and significant pressure from alkali consumption and wastewater treatment, resulting in insufficient environmental friendliness. CN110474123A discloses a method for recycling waste LFP cathode materials. This method achieves lithium recovery through a process of salt solution leaching + alkali precipitation for impurity removal + acid regeneration. However, it faces problems such as cumbersome leaching agent regeneration steps, high reagent consumption, and accumulation of solution impurities. CN113912033A discloses a method for recycling mixed positive and negative electrode powder from waste lithium iron phosphate batteries with pre-extraction lithium extraction. This method successively recovers high-concentration iron phosphate dihydrate through processes such as front-end liquid alkali aluminum removal, selective lithium extraction with sulfuric acid and oxidant, high-concentration acid leaching of iron phosphate, and graphite mixed slag. Although this method improves lithium recovery, it consumes large amounts of alkali and oxidant, resulting in high costs, and the residual impurities in the leachate affect the quality of lithium carbonate. CN119220819B discloses a method for preparing battery-grade lithium carbonate and battery-grade iron phosphate using waste powder from the positive and negative electrodes of lithium iron phosphate batteries. This method utilizes inorganic acids, oxidants, and modifiers (mainly soluble iron, calcium, and magnesium salts) for selective lithium extraction, yielding a high-concentration lithium solution. After deep purification with sodium hydroxide, battery-grade lithium carbonate is produced. The mixed slag of iron phosphate and graphite is converted into battery-grade iron phosphate after acid dissolution and deep purification. However, the deep purification process with sodium hydroxide not only consumes a large amount of alkali, but the excessively high sodium ion concentration in the solution also significantly increases the difficulty of preparing battery-grade lithium carbonate. CN120728070B discloses a method for cyclic leaching of mixed lithium iron phosphate powder with iron salts and full-component recovery. This method extracts lithium through salt leaching and hydrogen peroxide oxidation. The oxidation process requires the addition of hydrogen peroxide and hydrochloric acid, resulting in high costs for industrial applications and stringent oxidation conditions.

[0004] Furthermore, with the increasing demand for recycling non-lithium resources from waste positive and negative electrode powders, iron phosphate, as a precursor to lithium-ion battery cathode materials, is gradually attracting attention for its recycling and reuse. However, the downstream applications of iron phosphate are relatively concentrated, mainly targeting a few lithium-ion battery cathode material manufacturers. These manufacturers have strict requirements on the purity, particle size, specific surface area, and other physicochemical properties of the product, resulting in a narrow sales channel for its recycled products and difficulty in consistently realizing its economic added value. In actual recycling processes, iron phosphate in waste lithium iron phosphate powder often mixes with graphite and impurity metal ions (such as Al). 3+ Cu 2+The coexistence of iron phosphate and other components leads to severe intercalation between iron phosphate and other components. Even with partial dissolution through acid dissolution, iron and phosphorus impurities or fine particles that are difficult to settle and separate are easily present, resulting in low separation efficiency, poor product purity, and limited recyclability. Although some processes attempt to achieve structural reconstruction through calcination or thermal decomposition, they generally suffer from high energy consumption, large carbon emissions, and high risks of secondary pollution, making them difficult to promote under large-scale production conditions.

[0005] In summary, current wet recycling technology for mixed powders from spent lithium iron phosphate batteries still faces a series of key challenges and bottlenecks, mainly concentrated in the following aspects:

[0006] (1) Imbalance between economic efficiency and environmental protection in the leaching process: Existing technologies generally rely on external oxidants such as hydrogen peroxide to achieve leaching or regeneration. Not only are the reagent costs high, but their decomposition or reaction can easily generate saline wastewater, increasing the burden of subsequent treatment and environmental protection costs, thus failing to form an economic closed loop.

[0007] (2) The recovery path of ferric phosphate component is singular and the added value is low: the mainstream process aims to recover ferric phosphate as a low-value precursor, but the market capacity of its products is limited and the quality control requirements are stringent. In the mixed powder, ferric phosphate is tightly intercalated with graphite and impurity metals. Simple acid dissolution easily leads to difficulties in separating iron and phosphorus and poor product purity. On the other hand, the method of high-temperature roasting and reconstruction faces the contradiction of excessive energy consumption and carbon emissions, and fails to achieve the high-value-added directional conversion of this component.

[0008] (3) Insufficient process coordination and large amount of waste liquid: Most processes are to recover different components in steps, with loose links between links. The pH needs to be adjusted repeatedly and a large amount of acid and alkali is consumed for deep impurity removal, resulting in large reagent consumption, easy accumulation of liquid phase impurities, and large amount of waste liquid. The integration and greening level of the overall process need to be improved.

[0009] (4) Resource recycling is “complete but not excellent” and the overall economic benefits are not good: Although existing technologies have focused on lithium recycling, they have not given enough consideration to the synergistic recycling and high-value utilization of other valuable components such as iron, phosphorus, aluminum, copper and graphite, or there is a dilemma that the recycling rate and product value are difficult to balance. A complete solution that can efficiently extract all valuable components and significantly reduce the overall cost and environmental footprint through internal material recycling has not yet been formed.

[0010] Therefore, there is an urgent need to develop an integrated wet process that can simultaneously achieve efficient regeneration and recycling of leaching agents, high-value conversion of iron phosphate components, synergistic recovery of all elements, and low wastewater discharge. This is the key to breaking through current technical bottlenecks and improving the economic and environmental benefits of recycling waste lithium iron phosphate batteries, and it is also the core starting point of this invention. Summary of the Invention

[0011] This invention aims to overcome the shortcomings of existing technologies and provide an integrated wet process for the low-waste, high-value recycling of lithium iron phosphate battery powder. This method constructs a closed-loop "leaching-regeneration-leaching" system and innovatively converts iron phosphate into ferric chloride and phosphoric acid products, achieving efficient leaching of valuable metals, in-situ regeneration of the leaching agent, and targeted recovery of all components. This significantly reduces raw material and energy costs, and decreases emissions of waste gas, wastewater, and solid waste, resulting in good environmental and economic benefits.

[0012] The technical solution of this invention is to propose an integrated wet process for the low-waste, high-value recycling of lithium iron phosphate battery mixed powder, which includes:

[0013] S1: Using ferric chloride solution as the leaching agent, the mixture of positive and negative electrode powders from waste lithium iron phosphate batteries is stirred and reacted to obtain a leachate containing Li, Fe, Al, and Cu and an insoluble residue; chlorine gas is introduced into the leachate to oxidize and regenerate ferrous iron into ferric iron, which is then used in the next round of leaching. The oxidation, regeneration, and leaching are repeated until the preset number of cycles are reached.

[0014] The leaching residues obtained from each round are grouped and leached again using the leaching agent regenerated in the last round to obtain the final leachate and the final leaching residue;

[0015] S2: Iron powder is added to the final leachate to replace crude copper; the crude copper is treated with acid and separated to obtain industrial-grade copper powder and an acidic solution containing ferrous ions; chlorine gas is passed into the solution for oxidation; then alkali is added to selectively precipitate Fe and Al to obtain iron-aluminum precipitate and lithium-rich mother liquor; the mother liquor is used for the preparation of battery-grade lithium carbonate; the iron-aluminum precipitate is added to alkali solution and separated to obtain iron hydroxide precipitate and sodium aluminate solution; the former is acid-dissolved and reused as a leaching agent, and the latter is concentrated and crystallized to obtain sodium aluminate product;

[0016] S3: The final leaching residue is dissolved in hydrochloric acid to separate ferric phosphate solution and crude graphite; the crude graphite is washed and calcined in an inert atmosphere to obtain high-purity graphite;

[0017] The ferric phosphate solution was subjected to multi-stage countercurrent extraction, and the resulting organic phase was back-extracted to obtain a ferric chloride solution. After evaporation and concentration, ferric chloride hexahydrate was obtained. The raffinate was subjected to multi-stage countercurrent extraction to obtain a phosphorus-loaded organic phase. The phosphorus-loaded organic phase was back-extracted to obtain a phosphoric acid solution, which was then concentrated to prepare industrial-grade phosphoric acid.

[0018] The ferric phosphate solution (containing iron and phosphoric acid) obtained after acid dissolution is key to achieving high-value production in this invention. In some embodiments of this invention, the treatment process of the ferric phosphate solution in S3 includes: firstly, adjusting the pH of the ferric phosphate solution to 0.1–0.5, and then performing multi-stage countercurrent extraction using a first extraction system with di(2-ethylhexyl) phosphate (P204) as the main component to selectively extract iron ions to the organic phase, achieving iron-phosphorus separation. The iron-loaded organic phase is back-extracted with hydrochloric acid to obtain a pure ferric chloride solution, which can be evaporated and crystallized to obtain ferric chloride hexahydrate. The raffinate from the previous multi-stage countercurrent extraction mainly contains phosphate ions. After adjusting its pH to -0.1–-0.5, a second extraction system with tributyl phosphate (TBP) as the main component is used for multi-stage countercurrent extraction to enrich phosphoric acid. The phosphoric acid-loaded organic phase is then back-extracted with water and concentrated to obtain industrial-grade phosphoric acid. Both organic phases after extraction can be regenerated and recycled.

[0019] This invention utilizes the isomorphic induction effect of ferric chloride to rapidly leach lithium iron phosphate materials. Simultaneously, it combines this with direct oxidation by chlorine to achieve multiple regenerations of the leaching agent. Solid-liquid separation yields graphite-containing FePO4 leaching residue and a high-concentration lithium chloride leaching solution containing iron, copper, and aluminum. The high-concentration lithium chloride leaching solution is then subjected to a series of precipitation processes with sodium hydroxide solution to remove impurities, resulting in a high-purity lithium chloride solution. The graphite-containing FePO4 leaching residue is first acid-dissolved to obtain regenerated graphite. Then, the pH of the ferric phosphate solution is adjusted using acid to extract iron and phosphorus. After back-extraction, washing, and other post-processing, industrially marketable ferric chloride and phosphoric acid products are obtained. This effectively achieves the complete recovery of Li, Fe, P, Cu, Al, and graphite from waste lithium iron phosphate battery powder.

[0020] The beneficial effects of this invention are:

[0021] (1) Addressing the core pain points of traditional wet process, such as high reagent consumption, large waste liquid volume, and high cost, this invention innovatively adopts a "ferric chloride-chlorine gas coupled multi-stage circulating leaching" mechanism. Its core advantage lies in completely eliminating the expensive hydrogen peroxide oxidation method that generates saline wastewater. This system not only has high oxidation efficiency and thorough regeneration, but also forms a closed-loop path of "leaching-regeneration-leaching," allowing the leaching agent (FeCl3) to be repeatedly recycled. Experimental verification shows that after 10 consecutive cycles, the total lithium leaching rate is still as high as 99.3%, proving its excellent long-term operational stability, while significantly reducing the consumption of fresh chemicals and wastewater discharge load from the source.

[0022] (2) Unlike the traditional approach that only recovers lithium or iron phosphate, this invention establishes a complete recycling technology chain covering all valuable components such as Li, Fe, P, Cu, Al, and graphite. In particular, it breaks through the limitations of iron phosphate utilization: through an "acid dissolution-extraction" process, the iron phosphate precursor, with its limited market capacity, is selectively converted into widely used and high-value-added ferric chloride hexahydrate and industrial-grade phosphoric acid. The recovery rates of iron and phosphorus both exceed 99%, greatly expanding their downstream applications and overcoming the limitation of existing processes where iron phosphate products are restricted to a single sales channel by battery companies, thereby increasing the economic value-added of the entire recycling process.

[0023] (3) In the key steps of the high-value conversion of iron phosphate, this invention adopts a stepwise synergistic strategy of P204 and TBP extraction systems. By precisely controlling the pH, iron is first extracted efficiently and selectively using the P204 system under strong acidity, with a high separation coefficient. A pure FeCl3 solution can be obtained by back-extraction with hydrochloric acid. Then, phosphoric acid is extracted using the TBP system under certain pH conditions, and a phosphoric acid solution is obtained by back-extraction with water. The organic phases after both extraction stages can be regenerated and recycled, resulting in low reagent consumption. This unit technology not only achieves high-purity separation of iron and phosphorus (with less impurity entrainment), but also reduces the generation of secondary waste through internal circulation, which is an important guarantee for the green and continuous operation of the entire process.

[0024] (4) The entire process demonstrates a high degree of system integration: the leaching agent achieves closed-loop circulation, the extractant is internally regenerated, forming a multi-level material circulation network, which systematically reduces operating costs and environmental risks. In addition, in response to the challenge of large fluctuations in raw material composition and particle size in actual recycling, the innovative "group leaching-re-leaching" dynamic strategy of this invention can significantly enhance the process's adaptability and processing capacity to raw materials by flexibly adjusting the number of batches and reaction parameters, ensuring the stability of the recovery rate and the robustness of the process, laying a solid foundation for large-scale engineering applications. Attached Figure Description

[0025] Figure 1 This is a complete process flow diagram for the recycling of the method of the present invention.

[0026] Figure 2 This is a flow chart of the multi-cycle leaching process in this invention.

[0027] Figure 3 This is a flow chart of the final high-value recovery process of the leachate in this invention.

[0028] Figure 4 This is a flow chart of the final high-value recovery process of the leaching residue in this invention. Detailed Implementation

[0029] The technical solution of the present invention will be further described below with reference to the accompanying drawings and specific embodiments.

[0030] The examples provided are for illustrative purposes only and are not intended to limit the scope of the invention. Technical features of each embodiment in this invention can be combined appropriately without conflict. All embodiments obtained based on the examples in this invention without making any inventive breakthroughs are within the protection scope of this invention.

[0031] This invention proposes a low-cost, high-value wet process for recycling mixed lithium iron phosphate battery powder, which mainly includes a circulating leaching step, a leaching solution impurity removal step, and a leaching residue resource utilization step. Figure 1 This is a flow chart of the entire recycling process, including the multi-cycle leaching process as follows: Figure 2 As shown, the process flow is as follows: Figure 3 As shown, the final high-value recovery process of the leaching residue is as follows: Figure 4 As shown. See also Figures 1 to 4 As shown, the specific process flow of this invention is as follows:

[0032] S1. Cyclic Leaching: The positive and negative electrode mixed powder obtained after discharging, crushing, and sorting waste lithium iron phosphate batteries is used as raw material. The positive and negative electrode mixed powder can be obtained from a battery dismantling plant and is allowed to contain small amounts of residual electrolyte, binder, and current collector metal. In one embodiment of the invention, the main components of the positive and negative electrode mixed powder are lithium iron phosphate and graphite, and it contains small amounts of binder, current collector (aluminum foil, copper foil), and impurities such as residual electrolyte, wherein the copper content is ≤15 wt.% and the aluminum content is ≤10 wt.%.

[0033] FeCl3 solution was used as the leaching agent and mixed with the mixed powder. The molar ratio of ferric ions in the leaching agent to lithium in the mixed powder was controlled at 1.0~2.0, and the solid-liquid ratio of the mixed powder to the leaching agent was controlled at 100~600 g / L. The mixture was stirred at 20~60℃ for 20~60 min to transfer valuable metal ions such as lithium, iron, aluminum, and copper into the solution, thereby achieving the decomposition of lithium iron phosphate. The main leaching reactions are as follows:

[0034]

[0035] After the reaction was completed, solid-liquid separation was performed to obtain Li-rich products. + Fe 2+ / Fe 3+ Al 3+ Cu 2+ The leachate and the leachate residue, which mainly contains ferric phosphate and graphite.

[0036] Chlorine gas (Cl2) is introduced into the aforementioned leachate, and an oxidative regeneration reaction is carried out under stirring. The chlorine gas can be derived from by-product chlorine gas in the chlor-alkali industry, achieving resource utilization of industrial waste. Before chlorination, the solution pH is adjusted to <1, and the chlorine gas flow rate is controlled to be 0.5~1.0 times the theoretical molar amount of ferrous ions in the solution. The reaction temperature is 40~80℃, and the time is 60~120 min. The chlorine gas reacts with the ferrous ions (Fe2+) generated during the leaching process. 2+ ) is re-oxidized to ferric ions (Fe) 3+ To restore the oxidizing capacity of the leachate, the main regeneration reaction is as follows:

[0037]

[0038] The resulting regenerated leachate is used to leach the next batch of mixed powder. This "leaching-oxidation regeneration" process can be repeated a preset number of times (such as 5 or 10 times) until the activity of the leachate reaches a set threshold, thereby significantly improving the utilization rate of the leachate.

[0039] Collect and combine the leaching residue generated from multiple leaching cycles. Using the regenerated leaching agent remaining after the last cycle, group and batch-wise re-leaching treatment is performed on the leaching residue. For example, in a process with a solid-liquid ratio of 200 g / L and 10 leaching cycles, the total leaching residue can be divided into three groups (e.g., at a mass ratio of 3:3:4) and re-leached sequentially; in a process with a solid-liquid ratio of 400 g / L and 5 cycles, the leaching residue can be divided into two groups (e.g., at a mass ratio of 3:2) and re-leached. The re-leaching reaction temperature is controlled at 20–60 °C, and the time is 20–60 min.

[0040] S2. Leachate purification: Add metallic iron powder to the final leachate to replace the Cu. 2+ The amount of iron powder added is equal to the Cu content in the leaching solution. 2+ The molar amount is 1.0 to 2.0 times. After the reaction is complete, solid-liquid separation is performed to obtain crude copper and a copper-removed liquid containing ferrous ions. The crude copper is acid-washed with hydrochloric acid at a concentration of 0.5 to 2.0 mol / L to remove surface impurities. After solid-liquid separation and drying, industrial-grade copper powder is obtained. The acid washing solution is then added to the copper-removed liquid.

[0041]

[0042] Chlorine gas is introduced into the combined ferrous ion-containing solution to oxidize the ferrous ions to ferric ions. The molar amount of chlorine gas introduced is 0.5 to 1.0 times the molar amount of ferrous ions in the solution. The oxidation reaction temperature is 40 to 80℃ and the gas introduction time is 60 to 120 min.

[0043] Subsequently, solid sodium hydroxide was added to the oxidized solution to initiate a precipitation reaction. The amount of solid sodium hydroxide added was 1.0 to 1.5 times the total molar amount of iron and aluminum ions in the solution, causing Fe... 3+ And Al 3+ Selective precipitation yields hydroxides. Solid-liquid separation is performed after the reaction to obtain a mixed iron-aluminum precipitate and a Li-rich precipitate. + The mother liquor, which can be directly used for the subsequent preparation of battery-grade lithium carbonate, is obtained. The resulting iron-aluminum mixed precipitate is reacted with a sodium hydroxide solution of 0.5–5.0 mol / L, the amount of sodium hydroxide being 1.2–2.0 times the theoretical molar amount of aluminum in the precipitate. This causes the aluminum hydroxide to dissolve, forming sodium aluminate, while the iron hydroxide remains unchanged. Solid-liquid separation is performed again to obtain industrial-grade solid iron hydroxide and sodium aluminate solution, respectively. The sodium aluminate solution is evaporated, concentrated, and crystallized to finally obtain the solid sodium aluminate product.

[0044]

[0045]

[0046]

[0047]

[0048] S3. Resource utilization of leaching residue: The final leaching residue obtained in step S1 (mainly composed of ferric phosphate and graphite) is mixed with a hydrochloric acid solution with a concentration of 3-5 mol / L, and the mixture is subjected to acid dissolution reaction at 40-80℃ for 40-120 min. After the reaction is completed, solid-liquid separation is performed to obtain crude graphite and ferric phosphate solution.

[0049]

[0050] The obtained crude graphite was washed with deionized water until neutral, and then calcined at 450~700℃ for 1~3 h under N2 protection to remove residual organic matter and obtain high-purity regenerated graphite.

[0051] The pH of the ferric phosphate solution was adjusted to 0.1–0.5. Di(2-ethylhexyl) phosphate (P204) diluted with kerosene was used as the first extractant, with a volume fraction of 20%–40%. The organic phase was mixed with the aqueous phase at a volume ratio of (1–3):1, and a four-stage countercurrent extraction was performed at 25–40°C, with each stage lasting 5–10 min, selectively extracting iron ions from the solution. After extraction, the iron-loaded organic phase and the raffinate, mainly composed of phosphoric acid, were separated. The iron-loaded organic phase was back-extracted using a 0.5–6 mol / L hydrochloric acid solution. During back-extraction, the organic phase to aqueous phase volume ratio was 1:(2–5), and a ten-stage countercurrent back-extraction was performed at 35–50°C, with each stage lasting 5–10 min, yielding a pure ferric chloride aqueous solution. This solution was then evaporated, concentrated, and crystallized to obtain ferric chloride hexahydrate. The organic phase after back-extraction is returned to the extraction process for recycling.

[0052] The pH of the phosphorus-containing raffinate was adjusted to -0.1 to -0.5. Tributyl phosphate (TBP) diluted with kerosene was used as the second extractant, with a TBP volume fraction of 60%–80%. The organic phase was mixed with the aqueous phase at a volume ratio of (3–6):1, and subjected to 8 stages of countercurrent extraction at 25–40°C, with each stage lasting 5–10 min, selectively extracting phosphoric acid from the solution. The phosphoric acid-loaded organic phase was then separated after extraction. The phosphoric acid-loaded organic phase was back-extracted with water. During back-extraction, the organic phase to aqueous phase volume ratio was (5–20):1, and 8 stages of countercurrent back-extraction were performed at 35–50°C, with each stage lasting 5–10 min, yielding a dilute phosphoric acid solution. This dilute phosphoric acid solution was concentrated to obtain industrial-grade phosphoric acid. The back-extracted organic phase was also returned to the extraction process for recycling.

[0053] The overall method will be described in detail below with reference to several specific embodiments.

[0054] Example 1

[0055] (1) Circulating leaching

[0056] 8 g of mixed black powder was added to 40 mL of 1.32 mol / L FeCl3 solution (iron-lithium molar ratio 1.50, solid-liquid ratio 200 g / L), and stirred in a 40℃ water bath for 30 min. After leaching, solid-liquid separation was performed, and the pH of the solution was adjusted to <1. The leachate was then placed in a 100 mL polytetrafluoroethylene high-temperature and high-pressure reactor, and chlorine gas was introduced, with the amount of chlorine gas controlled to be 0.8 times the molar amount of lithium in the battery powder. The reaction temperature was maintained at 40℃ for 60 min. The stirring rate was 500 rpm during the reaction. The amount of ferrous ions in the leachate after the reaction was determined by potassium dichromate titration. The results showed that there were basically no ferrous ions remaining in the solution, proving that Fe... 2+ Completely oxidized to Fe 3+ This enables the regeneration of the leachate.

[0057] The second batch of 8 g mixed powder was treated with the regenerated FeCl3 solution, and the leaching operation was repeated for a total of 10 cycles. The leaching residue from the first three cycles was collected as Group A, the leaching residue from cycles 4-6 was collected as Group B, and the leaching residue from cycles 7-10 was collected as Group C. These three groups were then oxidatively leached sequentially with the leachate after 10 cycles, without adjusting the solid-liquid ratio. The final leachate and leaching residue were collected and analyzed by inductively coupled plasma atomic emission spectrometry (ICP-OES). The concentrations of each metal ion in the final leachate were: Li 60.57 g / L, Cu 118.07 g / L, Fe 66.53 g / L, and Al 19.98 g / L. The total Li recovery rate was 99.3%, the total Cu recovery rate was 78.3%, and the Al recovery rate was 80.6%.

[0058] (2) Copper removal

[0059] At room temperature, reducing iron powder was added to the final leachate at a molar ratio of n(Fe):n(Cu) = 1.2, and the mixture was mechanically stirred for 30 min to complete the displacement reaction. After solid-liquid separation and analysis, copper ions in the lithium-containing mother liquor were completely removed (displacement rate 100%). Subsequently, the obtained crude copper was acid-washed with 2 mol / L hydrochloric acid at a solid-liquid ratio of 400 g / L, followed by mechanical stirring and washing at room temperature for 1 h to remove excess iron powder. After the above process, the concentrations of each ion in the leachate were: Li 40.38 g / L, Fe 113.2 g / L, and Al 13.3 g / L, ultimately yielding copper powder with a purity of 99.8% and a recovery rate of 96.9%.

[0060] (3) Removal of iron and aluminum

[0061] Chlorine gas is introduced into the combined solution, and the amount of chlorine gas introduced is equal to the amount of Fe in the solution. 2+ 0.8 times the molar amount, reacted at 40℃ for 60 min, so that Fe2+ Completely oxidized to Fe 3+ Solid NaOH was added to the oxidized leachate, with the amount of sodium hydroxide added controlled to be 1.4 times the molar ratio of iron to aluminum, forming a co-precipitate of Fe(OH)3 and Al(OH)3. After centrifugation, a lithium-containing supernatant and metal precipitate were obtained. The Fe concentration in the post-precipitation solution was 0.02 g / L, the Al concentration was 0.06 g / L, the iron precipitation rate was 99.7%, and the aluminum precipitation rate was 92.5%, proving that the iron and aluminum precipitation was relatively complete. 1 mol / L NaOH solution was added to the mixed precipitate and stirred at room temperature for 30 min to further react with Al(OH)3. Solid-liquid separation was then performed to obtain sodium aluminate solution and solid iron hydroxide. The amount of sodium hydroxide added was approximately 1.3 times the molar amount of aluminum in the precipitate. The remaining iron hydroxide precipitate was dissolved in 3 mol / L hydrochloric acid, then concentrated under reduced pressure. The ferric chloride solution was recovered for a new round of leaching, forming a closed-loop utilization of the leaching agent.

[0062] (4) High-value recovery of leaching residue

[0063] A 4 mol / L dilute hydrochloric acid solution was added to the final leaching residue at a solid-liquid ratio of 200 g / L, and the reaction was carried out at 60 °C for 1.5 h. Solid-liquid separation was performed to obtain crude graphite and a ferric phosphate solution. The crude graphite was washed repeatedly with small amounts of 0.1 mol / L dilute hydrochloric acid, and then calcined at 600 °C for 2 h under a nitrogen atmosphere to remove binders such as PVDF, yielding relatively pure recycled graphite.

[0064] The washing solution was mixed with ferric phosphate solution to form ferric phosphate solution. The concentrations of each element were measured as follows: Fe 57.02 g / L, P 31.82 g / L, Al 0.54 g / L, and Cu 3.68 g / L. Sodium hydroxide solid was used to adjust the pH of the solution to 0.5. The extractant was prepared as kerosene-diluted di(2-ethylhexyl) phosphate (P2O4), with a P2O4 volume fraction of 30%. Extraction was performed at room temperature with an organic phase to aqueous phase volume ratio of 2:1 and an extraction time of 5 min. After four stages of countercurrent extraction, the Fe in the ferric phosphate solution was almost completely extracted. The concentrations of each metal ion in the raffinate 1 were: Fe 0.29 g / L, P 30.23 g / L, Al 0.08 g / L, and Cu 2.12 g / L. In this step, more than 99.5% of the iron in the ferric phosphate solution was extracted into the organic phase, while the amount of P extracted did not exceed 5%.

[0065] The iron-containing organic phase was back-extracted using 6 mol / L hydrochloric acid as the aqueous phase, maintaining an organic-to-aqueous phase ratio of 2:1. The process was carried out at 35°C through 10 stages of countercurrent back-extraction to completely remove Fe from the organic phase. Each stage lasted 5 minutes. After back-extraction, the concentrations of each element in the back-extraction solution were as follows: Fe 56.17 g / L, P 1.58 g / L, Al 0.45 g / L, and Cu 1.53 g / L. The overall back-extraction rate reached 99%.

[0066] After the solution is concentrated to 60% of its volume, it is filtered to remove ferric phosphate. Then, it is cooled and crystallized to obtain ferric chloride hexahydrate with a certain purity (purity ≥ 95%).

[0067] Kerosene-diluted tributyl phosphate (TBP) with a volume fraction of 60% was used as the extractant to extract raffinate 1. Before extraction, the pH was adjusted to -0.5 with hydrochloric acid to inhibit the dissociation of phosphate molecules. The ratio of the organic phase to the aqueous phase was 4:1, with 8 extraction stages and an extraction time of 5 min. After extraction, phosphate was extracted into the organic phase. The elemental concentrations in raffinate 2 from this step were: Fe 0.24 g / L, P 0.30 g / L, Al 0.07 g / L, and Cu 2.08 g / L. In this step, over 99% of the phosphorus was extracted into the organic phase, with only a small amount of impurities.

[0068] The phosphorus-containing organic phase was back-extracted using pure water at a volume ratio of 5:1. Each back-extraction stage lasted 5 minutes at a temperature of 35°C. After eight stages of countercurrent back-extraction, a high-purity dilute phosphoric acid solution was obtained. The concentrations of the elements were: Fe 0.04 g / L, P 29.50 g / L, Al 0.003 g / L, and Cu 0.16 g / L, with a calculated phosphoric acid content of 83.06 g / L. Further concentration to 10% of the original volume yielded industrial-grade phosphoric acid with a purity of 85%.

[0069] Example 2

[0070] (1) Circulating leaching

[0071] Eight g of positive and negative electrode mixed powder (with the same composition as in Example 1) obtained from the treatment of waste lithium iron phosphate batteries was mixed with 30 ml of FeCl3 solution to control the Fe content in the leaching agent. 3+The molar ratio of Li in the mixed powder was 1.3, and the solid-liquid ratio was 300 g / L. The reaction was carried out with stirring in a 40℃ water bath for 30 min. After the reaction, the solid and liquid phases were separated, and the pH of the leachate was adjusted to <1 with concentrated hydrochloric acid and transferred to a reactor. Chlorine gas was introduced into the leachate at a rate 0.6 times the theoretical molar amount of iron in the mixed powder, and the oxidation reaction was carried out at 60℃ for 90 min with a stirring rate of 500 rpm. The oxidized solution was tested and found to contain no Fe. 2+ The residue is regenerated by the leaching agent.

[0072] Using this regenerated leaching agent, a new batch of mixed powder was leached under the above conditions, for a total of 8 cycles. The leaching residues from cycles 1-4 were combined into group A, and the leaching residues from cycles 5-8 were combined into group B. Using the regenerated leaching agent after the 8th cycle, without adjusting the solid-liquid ratio, both groups A and B were leached again (40℃, 30 minutes). The two leaching solutions were combined to obtain the final leachate. Inductively coupled plasma atomic emission spectrometry (ICP-OES) analysis showed the following key metal concentrations in the final leachate: Li 48.6 g / L, Fe 58.7 g / L, Al 16.3 g / L, Cu 95.4 g / L. The calculated total leaching rate of Li was 96.1%.

[0073] The cyclic leaching process was the same as in Example 1, except that the oxidation conditions were changed: chlorine gas was introduced, and the amount of chlorine gas introduced was controlled to be 0.5 times the molar amount of lithium in the battery powder. The reaction temperature was maintained at 60°C, and the reaction time was 90 min. After ten cycles of leaching, the concentrations of each metal ion were: Li 58.62 g / L, Cu 106.77 g / L, Fe 68.74 g / L, and Al 18.67 g / L. The total Li recovery rate was 96.6%, the total Cu recovery rate was 70.8%, and the Al recovery rate was 75.3%. This indicates that reducing the amount of chlorine gas introduced during the oxidation process leads to a decrease in the oxidation efficiency, resulting in a decrease in the final leaching rate of each metal ion.

[0074] (2) Copper removal

[0075] Iron powder was added to the final leachate at a molar ratio of n(Fe):n(Cu) = 1.0, and the mixture was stirred at room temperature for 40 min. After solid-liquid separation, the crude copper was washed with 2 mol / L hydrochloric acid at a solid-liquid ratio of 500 g / L for 30 min, and dried to obtain copper powder with a purity of 99.5% and a copper recovery rate of 95.2%. The copper-removed solution was combined with the pickling solution.

[0076] (3) Removal of iron and aluminum

[0077] Chlorine gas is introduced into the combined solution, and the amount of chlorine gas introduced is equal to the amount of Fe in the solution. 2+ 0.8 times the molar amount, reacted at 40℃ for 60 minutes, so that Fe 2+ Completely oxidized to Fe3+ Subsequently, solid sodium hydroxide was added, the amount of which was equal to the Fe content in the solution. 3+ And Al 3+ Add 1.2 times the total molar amount and stir to allow for complete precipitation. After centrifugation, a mixed iron-aluminum precipitate and a lithium-rich mother liquor are obtained. The concentrations of Fe and Al in the mother liquor after precipitation are both below 0.1 g / L.

[0078] The mixed precipitate was mixed with a 1.0 mol / L sodium hydroxide solution, with the amount of NaOH being 1.5 times the molar amount of Al in the precipitate. The mixture was stirred at room temperature for 40 min. After solid-liquid separation, solid ferric hydroxide and sodium aluminate solution were obtained. The sodium aluminate solution was evaporated, concentrated, and cooled to crystallize, yielding white solid sodium aluminate. The solid ferric hydroxide was dissolved in acid and reused as a leaching agent.

[0079] (4) High-value recovery of leaching residue

[0080] The final leaching residue obtained from the cyclic leaching step was mixed with 3 mol / L hydrochloric acid at a solid-liquid ratio of 200 g / L and acid-dissolved at 70℃ for 120 min. Solid-liquid separation was performed to obtain crude graphite and ferric phosphate solution. The crude graphite was washed with water, dried, and then calcined at 650℃ for 2 h in a nitrogen atmosphere to obtain relatively pure regenerated graphite.

[0081] The pH of the ferric phosphate solution obtained from acid dissolution was adjusted to 0.3. Using a 40% (v / v) P204-kerosene solution as the organic phase, a mixture was prepared at a volume ratio of organic phase:aqueous phase (O:A) = 2:1. A four-stage countercurrent extraction was performed at 30°C, with a single-stage contact time of 8 minutes. After extraction, the iron extraction rate was greater than 99.8%. The iron-loaded organic phase was then back-extracted using 4 mol / L hydrochloric acid at an O:A ratio of 1:3, with a ten-stage countercurrent extraction at 45°C, each stage lasting 10 minutes. The resulting FeCl3 solution was evaporated, concentrated, and cooled to crystallize, yielding pale yellow ferric chloride hexahydrate (FeCl3·6H2O) crystals with a purity ≥98%.

[0082] The pH of the phosphorus-containing raffinate after iron extraction was adjusted to -0.3. Using an 80% (v / v) TBP-kerosene solution as the organic phase, an 8-stage countercurrent extraction was performed at room temperature with an O:A ratio of 6:1, with each stage lasting 10 min. The phosphoric acid extraction rate was greater than 99.5%. The phosphoric acid-loaded organic phase was then back-extracted with deionized water at an O:A ratio of 15:1, with an 8-stage countercurrent extraction at 40°C, lasting 10 min per stage. A dilute phosphoric acid solution was obtained, which was then concentrated under reduced pressure to obtain industrial-grade phosphoric acid with a concentration ≥85%.

[0083] The above-described embodiments are merely illustrative of several implementations of the present invention, and while the descriptions are relatively specific and detailed, they should not be construed as limiting the scope of the present invention. Those skilled in the art can make various modifications and improvements without departing from the concept of the present invention, and these modifications and improvements all fall within the scope of protection of the present invention.

Claims

1. An integrated wet process for the low-waste, high-value recycling of lithium iron phosphate battery mixed powder, characterized in that, include: S1: Using ferric chloride solution as the leaching agent, the mixture of positive and negative electrode powders from waste lithium iron phosphate batteries is stirred and reacted to obtain a leachate containing Li, Fe, Al, and Cu and an insoluble residue; chlorine gas is introduced into the leachate to oxidize and regenerate ferrous iron into ferric iron, which is then used in the next round of leaching. The oxidation, regeneration, and leaching are repeated until the preset number of cycles are reached. The leaching residues obtained from each round are grouped and leached again using the leaching agent regenerated in the last round to obtain the final leachate and the final leaching residue; S2: Iron powder is added to the final leachate to replace crude copper; the crude copper is treated with acid and separated to obtain industrial-grade copper powder and an acidic solution containing ferrous ions; chlorine gas is passed into the solution for oxidation; then alkali is added to selectively precipitate Fe and Al to obtain iron-aluminum precipitate and lithium-rich mother liquor; the mother liquor is used for the preparation of battery-grade lithium carbonate; the iron-aluminum precipitate is added to alkali solution and separated to obtain iron hydroxide precipitate and sodium aluminate solution; the former is acid-dissolved and reused as a leaching agent, and the latter is concentrated and crystallized to obtain sodium aluminate product; S3: The final leaching residue is dissolved in hydrochloric acid to separate ferric phosphate solution and crude graphite; the crude graphite is washed and calcined in an inert atmosphere to obtain high-purity graphite; The ferric phosphate solution was subjected to multi-stage countercurrent extraction, and the resulting organic phase was back-extracted to obtain a ferric chloride solution. After evaporation and concentration, ferric chloride hexahydrate was obtained. The raffinate was subjected to multi-stage countercurrent extraction to obtain a phosphorus-loaded organic phase. The phosphorus-loaded organic phase was back-extracted to obtain a phosphoric acid solution, which was then concentrated to prepare industrial-grade phosphoric acid.

2. The method according to claim 1, characterized in that, In step S1, during each leaching round, the molar ratio of ferric ions in the leaching agent to lithium in the positive and negative electrode mixed powder is 1.0~2.0; the positive and negative electrode mixed powder and the leaching agent are leached at a solid-liquid ratio of 100~600g / L, the reaction temperature is 20~60℃, and the reaction time is 20~60 min.

3. The method according to claim 1, characterized in that, In step S1, when oxidizing and regenerating the leachate, the pH of the leachate must first be adjusted to <1; the molar amount of chlorine gas introduced must be controlled to be 0.5 to 1.0 times the molar amount of ferrous ions in the solution; the oxidation temperature must be 40 to 80°C; and the reaction time must be 60 to 120 min.

4. The method according to claim 1, characterized in that, In step S2, the amount of iron powder added is 1.0-2.0 times the copper ion content in the solution; the molar amount of chlorine gas introduced is 0.5-1.0 times the molar amount of ferrous ions in the solution; the oxidation reaction temperature is 40-80℃; and the aeration time is 60-120 min.

5. The method according to claim 1, characterized in that, In step S2, the added alkali is sodium hydroxide; initially, solid sodium hydroxide is added for precipitation, and the amount added is 1.0 to 1.5 times the total molar amount of iron and aluminum ions in the leachate; the alkali solution added to the obtained iron and aluminum precipitate is a sodium hydroxide solution, and the amount used is 1.2 to 2.0 times the molar amount of aluminum in the precipitate; the OH- content in the alkali solution is... - The concentration is 0.5~5.0 mol / L.

6. The method according to claim 1, characterized in that, In step S3, the amount of hydrochloric acid used for acid dissolution is 3 to 6 times the molar amount of iron in the final leaching residue, the concentration is 3 to 5 mol / L, the acid dissolution temperature is 40 to 80℃, and the time is 40 to 120 min; the calcination of crude graphite after acid washing is carried out in a nitrogen atmosphere, the calcination temperature is 450℃ to 700℃, and the calcination time is 1 to 3 h.

7. The method according to claim 1, characterized in that, In step S3, the pH of the ferric phosphate solution is first adjusted to 0.1-0.5, and multi-stage countercurrent extraction is performed using the first extractant to selectively extract iron ions, resulting in an iron-loaded organic phase and a phosphorus-containing raffinate. The first extractant used was kerosene-diluted di(2-ethylhexyl) phosphate, with a volume fraction of 20%-40%. The volume ratio of organic phase to aqueous phase was (1-3):

1. The extraction stage was 4 stages, the single-stage extraction time was 5-10 min, and the extraction temperature was 25-40℃.

8. The method according to claim 1 or 7, characterized in that, In step S3, the back-extraction of the obtained organic phase includes: when back-extracting the iron-loaded organic phase, a hydrochloric acid solution with a concentration of 3~6 mol / L is used as the back-extraction agent, the volume ratio of the organic phase to the aqueous phase is 1:(2~5), the number of back-extraction stages is 10, the single-stage back-extraction time is 5~10 min, and the back-extraction temperature is 35~50℃.

9. The method according to claim 1, characterized in that, In step S3, the multi-stage countercurrent extraction of the obtained raffinate includes: first adjusting the pH of the raffinate to -0.1 to -0.5, and then using a second extractant for multi-stage countercurrent extraction to selectively extract phosphoric acid; the second extractant used is kerosene-diluted tributyl phosphate, the volume fraction of tributyl phosphate is 60% to 80%, the volume ratio of organic phase to aqueous phase is (3-6):1, the number of extraction stages is 8, the single-stage extraction time is 5-10 min, and the extraction temperature is 25-40℃.

10. The method according to claim 1 or 9, characterized in that, In step S3, when back-extracting the phosphorus-loaded organic phase, water is used as the back-extracting agent; the volume ratio of organic phase to water phase is (5~20):1, the number of back-extraction stages is 8, the single-stage back-extraction time is 5~10 min, and the back-extraction temperature is 35~50℃.

Images