Process for the preparation of phosphates

Abstract

The present application relates to a kind of phosphate preparation method, belong to chemical industry technical field.The phosphate preparation method includes:A. extraction 1;B. defluorination 1;C. defluorination 2;D. extraction 2;E. back extraction 1;F. pH adjustment;G. evaporation concentration to specific gravity 1.45~1.75, cooling crystallization, solid-liquid separation obtains phosphate salt.Compared with traditional chemical method, the method of the present application has the advantages of low energy consumption, small pollution, high resource utilization rate.The present application shortens the production process of phosphate, and sample diversity is good.The phosphate obtained by the present application has high purity.The method of the present application does not produce raffinate acid.In addition, the method of the present application can also extract other valuable elements such as magnesium, aluminum etc.from phosphate rock, realize the comprehensive utilization of resources.

Description

Technical Field

[0001] This invention relates to a method for preparing phosphates, belonging to the field of chemical technology. Background Technology

[0002] Phosphates are important chemical raw materials, widely used in agricultural fertilizers, food processing, detergents, and metallurgy. There are many types of phosphates, including lithium dihydrogen phosphate, lithium hydrogen phosphate, lithium phosphate, sodium dihydrogen phosphate, disodium hydrogen phosphate, sodium phosphate, magnesium dihydrogen phosphate, magnesium hydrogen phosphate, magnesium phosphate, potassium dihydrogen phosphate, dipotassium hydrogen phosphate, potassium phosphate, aluminum dihydrogen phosphate, aluminum hydrogen phosphate, aluminum phosphate, calcium dihydrogen phosphate, calcium hydrogen phosphate, calcium phosphate, manganese dihydrogen phosphate, manganese hydrogen phosphate, manganese phosphate, ferric dihydrogen phosphate, ferric hydrogen phosphate, ferric phosphate, cobalt dihydrogen phosphate, cobalt hydrogen phosphate, cobalt phosphate, nickel dihydrogen phosphate, nickel hydrogen phosphate, and phosphorus. Nickel dihydrogen phosphate, copper dihydrogen phosphate, copper hydrogen phosphate, copper phosphate, zinc dihydrogen phosphate, zinc hydrogen phosphate, zinc phosphate, chromium dihydrogen phosphate, dichromium hydrogen phosphate, chromium phosphate, germanium dihydrogen phosphate, digermanium hydrogen phosphate, germanium phosphate, zirconium dihydrogen phosphate, dizirconium hydrogen phosphate, zirconium phosphate, niobium dihydrogen phosphate, diniobium hydrogen phosphate, niobium phosphate, tin dihydrogen phosphate, tin hydrogen phosphate, tin phosphate, antimony dihydrogen phosphate, antimony hydrogen phosphate, antimony phosphate, mercury dihydrogen phosphate, mercury hydrogen phosphate, barium dihydrogen phosphate, barium hydrogen phosphate, etc.

[0003] Phosphate production methods mainly include neutralization, extraction, ion exchange, metathesis, direct method, crystallization, electrolysis, solid-phase synthesis, solution method, thermal phosphoric acid production, wet phosphoric acid production, alkali source method, hydrothermal method, alkali-thermal method, solvothermal method, precipitation method, sol-gel method, etc., each with its own specific characteristics. However, existing phosphate production methods suffer from high energy consumption, significant pollution, and low resource utilization. Therefore, developing a new, environmentally friendly production method is of great significance for meeting the growing market demand while protecting the environment.

[0004] CN2023108576441 discloses a lithium dihydrogen phosphate (LDH) production process, including lithium solvent extraction, evaporation and concentration of lithium DHT aqueous solution, and lithium DHT refining. The lithium solvent extraction employs a β-diketone and neutral phosphorus extractant co-extraction technology, allowing the lithium-containing wastewater to be chloride-free. The lithium content in the wastewater can be as low as 0.1 g / L, and the lithium concentration in the raffinate can be reduced to below 10 ppm. The lithium-loaded organic phase is back-extracted using a phosphoric acid aqueous solution to obtain a lithium DHT aqueous solution. The lithium DHT aqueous solution is then concentrated by evaporation, cooled, crystallized, and separated into solid and liquid components to obtain crude lithium DHT. The crude lithium DHT is then refined to obtain battery-grade lithium DHT. This method can significantly reduce the production cost of lithium DHT while reducing waste generation, resulting in significant economic and social benefits. However, it has drawbacks, such as the need for saponification of the organic phase with 2.8 ml of 30% (w / w) sodium hydroxide during phosphoric acid extraction, increasing operating costs, and the presence of sodium hydroxide in the extractant after saponification, requiring subsequent sodium removal with hydrochloric acid.

[0005] CN2007100761125 discloses a method for synthesizing lithium dihydrogen phosphate, belonging to the field of inorganic material preparation technology. In a reactor, phosphoric acid and a lithium compound are reacted in an aqueous solution with stirring until the pH reaches 1.0–4.0, yielding an aqueous solution of lithium dihydrogen phosphate. This solution is then heated and concentrated to saturation. An extractant is added to the solution, and the mixture is stirred to extract and crystallize the lithium dihydrogen phosphate. The crystals are then filtered and dried to obtain lithium dihydrogen phosphate crystals. The extractant is a solvent with a solubility ratio of phosphate to lithium dihydrogen phosphate of 1–5. The extractant can be one or a mixture of methanol, ethanol, propanol, isopropanol, and glycerol. Adding the extractant to a saturated solution for extraction and crystallization overcomes the disadvantage of evaporative crystals being too fine, and crystallization and washing are completed in one step, simplifying the synthesis process and facilitating industrial production. However, the extraction is only used in the crystallization process to reduce the solubility of lithium dihydrogen phosphate and control the crystal form of the crystalline product.

[0006] CN2014100350541 discloses a method for preparing sodium dihydrogen phosphate using wet-process phosphoric acid. The method involves reacting carbamide urea—CO(NH2)2 with wet-process phosphoric acid to obtain an intermediate, which is then reacted with sodium hydroxide to produce the sodium dihydrogen phosphate product. This invention features a short process route, low energy consumption, stable product quality, low production cost, convenient operation, and safe production. The byproduct slurry can be completely recycled. The entire production process is environmentally friendly and pollution-free, with no waste gas, wastewater, or waste residue emissions. It responds to the policy calls for energy conservation, emission reduction, and clean production, overcoming the problems of complex processes, unstable product quality, high energy consumption, and environmental pollution inherent in existing technologies. The purity of the obtained sodium dihydrogen phosphate product is ≥98%. However, its purity is not high enough, and the sample diversity is poor.

[0007] CN2022117104426 discloses a production process for potassium dihydrogen phosphate, specifically including the following steps: Step 1: Add residual phosphoric acid and filtrate G to neutralization tank #1, add a pH adjuster, and after the reaction, send the slurry to a concentration tank to concentrate to a slurry density of 1.32-1.40. Filter the slurry to obtain filtrate A and filter residue B; Step 2: Send filtrate A from Step 1 to a crystallization tank to cool and precipitate crystals, separate to obtain crystals and filtrate C, and dry the crystals to obtain potassium dihydrogen phosphate product; Step 3: Send filter residue B and filter residue H from Step 1 to neutralization tank #2, add potassium solution I to react, add a pH adjuster, filter after the reaction to obtain filtrate D, wash the filter cake to obtain washing liquid E, and send the washed filter cake F to a compound fertilizer or superphosphate plant as an additive. This invention solves the technical problem in the prior art that the neutralization method uses industrial-grade purified phosphoric acid and potassium hydroxide to produce potassium dihydrogen phosphate, resulting in high raw material costs. However, it also has the problems of insufficient purity and poor sample diversity.

[0008] CN2021116639462 discloses a process for producing potassium dihydrogen phosphate by extraction, comprising the following steps: S100: adding raw materials phosphoric acid and potassium chloride to a mother liquor, raising the temperature of the mother liquor until the potassium chloride is completely dissolved; S200: adding an extractant to the mother liquor and stirring the reaction for a specified time, the extraction and mixing process using a microreactor; S300: allowing the mother liquor to stand and clarify until oil and water separation; S400: separating the oil phase using conventional methods; S500: evaporating the water in the mother liquor by heating; S600: cooling the mother liquor to crystallize, filtering out the crystals and drying to obtain the product. This invention has the advantages of low production cost and high product purity. Testing shows that the purity of potassium dihydrogen phosphate produced by this method is approximately 98.44%. However, it also has problems such as insufficient purity and poor sample diversity.

[0009] CN2021115132376 discloses an extractant for the production of potassium dihydrogen phosphate by solvent extraction, its preparation method, and its application. The extractant includes a main extractant and a diluent, wherein the main extractant is an asymmetric tertiary amine containing a carbonyl group. When using the extractant of this invention for the production of potassium dihydrogen phosphate by solvent extraction, it can improve the selectivity for hydrochloric acid, extracting more hydrochloric acid and less phosphoric acid; it improves the phase separation performance of the two phases, allowing for rapid phase separation without emulsification or two-phase fire zones, and features high phosphorus yield, low extractant loss, and good product quality. However, it also has problems such as insufficient purity and poor sample diversity.

[0010] CN2012100695247 discloses a method for producing potassium dihydrogen phosphate, comprising: 1. dissolving potassium chloride and phosphoric acid in water to prepare a product mother liquor; dissolving ammonia and ammonium chloride in water to prepare a by-product mother liquor; 2. extracting the product mother liquor with an extractant in a micro-mixer, and then separating the loaded extractant and raffinate; 3. cooling the raffinate to precipitate potassium dihydrogen phosphate crystals, and then performing solid-liquid separation, with the solid being the product and the liquid being the product mother liquor for reuse; 4. back-extracting the loaded extractant with the by-product mother liquor in another micro-mixer, and then performing phase separation to obtain a regenerated extractant and back-extract; 5. reusing the obtained regenerated extractant, and cooling the back-extract to precipitate ammonium chloride crystals, and then performing solid-liquid separation, with the solid being the by-product and the liquid being the by-product mother liquor for reuse. This invention avoids cross-contamination between the product mother liquor and the by-product mother liquor due to entrainment and selectivity issues through process control, and decouples the extraction process from the crystallization process through equipment enhancement to avoid the generation of fine crystals, thus producing high-quality potassium dihydrogen phosphate. However, it also has problems such as insufficient purity and poor sample diversity.

[0011] CN2017111163595 discloses a method for producing industrial-grade potassium dihydrogen phosphate using a wet-process phosphoric acid micro-reaction method, comprising the following steps: (1) phosphoric acid micro-extraction, (2) micro-reaction, (3) preparation of potassium dihydrogen phosphate, and (4) post-treatment. This method couples the extraction purification technology of wet-process phosphoric acid with the preparation technology of potassium dihydrogen phosphate using microchemical technology, thereby reducing equipment investment, improving the quality of potassium dihydrogen phosphate, and increasing the efficient utilization rate of phosphoric acid dihydrate. The organic phase of phosphoric acid purified by organic solvent extraction reacts directly with a potassium salt solution to obtain qualified industrial-grade potassium dihydrogen phosphate. However, it also suffers from insufficient purity and poor sample diversity.

[0012] CN2017105096269 discloses a method and apparatus for producing potassium dihydrogen phosphate from wet-process phosphoric acid. The process includes: wet-process phosphoric acid production; wet-process phosphoric acid pretreatment; neutralization; belt conveyor filtration; metathesis; filter press filtration; index adjustment; crystallization; separation and drying; the solid is fed into a fluidized bed dryer for drying to obtain industrial-grade potassium dihydrogen phosphate, and the mother liquor is recycled. This invention uses wet extraction of phosphoric acid to replace the thermal phosphoric acid process for preparing industrial-grade potassium dihydrogen phosphate, achieving the goal of reducing production costs and improving product market competitiveness. However, it also has problems such as insufficient purity and poor sample diversity.

[0013] CN2013100032277 discloses a method for producing industrial potassium dihydrogen phosphate, which involves sequentially adding phosphoric acid and an extractant to a potassium chloride solution, followed by extraction, cooling and crystallization, and then liquid-solid separation to obtain the product. This invention uses a mixture of trialkylamine and kerosene as the extractant. By controlling the pH value of the potassium chloride solution and various process parameters, the raw materials react fully, and the extractant has high extraction efficiency for HCl. After extraction, the reaction system is partially miscible and a large amount of potassium dihydrogen phosphate solid precipitates. Subsequent crystallization and liquid-solid separation are sufficient to obtain a high-quality potassium dihydrogen phosphate product. The entire process is mild, requires only one phase separation operation, has a simple flow, low energy consumption and production cost, and produces no waste. It is a clean production method with a simple flow, low production cost, and continuous production capability. However, it also has drawbacks such as insufficient purity and poor sample diversity.

[0014] CN2010105439939 discloses a method for producing industrial potassium dihydrogen phosphate and NPK compound fertilizer using wet-process purified phosphoric acid as raw material. The method uses wet-process purified phosphoric acid and potassium chloride as raw materials, and octanol and octanoic acid as extractants for extraction. The aqueous phase is cooled and crystallized to obtain potassium dihydrogen phosphate product. The organic phase is back-extracted with a saturated ammonium chloride solution under stirring. The mother liquor after back-extraction is cooled and crystallized to obtain the byproduct NPK compound fertilizer. This invention yields products with high purity. By using octanol and octanoic acid as extractants to modify the chemical composition and pH value of the aqueous phase, it can selectively separate specific components as well as perform broad-spectrum separation of multiple components. The extraction efficiency is high, the selectivity is good, and the extractants are easy to regenerate through back-extraction. However, it also has drawbacks such as insufficient purity and poor sample diversity.

[0015] CN011289716 discloses a method for producing potassium dihydrogen phosphate. This method employs a wet process, consisting of potassium bisulfate preparation, solution extraction, and potassium dihydrogen phosphate crystallization and separation. The operation follows the steps of sulfonation, hydrolysis, extraction, filtration, filtrate concentration, crystallization and separation, drying, and fractionation. The reaction conditions are mild, and there are no special requirements for equipment materials. The main advantage lies in the use of the segmented neutralization method for producing dicalcium phosphate, which has very low levels of iron, aluminum, fluorine, and sulfur impurities. Using this as raw material results in very little potassium fluorosilicate produced during production, leading to a high potassium yield and ensuring the quality of both potassium dihydrogen phosphate and phosphoric acid—something no other wet process can achieve. Experiments have shown that this process is advanced, reliable, and highly practical, representing the lowest-cost method for producing potassium dihydrogen phosphate, and the product quality fully meets national standards. However, it also has drawbacks, including insufficient purity and poor sample diversity.

[0016] CN202310186472X discloses a method for preparing potassium dihydrogen phosphate from wet-process phosphoric acid containing an extractant. In this method, KOH aqueous solution is used instead of water for back-extraction in the wet-process phosphoric acid containing the extractant to directly produce industrial-grade KH2PO4. The specific steps are as follows: wet-process phosphoric acid containing the extractant and potassium oxide solution are reacted at a temperature controlled at 80-90℃, and the stirrer is set at 50-70 rpm. The raw materials are pumped to a mixer for mixing, and then fed into a reaction vessel for reaction. After the reaction, the upper solvent is recovered, the pH of the lower reaction liquid is adjusted, and hydrogen peroxide is added for cooling and crystallization. After solid-liquid separation, the material is dried to obtain the final product. This invention mainly uses wet-process phosphoric acid containing an extractant to prepare potassium dihydrogen phosphate, and uses potassium hydroxide aqueous solution instead of water for back-extraction. After back-extraction, potassium dihydrogen phosphate is directly obtained by cooling and crystallization. This process is simple, avoids the introduction of back-extraction water, and eliminates the concentration step, which can significantly reduce the energy consumption and cost of potassium dihydrogen phosphate production. However, it also has problems such as insufficient purity and poor sample diversity.

[0017] CN2016109885601 discloses a method for producing potassium dihydrogen phosphate using wet-process phosphoric acid. The method involves sequentially adding phosphorus, slurry, and barium carbonate to wet-process phosphoric acid to reduce the concentration of sulfate anions in the phosphoric acid to 0.5 g / L. After settling, the supernatant is collected as refined desulfurized dilute phosphoric acid. This refined desulfurized dilute phosphoric acid is extracted with a composite extractant composed of octylamine and octanol in any proportion. After settling and phase separation, the organic phase coupled with phosphoric acid is the extract, and the remaining phosphoric acid is the raffinate. Potassium chloride is added to the extract to react, and after settling and phase separation, the lower layer is a crude potassium dihydrogen phosphate solution, and the upper layer is an organic phase rich in hydrochloric acid and containing small amounts of phosphoric acid and potassium chloride. After filtration of the crude potassium dihydrogen phosphate solution, potassium hydroxide is added to the filtrate to adjust the pH to approximately 6. After settling and impurities are removed, a refined potassium dihydrogen phosphate solution is obtained. Purified phosphoric acid is added to the refined potassium dihydrogen phosphate solution to adjust the pH to 4-5. The resulting solution is concentrated, crystallized, filtered, and dried to obtain the potassium dihydrogen phosphate product. However, it also has problems such as insufficient purity and poor sample diversity.

[0018] CN2022102624146 discloses a process for preparing feed-grade dicalcium phosphate, comprising the following steps: (1) using phosphate rock with a grade of 14-25%, purifying it through dihydrate extraction to obtain phosphate I, then neutralizing phosphate I with lime milk to generate dicalcium phosphate slurry, separating the phosphate slurry into solid and liquid to obtain dicalcium phosphate semi-finished product, drying it to obtain dicalcium phosphate with a moisture content of less than 4%; (2) using phosphate rock with a grade higher than 29%, extracting it through hemihydrate-dihydrate extraction, then concentrating, defluorinating, and dearsenicing to obtain phosphate II with a concentration greater than 50%; (3) reacting phosphate II obtained in step (2) with phosphate obtained in step (1) to generate dicalcium phosphate semi-finished product, aging it, drying it, and then sieving it to obtain dicalcium phosphate finished product. This invention can make extensive use of low-grade phosphate rock that is difficult to apply in traditional processes, reducing the proportion of high-grade phosphate rock used, and the unit product is equivalent to saving about 45% of coal consumption. However, it also has the problems of insufficient purity and poor sample diversity.

[0019] CN2021116818104 discloses a process for producing feed-grade calcium dihydrogen phosphate, comprising the steps of raw material preparation, solid-liquid reaction, filtration, and drying. The advantages of this invention compared to existing technologies are: this method can extract dilute phosphoric acid from phosphate rock, then generate calcium dihydrogen phosphate through defluorination and liquid-phase calcium-containing compounds, and combine this with thermal phosphoric acid and wet defluorination to concentrate phosphoric acid for solid-liquid reaction, which can significantly reduce the cost of calcium dihydrogen phosphate production. However, it also suffers from problems such as insufficient purity and poor sample diversity. Summary of the Invention

[0020] The purpose of this invention is to provide a method for preparing phosphates.

[0021] To achieve the objective of this invention, the method includes:

[0022] A. Extraction 1: Wet phosphoric acid is extracted using extractant 1. After extraction, the phases are separated to obtain oil phase 1 and aqueous phase 1. The extractant 1 is an extractant for acidic metal ions.

[0023] B. Defluorination 1: The aqueous phase 1 is concentrated and defluorinated. The concentration and defluorination temperature is 40-100℃, the pressure is 10-90kPa, the concentration endpoint has a phosphoric acid P2O5 content of 30-50%, and a specific gravity of 1.4-2.2.

[0024] C. Defluorination 2: The wet-process phosphoric acid after defluorination 1 is chemically defluorinated a second time using a defluorinating agent;

[0025] D. Extraction 2: Extract the wet phosphoric acid after the secondary defluorination in step C with extractant 2. After extraction, separate the phases to obtain oil phase 2 and aqueous phase 2.

[0026] E. Back-extraction 1: Back-extracting the oil phase 2 using back-extracting agent 2 to separate the phases into oil phase 3 and aqueous phase 3. The back-extracting agent 2 is a solution of at least one of the soluble sulfate, hydrogen sulfate, hydrochloride, and nitrate of metal M. M is Li, Na, Mg, K, Al, Mn, Fe, Co, Ni, Cu, Zn, Cr, Ge, Zr, Nr, Sn, Sb, Hg, or Ba, and the molar ratio of M ions to phosphate carried in oil phase 2 is 2:1 to 1:2.

[0027] F. pH adjustment: The pH of the aqueous phase 3 obtained in step E is adjusted to 2.0–11.0 using the carbonate, bicarbonate, and hydroxide of metal M;

[0028] G. Evaporate and concentrate to a specific gravity of 1.45-1.75, cool and crystallize, and separate the solid and liquid to obtain phosphate, wherein the purity of the phosphate is above 97%.

[0029] Step F can be adjusted according to different products. For example, dihydrogen phosphate has a low pH value; while hydrogen phosphate or phosphate requires the addition of a neutralizing agent to neutralize the free hydrogen ions.

[0030] In one specific embodiment, the volume ratio of extractant 1 to wet phosphoric acid in step A is 1:3 to 3:1, the extraction temperature is 10 to 80°C, the number of extraction stages in step A is preferably 1 to 6, and the extraction time is preferably 5 to 60 min.

[0031] The preferred temperature for phase separation in step A is 10–80°C, and the separation time is 0.5–10 min.

[0032] In one specific embodiment, the extractant 1 comprises at least one of phosphate esters, phosphonates, hypophosphite, naphthalenesulfonic acid, cycloalkanoic acid, and tertiary carbonate extractants; preferably, the extractant 1 is at least one of di(2-ethylhexyl) phosphate, 2-ethylhexyl phosphate, di(2,4,4-trimethylpentyl)hypophosphite, 2-ethylhexylphosphonic acid mono(2-ethylhexyl) ester, N,NN-n-octylaminedimethylenephenylphosphonic acid, N,NN-n-hexylaminedimethylenephenylphosphonic acid, toluenesulfonic acid, p-toluenesulfonic acid, benzenesulfonic acid, xylenesulfonic acid, dodecylbenzenesulfonic acid, dinonylnaphthalenesulfonic acid, cinnamic acid, fatty acids, lauric acid, and cycloalkanoic acid;

[0033] The extractant 2 includes at least one of lipids, sulfoxides, phosphate esters, ketones, alcohols, and organic amines. Preferably, the extractant 2 is at least one of ethyl acetate, amyl acetate, butyl acetate, dioctyl sulfoxide, diphenyl sulfoxide, hydrocarbon sulfoxide, di(2-ethylhexyl) hexyl phosphate, dioctyl phosphate, tributyl phosphate, methyl isobutyl ketone, cyclohexanone, isoamyl alcohol, sec-octanol, substituted primary alcohols, trialkylmethylamine, didecylamine, trioctylamine, and trinonylamine.

[0034] In one specific embodiment, the method further includes the treatment of the oil phase 1: the oil phase 1 is back-extracted using a back-extraction agent 1, and the oil phase 4 and the aqueous phase 4 are obtained after back-extraction and phase separation.

[0035] The oil phase 4 is returned and recycled as extractant 1;

[0036] The aqueous phase 4 is cooled and crystallized to obtain aluminum ammonium sulfate, magnesium ammonium sulfate solids and crystallization mother liquor 1. The crystallization mother liquor 1 is returned to the aqueous phase 4 for recycling.

[0037] The back-extraction agent 1 is at least one of sulfuric acid and soluble sulfate solution;

[0038] Preferably, the volume ratio of the back-extraction agent 1 to the extractant is 1:3 to 3:1, the number of back-extraction stages is 1 to 6, the back-extraction temperature is 10 to 80°C, the back-extraction time is 5 to 60 min, and the back-extraction agent 1 is preferably at least one of sulfuric acid and ammonium sulfate solution, with a concentration of 1 to 50%.

[0039] The preferred phase separation temperature for obtaining oil phase 4 and water phase 4 is 10–80°C, and the phase separation time is 0.5–10 min.

[0040] The preferred cooling endpoint temperature for the aqueous phase 4 cooling crystallization is 10–60°C.

[0041] In one specific embodiment, the defluorinating agent is at least one of sodium carbonate and ammonium bicarbonate, calcium carbonate, sodium phosphate, sodium hydroxide, calcium hydroxide, calcium oxide, and sodium bicarbonate, and the secondary defluorination temperature is preferably 10-80°C.

[0042] In one specific embodiment, the extraction temperature in step D is 10–80°C, the volume ratio of extractant 2 to wet phosphoric acid after secondary defluorination is 3:1–1:3, and the number of extraction stages is 1–6.

[0043] The phase separation temperature in step D is 10–80°C, and the phase separation time is 0.5–10 min.

[0044] Preferably, the aqueous phase 2 is returned to the extraction 2 in step D.

[0045] In one specific embodiment, the back-extraction temperature in step E is 10–80°C, the volume ratio of back-extraction agent 2 to oil phase 2 is 3:1 to 1:3, the phase separation temperature in step E is preferably 10–80°C, and the phase separation time is 0.5–10 min.

[0046] In one specific embodiment, the method further includes the treatment of the oil phase 3: back-extracting the oil phase 3 with ammonia or water to separate the phases into an oil phase 5 and an aqueous phase 5;

[0047] The aqueous phase 5 was evaporated and concentrated, cooled and crystallized, and then separated into solid and liquid phases to obtain solid ammonium salt and liquid crystallization mother liquor.

[0048] The oil phase 5 is returned to step D and recycled as extractant 2.

[0049] The preferred back-extraction temperature of oil phase 3 is 10-80℃, the volume ratio of ammonia water to organic phase is 3:1-1:3, and the molar ratio of NH3 in ammonia water to acid carried in oil phase 3 is 3:1-1:3.

[0050] Preferably, the aqueous phase 5 is evaporated and concentrated to a specific gravity of 1.1 to 1.7, the concentration temperature is 10 to 100°C, the pressure is 10 to 80 kPa, and the cooling point for crystallization is 10 to 60°C.

[0051] The liquid crystallization mother liquor obtained by evaporation, concentration, cooling, crystallization, and solid-liquid separation in aqueous phase 5 is returned to aqueous phase 5 for recycling.

[0052] Ammonia is preferably used to back-extract oil phase 3.

[0053] The solid-liquid separation described in this invention can be a common solid-liquid separation method, such as filtration or centrifugation.

[0054] In one specific embodiment, when the M ion in the wet-process phosphoric acid is in a low valence state and a phosphate with a high valence state of M ion needs to be prepared, the method further includes oxidizing the M ion before neutralization; the oxidant is at least one selected from oxygen, ozone, hydrogen peroxide, sodium peroxide, barium peroxide, chlorine, potassium permanganate, potassium dichromate, chlorine, sodium hypochlorite, perchloric acid, and nitric acid, and the oxidation temperature is preferably 20-80°C;

[0055] The preferred temperature for neutralization in step F is 10–100°C.

[0056] For example, wet-process phosphoric acid contains divalent ferrous ions, but the product requires trivalent ferric ions, which necessitates oxidation.

[0057] In one specific embodiment, the evaporation and concentration in step G is carried out at a temperature of 10–90°C and a pressure of 10–50 kPa; the crystallization temperature in step G is preferably 10–60°C; and the mother liquor after solid-liquid separation in step G is preferably returned to the evaporation and concentration in step G.

[0058] Beneficial effects:

[0059] To address the problems of low purity, long production processes, high energy consumption, poor sample diversity, and low utilization of other valuable elements in phosphate rock in existing phosphate products, this invention proposes a novel method for phosphate production. This method is based on extraction and back-extraction technology. Wet-process phosphoric acid is purified using extractant 1, followed by a secondary extraction using extractant 2. Different salts are used to back-extract extractant 2 to obtain a phosphate MH2PO4 solution. Different phosphates, MHPO4 and MPO4, are then obtained by adjusting the pH. Compared with traditional chemical methods, this invention offers advantages such as low energy consumption, low pollution, and high resource utilization.

[0060] The method of this invention can obtain different phosphates by selecting different back-extraction agents, which greatly shortens the phosphate production process and provides good sample diversity.

[0061] The phosphate obtained by this invention has high purity.

[0062] The method of the present invention does not produce residual acid.

[0063] Furthermore, the method of the present invention can also extract other valuable elements such as magnesium and aluminum from phosphate rock at the same time, thereby achieving comprehensive utilization of resources.

[0064] The implementation of this invention will provide a new and environmentally friendly method for the production of phosphates, which is expected to promote the development of related industries, meet market demand, and contribute to environmental protection. Attached Figure Description

[0065] Figure 1 This is a process flow diagram of a specific embodiment of the present invention. Detailed Implementation

[0066] To achieve the objective of this invention, the method includes:

[0067] A. Extraction 1: Wet phosphoric acid is extracted using extractant 1. After extraction, the phases are separated to obtain oil phase 1 and aqueous phase 1. The extractant 1 is an extractant for acidic metal ions.

[0068] B. Defluorination 1: The aqueous phase 1 is concentrated and defluorinated. The concentration and defluorination temperature is 40-100℃, the pressure is 10-90kPa, the concentration endpoint has a phosphoric acid P2O5 content of 30-50%, and a specific gravity of 1.4-2.2.

[0069] C. Defluorination 2: The wet-process phosphoric acid after defluorination 1 is chemically defluorinated a second time using a defluorinating agent;

[0070] D. Extraction 2: Extract the wet phosphoric acid after the secondary defluorination in step C with extractant 2. After extraction, separate the phases to obtain oil phase 2 and aqueous phase 2.

[0071] E. Back-extraction 1: Back-extracting the oil phase 2 using back-extracting agent 2 to separate the phases into oil phase 3 and aqueous phase 3. The back-extracting agent 2 is a solution of at least one of the soluble sulfate, hydrogen sulfate, hydrochloride, and nitrate of metal M. M is Li, Na, Mg, K, Al, Mn, Fe, Co, Ni, Cu, Zn, Cr, Ge, Zr, Nr, Sn, Sb, Hg, or Ba, and the molar ratio of M ions to phosphate carried in oil phase 2 is 2:1 to 1:2.

[0072] F. pH adjustment: The pH of the aqueous phase 3 obtained in step E is adjusted to 2.0–11.0 using the carbonate, bicarbonate, and hydroxide of metal M;

[0073] G. Evaporate and concentrate to a specific gravity of 1.45-1.75, cool and crystallize, and separate the solid and liquid to obtain phosphate, wherein the purity of the phosphate is above 97%.

[0074] In one specific embodiment, the volume ratio of extractant 1 to wet phosphoric acid in step A is 1:3 to 3:1, the extraction temperature is 10 to 80°C, the number of extraction stages in step A is preferably 1 to 6, and the extraction time is preferably 5 to 60 min.

[0075] The preferred temperature for phase separation in step A is 10–80°C, and the separation time is 0.5–10 min.

[0076] In one specific embodiment, the extractant 1 comprises at least one of phosphate esters, phosphonates, hypophosphite, naphthalenesulfonic acid, cycloalkanoic acid, and tertiary carbonate extractants; preferably, the extractant 1 is at least one of di(2-ethylhexyl) phosphate, 2-ethylhexyl phosphate, di(2,4,4-trimethylpentyl)hypophosphite, 2-ethylhexylphosphonic acid mono(2-ethylhexyl) ester, N,NN-n-octylaminedimethylenephenylphosphonic acid, N,NN-n-hexylaminedimethylenephenylphosphonic acid, toluenesulfonic acid, p-toluenesulfonic acid, benzenesulfonic acid, xylenesulfonic acid, dodecylbenzenesulfonic acid, dinonylnaphthalenesulfonic acid, cinnamic acid, fatty acids, lauric acid, and cycloalkanoic acid;

[0077] The extractant 2 includes at least one of lipids, sulfoxides, phosphate esters, ketones, alcohols, and organic amines; preferably, the extractant 2 is at least one of ethyl acetate, amyl acetate, butyl acetate, dioctyl sulfoxide, diphenyl sulfoxide, hydrocarbon sulfoxide, di(2-ethylhexyl) hexyl phosphate, dioctyl phosphate, tributyl phosphate, methyl isobutyl ketone, cyclohexanone, isoamyl alcohol, sec-octanol, substituted primary alcohols, trialkylmethylamine, didecylamine, trioctylamine, and trinonylamine.

[0078] In one specific embodiment, the method further includes the treatment of the oil phase 1: the oil phase 1 is back-extracted using a back-extraction agent 1, and the oil phase 4 and the aqueous phase 4 are obtained after back-extraction and phase separation.

[0079] The oil phase 4 is returned and recycled as extractant 1;

[0080] The aqueous phase 4 is cooled and crystallized to obtain aluminum ammonium sulfate, magnesium ammonium sulfate solids and crystallization mother liquor 1. The crystallization mother liquor 1 is returned to the aqueous phase 4 for recycling.

[0081] The back-extraction agent 1 is at least one of sulfuric acid and soluble sulfate solution;

[0082] Preferably, the volume ratio of the back-extraction agent 1 to the extractant is 1:3 to 3:1, the number of back-extraction stages is 1 to 6, the back-extraction temperature is 10 to 80°C, the back-extraction time is 5 to 60 min, and the back-extraction agent 1 is preferably at least one of sulfuric acid and ammonium sulfate solution, with a concentration of 1 to 50%.

[0083] The preferred phase separation temperature for obtaining oil phase 4 and water phase 4 is 10–80°C, and the phase separation time is 0.5–10 min.

[0084] The preferred cooling endpoint temperature for the aqueous phase 4 cooling crystallization is 10–60°C.

[0085] In one specific embodiment, the defluorinating agent is at least one of sodium carbonate and ammonium bicarbonate, calcium carbonate, sodium phosphate, sodium hydroxide, calcium hydroxide, calcium oxide, and sodium bicarbonate, and the secondary defluorination temperature is preferably 10-80°C.

[0086] In one specific embodiment, the extraction temperature in step D is 10–80°C, the volume ratio of extractant 2 to wet phosphoric acid after secondary defluorination is 3:1–1:3, and the number of extraction stages is 1–6.

[0087] The phase separation temperature in step D is 10–80°C, and the phase separation time is 0.5–10 min.

[0088] Preferably, the aqueous phase 2 is returned to the extraction 2 in step D.

[0089] In one specific embodiment, the back-extraction temperature in step E is 10–80°C, the volume ratio of back-extraction agent 2 to oil phase 2 is 3:1 to 1:3, the phase separation temperature in step E is preferably 10–80°C, and the phase separation time is 0.5–10 min.

[0090] In one specific embodiment, the method further includes the treatment of the oil phase 3: back-extracting the oil phase 3 with ammonia or water to separate the phases into an oil phase 5 and an aqueous phase 5;

[0091] The aqueous phase 5 was evaporated and concentrated, cooled and crystallized, and then separated into solid and liquid phases to obtain solid ammonium salt and liquid crystallization mother liquor.

[0092] The oil phase 5 is returned to step D and recycled as extractant 2.

[0093] The preferred back-extraction temperature of oil phase 3 is 10-80℃, the volume ratio of ammonia water to organic phase is 3:1-1:3, and the molar ratio of NH3 in ammonia water to acid carried in oil phase 3 is 3:1-1:3.

[0094] Preferably, the aqueous phase 5 is evaporated and concentrated to a specific gravity of 1.1 to 1.7, the concentration temperature is 10 to 100°C, the pressure is 10 to 80 kPa, and the cooling point for crystallization is 10 to 60°C.

[0095] The liquid crystallization mother liquor obtained by evaporation, concentration, cooling, crystallization, and solid-liquid separation in aqueous phase 5 is returned to aqueous phase 5 for recycling.

[0096] Ammonia is preferably used to back-extract oil phase 3.

[0097] The solid-liquid separation described in this invention can be a common solid-liquid separation method, such as filtration or centrifugation.

[0098] In one specific embodiment, the neutralization temperature in step F is preferably 10 to 100°C.

[0099] In one specific embodiment, the evaporation and concentration in step G is carried out at a temperature of 10–90°C and a pressure of 10–50 kPa; the crystallization temperature in step G is preferably 10–60°C; and the mother liquor after solid-liquid separation in step G is preferably returned to the evaporation and concentration in step H.

[0100] The specific embodiments of the present invention will be further described below with reference to examples, but the present invention is not limited to the scope of the embodiments described herein.

[0101] Example 1

[0102] Lithium dihydrogen phosphate: (1) Extraction 1: Metal cations in wet phosphoric acid were extracted using di(2-ethylhexyl) phosphate. The volume ratio of extractant to wet phosphoric acid was 3:1. The number of extraction stages was 3. The extraction temperature was 50℃ and the extraction time was 10min.

[0103] (2) Phase separation 1: The mixture of extractant and phosphoric acid after extraction is allowed to stand and separate to obtain oil phase 1 and aqueous phase 1. Oil phase 1 is organic extractant and a small amount of metal cations carried. The remaining aqueous phase 1 is phosphoric acid after metal cation removal. The oil phase is back-extracted and regenerated, and the aqueous phase enters the next stage. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0104] (3) Back-extraction 1: Ammonium sulfate solution is used as the back-extraction agent to back-extract oil phase 1. The volume ratio of back-extraction liquid to extractant is 1:1, the number of back-extraction stages is 1, the back-extraction temperature is 50℃, the back-extraction time is 60min, and the concentration of back-extraction agent is 30%.

[0105] (4) Phase separation 2: The extractant and the back-extractant solution after back-extraction are allowed to stand and separate to obtain oil phase 2 and water phase 2. Oil phase 2 is organic extractant, and the remaining water phase 2 is back-extractant and eluted metal ions. Oil phase 2 is returned to extraction 1 process, and water phase enters the next cooling crystallization 1 process. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0106] (5) Cooling crystallization 1: The liquid obtained after back-extraction is cooled and crystallized, with the final cooling temperature being 20℃.

[0107] (6) Filtration 1: The solid-liquid mixture obtained by cooling crystallization is filtered. The solid is ammonium aluminum sulfate and ammonium magnesium sulfate, and the liquid is crystallization mother liquor 1, which is returned to the cooling crystallization 1 section.

[0108] (7) Defluorination 1: The aqueous phase 1 obtained in (2) is concentrated and defluorinated. The concentration and defluorination temperature is 80℃, the pressure is 30kPa, the concentration endpoint is 40% phosphoric acid P2O5 content and specific gravity is 1.5.

[0109] (8) Defluorination 2: The concentrated defluorinated wet phosphoric acid is subjected to secondary chemical defluorination using a defluorinating agent, ammonium bicarbonate, at a temperature of 60°C.

[0110] (9) Extraction 2: Trialkylmethylamine was used as the extractant to extract the wet phosphoric acid after secondary defluorination in Defluorination 2. The extraction temperature was 60℃, the volume ratio of the composite extractant to the wet phosphoric acid after secondary defluorination was 1:3, and the number of extraction stages was 3.

[0111] (10) Phase separation 3: The mixed solution of composite extractant and wet phosphoric acid is allowed to stand and separate. The organic oil phase coupled with phosphoric acid is the extract, and the remaining aqueous phase is water and a small amount of phosphoric acid. The aqueous phase is returned to extraction 2. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0112] (11) Back-extraction 2: The oil phase 3 obtained by phase separation 3 is back-extracted with lithium sulfate. The back-extraction temperature is 60℃, the volume ratio of back-extractant to organic phase is 1:1, and the molar ratio of lithium ions to phosphoric acid carried in the extractant is 1:1.05.

[0113] (12) Phase separation 4: The mixture of extract and back-extraction solution is allowed to stand and separate to obtain oil phase 4 and aqueous phase 4. The main component of aqueous phase 4 is H2PO4. - Li + H2O and very small amounts of SO4 2-The main component of oil phase 4 is H2SO4 carried by a composite extractant. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0114] (13) Back-extraction 3: Ammonia water is added to the oil phase 4 obtained by phase separation 4 for secondary back-extraction. The secondary back-extraction temperature is 50℃, the volume ratio of back-extraction agent to organic phase is 1:1, and the molar ratio of back-extraction agent NH3 to acid carried in the extractant is 1.1:1.

[0115] (14) Phase separation 5: The substance obtained from back-extraction 3 is phase separated. The oil phase of the composite extractant after back-extraction is returned to the extraction 2 process, and the aqueous phase is used for subsequent operations.

[0116] (15) Evaporation and concentration 1: The aqueous ammonium sulfate solution obtained from phase separation 5 is evaporated and concentrated to a specific gravity of 1.35, the concentration temperature is 80℃, and the pressure is 30kPa;

[0117] (16) Cooling crystallization 2: The concentrated liquid is cooled and crystallized, with the final cooling temperature being 40℃;

[0118] (17) Filtration 2: The product obtained by cooling crystallization is filtered to obtain solid ammonium salt (NH4)2SO4, and the liquid crystallization mother liquor is returned to (15) Evaporation and Concentration 1 section.

[0119] (18) Evaporation and concentration 2: The aqueous phase 4 obtained in (12) is evaporated and concentrated to a specific gravity of 1.45, the concentration temperature is 80℃, and the pressure is 20kPa;

[0120] (19) Cooling crystallization 3: The concentrated liquid is cooled and crystallized to obtain relatively pure lithium dihydrogen phosphate at a crystallization temperature of 30℃.

[0121] Example 2

[0122] Sodium dihydrogen phosphate

[0123] (1) Extraction 1: 2-ethylhexyl phosphate 2-ethylhexyl ester was used to extract metal cations from wet phosphoric acid. The volume ratio of extractant to wet phosphoric acid was 3:1. The number of extraction stages was 3. The extraction temperature was 50℃ and the extraction time was 10min.

[0124] (2) Phase separation 1: The mixture of extractant and phosphoric acid after extraction is allowed to stand and separate to obtain oil phase 1 and aqueous phase 1. Oil phase 1 is organic extractant and a small amount of metal cations carried. The remaining aqueous phase 1 is phosphoric acid after metal cation removal. The oil phase is back-extracted and regenerated, and the aqueous phase enters the next stage. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0125] (3) Back-extraction 1: Ammonium sulfate is used as the back-extraction agent to back-extract oil phase 1. The volume ratio of back-extraction liquid to extractant is 1:1, the number of back-extraction stages is 1, the back-extraction temperature is 50℃, the back-extraction time is 60min, and the concentration of back-extraction agent is 30%.

[0126] (4) Phase separation 2: The extractant and the back-extractant solution after back-extraction are allowed to stand and separate to obtain oil phase 2 and water phase 2. Oil phase 2 is organic extractant, and the remaining water phase 2 is back-extractant and eluted metal ions. Oil phase 2 is returned to extraction 1 process, and water phase enters the next cooling crystallization 1 process. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0127] (5) Cooling crystallization 1: The liquid obtained after back-extraction is cooled and crystallized, with the final cooling temperature being 20℃.

[0128] (6) Filtration 1: The solid-liquid mixture obtained by cooling crystallization is filtered. The solid is ammonium aluminum sulfate and ammonium magnesium sulfate, and the liquid is crystallization mother liquor 1, which is returned to the cooling crystallization 1 section.

[0129] (7) Defluorination 1: The aqueous phase 1 obtained in (2) is concentrated and defluorinated. The concentration and defluorination temperature is 80℃, the pressure is 30kPa, the concentration endpoint is 40% phosphoric acid P2O5 content and specific gravity is 1.5.

[0130] (8) Defluorination 2: The concentrated defluorinated wet phosphoric acid is subjected to secondary chemical defluorination using a defluorinating agent, ammonium bicarbonate, at a temperature of 60°C.

[0131] (9) Extraction 2: Didecylamine was used as the extractant to extract the wet phosphoric acid after secondary defluorination in Defluorination 2. The extraction temperature was 60℃, the volume ratio of the composite extractant to the wet phosphoric acid after secondary defluorination was 1:3, and the number of extraction stages was 3.

[0132] (10) Phase separation 3: The mixed solution of composite extractant and wet phosphoric acid is allowed to stand and separate. The organic oil phase coupled with phosphoric acid is the extract, and the remaining aqueous phase is water and a small amount of phosphoric acid. Aqueous phase 3 is returned to wet phosphoric acid extraction 2 for the next extraction. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0133] (11) Back-extraction 2: Sodium sulfate is used to back-extract the oil phase 3 obtained from phase separation 3. The back-extraction temperature is 60℃, the volume ratio of back-extractant to organic phase is 1:1, and the molar ratio of sodium ions to phosphoric acid carried in the extractant is 1:1.05.

[0134] (12) Phase separation 4: The mixture of extract and back-extraction solution is allowed to stand and separate to obtain oil phase 4 and aqueous phase 4. The main component of aqueous phase 4 is H2PO4. - Na + H2O and very small amounts of SO4 2-The main component of oil phase 4 is H2SO4 carried by a composite extractant. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0135] (13) Back-extraction 3: Ammonia water is added to the oil phase 4 obtained by phase separation 4 for secondary back-extraction. The secondary back-extraction temperature is 50℃, the volume ratio of back-extraction agent to organic phase is 1:1, and the molar ratio of back-extraction agent NH3 to acid carried in the extractant is 1.1:1.

[0136] (14) Phase separation 5: The substance obtained from back-extraction 3 is phase separated. The oil phase of the composite extractant after back-extraction is returned to the extraction 2 process, and the aqueous phase is used for subsequent operations.

[0137] (15) Evaporation and concentration 1: The aqueous ammonium sulfate solution obtained from phase separation 5 is evaporated and concentrated to a specific gravity of 1.35, the concentration temperature is 80℃, and the pressure is 30kPa;

[0138] (16) Cooling crystallization 2: The concentrated liquid is cooled and crystallized, with the final cooling temperature being 40℃;

[0139] (17) Filtration 2: The product obtained by cooling crystallization is filtered to obtain solid ammonium salt (NH4)2SO4, and the liquid crystallization mother liquor is returned to (15) Evaporation and Concentration 1 section.

[0140] (18) Evaporation and concentration 2: The concentrated mother liquor obtained in (12) is evaporated and concentrated to a specific gravity of 1.50, a concentration temperature of 80℃, and a pressure of 20kPa;

[0141] (19) Cooling crystallization 3: Cool the concentrated liquid to obtain relatively pure sodium dihydrogen phosphate. The crystallization temperature is 30℃.

[0142] Example 3

[0143] Magnesium dihydrogen phosphate

[0144] (1) Extraction 1: Metal cations in wet phosphoric acid were extracted using di(2,4,4-trimethylpentyl)phosphoric acid. The volume ratio of extractant to wet phosphoric acid was 3:1, the number of extraction stages was 3, the extraction temperature was 50℃, and the extraction time was 10min.

[0145] (2) Phase separation 1: The mixture of extractant and phosphoric acid after extraction is allowed to stand and separate to obtain oil phase 1 and aqueous phase 1. Oil phase 1 is organic extractant and a small amount of metal cations carried. The remaining aqueous phase 1 is phosphoric acid after metal cation removal. The oil phase is back-extracted and regenerated, and the aqueous phase enters the next stage. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0146] (3) Back-extraction 1: Ammonium sulfate is used as the back-extraction agent to back-extract oil phase 1. The volume ratio of back-extraction liquid to extractant is 1:1, the number of back-extraction stages is 1, the back-extraction temperature is 50℃, the back-extraction time is 60min, and the concentration of back-extraction agent is 30%.

[0147] (4) Phase separation 2: The extractant and the back-extractant solution after back-extraction are allowed to stand and separate to obtain oil phase 2 and water phase 2. Oil phase 2 is organic extractant, and the remaining water phase 2 is back-extractant and eluted metal ions. Oil phase 2 is returned to extraction 1 process, and water phase enters the next cooling crystallization 1 process. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0148] (5) Cooling crystallization 1: The liquid obtained after back-extraction is cooled and crystallized, with the final cooling temperature being 20℃.

[0149] (6) Filtration 1: The solid-liquid mixture obtained by cooling crystallization is filtered. The solid is ammonium aluminum sulfate and ammonium magnesium sulfate, and the liquid is crystallization mother liquor 1, which is returned to the cooling crystallization 1 section.

[0150] (7) Defluorination 1: The aqueous phase 1 obtained in (2) is concentrated and defluorinated. The concentration and defluorination temperature is 80℃, the pressure is 30kPa, the concentration endpoint is 40% phosphoric acid P2O5 content and specific gravity is 1.5.

[0151] (8) Defluorination 2: The concentrated defluorinated wet phosphoric acid is subjected to secondary chemical defluorination using a defluorinating agent, ammonium bicarbonate, at a temperature of 60°C.

[0152] (9) Extraction 2: Trioctylamine was used as the extractant to extract wet phosphoric acid after secondary defluorination in defluorination 2. The extraction temperature was 60℃, the volume ratio of the composite extractant to the wet phosphoric acid after secondary defluorination was 1:3, and the number of extraction stages was 3.

[0153] (10) Phase separation 3: The mixed solution of composite extractant and wet phosphoric acid is allowed to stand and separate. The organic oil phase coupled with phosphoric acid is the extract, and the remaining aqueous phase is water and a small amount of phosphoric acid. Aqueous phase 3 is returned to wet phosphoric acid extraction 2 for the next extraction. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0154] (11) Back-extraction 2: Magnesium sulfate is used to back-extract the oil phase 3 obtained from phase separation 3. The back-extraction temperature is 60℃, the volume ratio of the back-extractant to the organic phase is 1:1, and the molar ratio of magnesium ions to phosphoric acid carried in the extractant is 1:2.1.

[0155] (12) Phase separation 4: The mixture of extract and back-extraction solution is allowed to stand and separate to obtain oil phase 4 and aqueous phase 4. The main component of aqueous phase 4 is H2PO4. - Mg 2+ H2O and very small amounts of SO4 2-The main component of oil phase 4 is H2SO4 carried by a composite extractant. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0156] (13) Back-extraction 3: Ammonia water is added to the oil phase 4 obtained by phase separation 4 for secondary back-extraction. The secondary back-extraction temperature is 50℃, the volume ratio of back-extraction agent to organic phase is 1:1, and the molar ratio of back-extraction agent NH3 to acid carried in the extractant is 1.1:1.

[0157] (14) Phase separation 5: The substance obtained from back-extraction 3 is phase separated. The oil phase of the composite extractant after back-extraction is returned to the extraction 2 process, and the aqueous phase is used for subsequent operations.

[0158] (15) Evaporation and concentration 1: The aqueous ammonium sulfate solution obtained from phase separation 5 is evaporated and concentrated to a specific gravity of 1.35, the concentration temperature is 80℃, and the pressure is 30kPa;

[0159] (16) Cooling crystallization 2: The concentrated liquid is cooled and crystallized, with the final cooling temperature being 40℃;

[0160] (17) Filtration 2: The product obtained by cooling crystallization is filtered to obtain solid ammonium salt (NH4)2SO4, and the liquid crystallization mother liquor is returned to (15) Evaporation and Concentration 1 section.

[0161] (18) Evaporation and concentration 2: The concentrated mother liquor obtained in (12) is evaporated and concentrated to a specific gravity of 1.50, a concentration temperature of 80℃, and a pressure of 20kPa;

[0162] (19) Cooling crystallization 3: Cool the concentrated liquid to obtain relatively pure magnesium dihydrogen phosphate. The crystallization temperature is 30℃.

[0163] Example 4

[0164] Aluminum dihydrogen phosphate

[0165] (1) Extraction 1: Metal cations in wet phosphoric acid were extracted using 2-ethylhexylphosphonic acid mono(2-ethylhexyl) ester. The volume ratio of extractant to wet phosphoric acid was 3:1, the number of extraction stages was 3, the extraction temperature was 50℃, and the extraction time was 10min.

[0166] (2) Phase separation 1: The mixture of extractant and phosphoric acid after extraction is allowed to stand and separate to obtain oil phase 1 and aqueous phase 1. Oil phase 1 is organic extractant and a small amount of metal cations carried. The remaining aqueous phase 1 is phosphoric acid after metal cation removal. The oil phase is back-extracted and regenerated, and the aqueous phase enters the next stage. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0167] (3) Back-extraction 1: Ammonium sulfate is used as the back-extraction agent to back-extract oil phase 1. The volume ratio of back-extraction liquid to extractant is 1:1, the number of back-extraction stages is 1, the back-extraction temperature is 50℃, the back-extraction time is 60min, and the concentration of back-extraction agent is 30%.

[0168] (4) Phase separation 2: The extractant and the back-extractant solution after back-extraction are allowed to stand and separate to obtain oil phase 2 and water phase 2. Oil phase 2 is organic extractant, and the remaining water phase 2 is back-extractant and eluted metal ions. Oil phase 2 is returned to extraction 1 process, and water phase enters the next cooling crystallization 1 process. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0169] (5) Cooling crystallization 1: The liquid obtained after back-extraction is cooled and crystallized, with the final cooling temperature being 20℃.

[0170] (6) Filtration 1: The solid-liquid mixture obtained by cooling crystallization is filtered. The solid is ammonium aluminum sulfate and ammonium magnesium sulfate, and the liquid is crystallization mother liquor 1, which is returned to the cooling crystallization 1 section.

[0171] (7) Defluorination 1: The aqueous phase 1 obtained in (2) is concentrated and defluorinated. The concentration and defluorination temperature is 80℃, the pressure is 30kPa, the concentration endpoint is 40% phosphoric acid P2O5 content and specific gravity is 1.5.

[0172] (8) Defluorination 2: The concentrated defluorinated wet phosphoric acid is subjected to secondary chemical defluorination using a defluorinating agent, ammonium bicarbonate, at a temperature of 60°C.

[0173] (9) Extraction 2: Trinonylamine was used as the extractant to extract wet phosphoric acid after secondary defluorination in defluorination 2. The extraction temperature was 60℃, the volume ratio of the composite extractant to the wet phosphoric acid after secondary defluorination was 1:3, and the number of extraction stages was 3.

[0174] (10) Phase separation 3: The mixed solution of composite extractant and wet phosphoric acid is allowed to stand and separate. The organic oil phase coupled with phosphoric acid is the extract, and the remaining aqueous phase is water and a small amount of phosphoric acid. Aqueous phase 3 is returned to wet phosphoric acid extraction 2 for the next extraction. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0175] (11) Back-extraction 2: Aluminum sulfate is used to back-extract the oil phase 3 obtained from phase separation 3. The back-extraction temperature is 60℃, the volume ratio of back-extractant to organic phase is 1:1, and the molar ratio of aluminum ions to phosphoric acid carried in the extractant is 1:3.15.

[0176] (12) Phase separation 4: The mixture of extract and back-extraction solution is allowed to stand and separate to obtain oil phase 4 and aqueous phase 4. The main component of aqueous phase 4 is H2PO4. - Al 3+ H2O and very small amounts of SO4 2-The main component of oil phase 4 is H2SO4 carried by a composite extractant. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0177] (13) Back-extraction 3: Ammonia water is added to the oil phase 4 obtained by phase separation 4 for secondary back-extraction. The secondary back-extraction temperature is 50℃, the volume ratio of back-extraction agent to organic phase is 1:1, and the molar ratio of back-extraction agent NH3 to acid carried in the extractant is 1.1:1.

[0178] (14) Phase separation 5: The substance obtained from back-extraction 3 is phase separated. The oil phase of the composite extractant after back-extraction is returned to the extraction 2 process, and the aqueous phase is used for subsequent operations.

[0179] (15) Evaporation and concentration 1: The aqueous ammonium sulfate solution obtained from phase separation 5 is evaporated and concentrated to a specific gravity of 1.35, the concentration temperature is 80℃, and the pressure is 30kPa;

[0180] (16) Cooling crystallization 2: The concentrated liquid is cooled and crystallized, with the final cooling temperature being 40℃;

[0181] (17) Filtration 2: The product obtained by cooling crystallization is filtered to obtain solid ammonium salt (NH4)2SO4, and the liquid crystallization mother liquor is returned to (15) Evaporation and Concentration 1 section.

[0182] (18) Evaporation and concentration 2: The concentrated mother liquor obtained in (12) is evaporated and concentrated to a specific gravity of 1.50, a concentration temperature of 80℃, and a pressure of 20kPa;

[0183] (19) Cooling crystallization 3: Cooling crystallize the concentrated liquid to obtain relatively pure aluminum dihydrogen phosphate at a crystallization temperature of 30℃.

[0184] Example 5

[0185] Potassium dihydrogen phosphate:

[0186] (1) Extraction 1: N,NN-n-octylaminedimethylphenylphosphonic acid was used as a metal cation extractant to extract metal cations in wet phosphoric acid. The volume ratio of extractant to wet phosphoric acid was 3:1, the number of extraction stages was 3, the extraction temperature was 50℃, and the extraction time was 10min.

[0187] (2) Phase separation 1: The mixture of extractant and phosphoric acid after extraction is allowed to stand and separate to obtain oil phase 1 and aqueous phase 1. Oil phase 1 is organic extractant and a small amount of metal cations carried. The remaining aqueous phase 1 is phosphoric acid after metal cation removal. The oil phase is back-extracted and regenerated, and the aqueous phase enters the next stage. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0188] (3) Back-extraction 1: Ammonium sulfate is used to back-extract oil phase 1. The volume ratio of back-extraction liquid to extractant is 1:1, the number of back-extraction stages is 1, the back-extraction temperature is 50℃, the back-extraction time is 60min, and the concentration of back-extraction agent is 30%.

[0189] (4) Phase separation 2: The extractant and the back-extraction solution after back-extraction are allowed to stand and separate. The oil phase is the organic extractant, and the remaining aqueous phase is the back-extraction agent and the eluted metal ions. The oil phase is returned to the extraction 1 process, and the aqueous phase enters the next cooling crystallization 1 process. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0190] (5) Cooling crystallization 1: The liquid obtained after back-extraction is cooled and crystallized, with the final cooling temperature being 20℃.

[0191] (6) Filtration 1: The solid-liquid mixture obtained by cooling crystallization is filtered. The solid is ammonium aluminum sulfate and ammonium magnesium sulfate, and the liquid is the crystallization mother liquor, which is returned to the cooling crystallization section.

[0192] (7) Defluorination 1: The aqueous phase 1 obtained in (2) is concentrated and defluorinated. The concentration and defluorination temperature is 80℃, the pressure is 30kPa, the concentration endpoint is 40% phosphoric acid P2O5 content and specific gravity is 1.5.

[0193] (8) Defluorination 2: The concentrated defluorinated wet phosphoric acid is subjected to secondary chemical defluorination using a defluorinating agent, ammonium bicarbonate, at a temperature of 60°C.

[0194] (9) Extraction 2: Ethyl acetate was used as the extractant to extract wet phosphoric acid after the second defluorination of defluorination 2. The extraction temperature was 60℃, the volume ratio of the composite extractant to dilute phosphoric acid was 1:3, and the number of extraction stages was 3.

[0195] (10) Phase separation 3: The mixed solution of composite extractant and dilute phosphoric acid is allowed to stand and separate. The organic oil phase coupled with phosphoric acid is the extract, and the remaining aqueous phase is water and a small amount of phosphoric acid. After the phase separation 3 operation, the oil phase 3 and the aqueous phase 3 are obtained. The aqueous phase 3 is returned to the extraction stage 2 for the next extraction. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0196] (11) Back-extraction 2: KNO3 is used for back-extraction, the back-extraction temperature is 60℃, the volume ratio of back-extractant to organic phase is 1:1, and the molar ratio of potassium ions to phosphoric acid carried in the extractant is 1:1.05.

[0197] (12) Phase separation 4: The mixture of extract and back-extraction solution is allowed to stand and separate into phases. The oil phase is subjected to a second back-extraction. The main component of the aqueous phase is H2PO4. - K + H2O and trace amounts of NO3 -The main component in the oil phase is HNO3 carried by the composite extractant. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0198] (13) Back-extraction 3: Add ammonia water to the oil phase obtained by phase separation 4 for secondary back-extraction. The secondary back-extraction temperature is 50℃, the volume ratio of back-extraction agent to organic phase is 1:1, and the molar ratio of back-extraction agent NH3 to acid carried in the extractant is 1.1:1.

[0199] (14) Phase separation 5: The substance obtained from back-extraction 3 is phase separated. The oil phase of the composite extractant after back-extraction is returned to the extraction 2 process, and the aqueous phase is used for subsequent operations.

[0200] (15) Evaporation and concentration 1: The ammonium nitrate solution obtained from the second back-extraction is evaporated and concentrated to a specific gravity of 1.28, the concentration temperature is 80℃, and the pressure is 30kPa;

[0201] (16) Cooling crystallization 2: The concentrated liquid is cooled and crystallized, with the final cooling temperature being 40℃;

[0202] (17) Filtration 2: The product obtained by cooling crystallization is filtered to obtain solid ammonium salt NH4NO3, and the liquid crystallization mother liquor is returned to the concentration section.

[0203] (18) Evaporation and concentration 2: The aqueous phase 4 obtained by ((12) phase separation 4) is evaporated and concentrated to a specific gravity of 1.45, the concentration temperature is 80℃, and the pressure is 20kPa;

[0204] (19) Cooling crystallization 3: Cool the concentrated liquid to obtain relatively pure potassium dihydrogen phosphate. The crystallization temperature is 30℃.

[0205] Example 6

[0206] calcium dihydrogen phosphate

[0207] (1) Extraction 1: N,NN-n-hexylamine dimethylene phenylphosphonic acid was used to extract metal cations from wet phosphoric acid. The volume ratio of extractant to wet phosphoric acid was 3:1. The number of extraction stages was 3. The extraction temperature was 50℃ and the extraction time was 10min.

[0208] (2) Phase separation 1: The mixture of extractant and phosphoric acid after extraction is allowed to stand and separate to obtain oil phase 1 and aqueous phase 1. Oil phase 1 is organic extractant and a small amount of metal cations carried. The remaining aqueous phase 1 is phosphoric acid after metal cation removal. The oil phase is back-extracted and regenerated, and the aqueous phase enters the next stage. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0209] (3) Back-extraction 1: Ammonium sulfate is used as the back-extraction agent to back-extract oil phase 1. The volume ratio of back-extraction liquid to extractant is 1:1, the number of back-extraction stages is 1, the back-extraction temperature is 50℃, the back-extraction time is 60min, and the concentration of back-extraction agent is 30%.

[0210] (4) Phase separation 2: The extractant and the back-extractant solution after back-extraction are allowed to stand and separate to obtain oil phase 2 and water phase 2. Oil phase 2 is organic extractant, and the remaining water phase 2 is back-extractant and eluted metal ions. Oil phase 2 is returned to extraction 1 process, and water phase enters the next cooling crystallization 1 process. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0211] (5) Cooling crystallization 1: The liquid obtained after back-extraction is cooled and crystallized, with the final cooling temperature being 20℃.

[0212] (6) Filtration 1: The solid-liquid mixture obtained by cooling crystallization is filtered. The solid is ammonium aluminum sulfate and ammonium magnesium sulfate, and the liquid is crystallization mother liquor 1, which is returned to the cooling crystallization 1 section.

[0213] (7) Defluorination 1: The aqueous phase 1 obtained in (2) is concentrated and defluorinated. The concentration and defluorination temperature is 80℃, the pressure is 30kPa, the concentration endpoint is 40% phosphoric acid P2O5 content and specific gravity is 1.5.

[0214] (8) Defluorination 2: Use a defluorinating agent to perform secondary chemical defluorination on concentrated defluorinated wet phosphoric acid. The defluorinating agent is ammonium bicarbonate, and the defluorination temperature is 60℃.

[0215] (9) Extraction 2: Amyl acetate was used as the extractant to extract wet phosphoric acid after secondary defluorination in defluorination 2. The extraction temperature was 60℃, the volume ratio of the composite extractant to the wet phosphoric acid after secondary defluorination was 1:3, and the number of extraction stages was 3.

[0216] (10) Phase separation 3: The mixed solution of composite extractant and wet phosphoric acid is allowed to stand and separate. The organic oil phase coupled with phosphoric acid is the extract, and the remaining aqueous phase is water and a small amount of phosphoric acid. Aqueous phase 3 is returned to wet phosphoric acid extraction 2 for the next extraction. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0217] (11) Back-extraction 2: Calcium nitrate is used to back-extract the oil phase 3 obtained from phase separation 3. The back-extraction temperature is 60℃, the volume ratio of the back-extractant to the organic phase is 1:1, and the molar ratio of calcium ions to phosphoric acid carried in the extractant is 1:2.1.

[0218] (12) Phase separation 4: The mixture of extract and back-extraction solution is allowed to stand and separate to obtain oil phase 4 and aqueous phase 4. The main component of aqueous phase 4 is H2PO4. - Ca 2+ H2O and trace amounts of NO3 -The main component of the oil phase 4 is HNO3 carried by the composite extractant. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0219] (13) Back-extraction 3: Ammonia water is added to the oil phase 4 obtained by phase separation 4 for secondary back-extraction. The secondary back-extraction temperature is 50℃, the volume ratio of back-extraction agent to organic phase is 1:1, and the molar ratio of back-extraction agent NH3 to acid carried in the extractant is 1.1:1.

[0220] (14) Phase separation 5: The substance obtained from back-extraction 3 is phase separated. The oil phase of the composite extractant after back-extraction is returned to the extraction 2 process, and the aqueous phase is used for subsequent operations.

[0221] (15) Evaporation and concentration 1: The aqueous ammonium sulfate solution obtained from phase separation 5 is evaporated and concentrated to a specific gravity of 1.35, the concentration temperature is 80℃, and the pressure is 30kPa;

[0222] (16) Cooling crystallization 2: The concentrated liquid is cooled and crystallized, with the final cooling temperature being 40℃;

[0223] (17) Filtration 2: Filter the product obtained by cooling crystallization to obtain solid ammonium salt NH4NO3, and return the liquid crystallization mother liquor to (15) Evaporation and Concentration 1 section.

[0224] (18) Evaporation and concentration 2: The concentrated mother liquor obtained in (12) is evaporated and concentrated to a specific gravity of 1.50, a concentration temperature of 80℃, and a pressure of 20kPa;

[0225] (19) Cooling crystallization 3: Cool the concentrated liquid to obtain relatively pure calcium dihydrogen phosphate. The crystallization temperature is 30℃.

[0226] Example 7

[0227] Ferric dihydrogen phosphate

[0228] (1) Extraction 1: Metal cations in wet phosphoric acid were extracted using toluenesulfonic acid. The volume ratio of extractant to wet phosphoric acid was 3:1. The number of extraction stages was 3. The extraction temperature was 50℃ and the extraction time was 10min.

[0229] (2) Phase separation 1: The mixture of extractant and phosphoric acid after extraction is allowed to stand and separate to obtain oil phase 1 and aqueous phase 1. Oil phase 1 is organic extractant and a small amount of metal cations carried. The remaining aqueous phase 1 is phosphoric acid after metal cation removal. The oil phase is back-extracted and regenerated, and the aqueous phase enters the next stage. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0230] (3) Back-extraction 1: Ammonium sulfate is used as the back-extraction agent to back-extract oil phase 1. The volume ratio of back-extraction liquid to extractant is 1:1, the number of back-extraction stages is 1, the back-extraction temperature is 50℃, the back-extraction time is 60min, and the concentration of back-extraction agent is 30%.

[0231] (4) Phase separation 2: The extractant and the back-extractant solution after back-extraction are allowed to stand and separate to obtain oil phase 2 and water phase 2. Oil phase 2 is organic extractant, and the remaining water phase 2 is back-extractant and eluted metal ions. Oil phase 2 is returned to extraction 1 process, and water phase enters the next cooling crystallization 1 process. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0232] (5) Cooling crystallization 1: The liquid obtained after back-extraction is cooled and crystallized, with the final cooling temperature being 20℃.

[0233] (6) Filtration 1: The solid-liquid mixture obtained by cooling crystallization is filtered. The solid is ammonium aluminum sulfate and ammonium magnesium sulfate, and the liquid is crystallization mother liquor 1, which is returned to the cooling crystallization 1 section.

[0234] (7) Defluorination 1: The aqueous phase 1 obtained in (2) is concentrated and defluorinated. The concentration and defluorination temperature is 80℃, the pressure is 30kPa, the concentration endpoint is 40% phosphoric acid P2O5 content and specific gravity is 1.5.

[0235] (8) Defluorination 2: Use a defluorinating agent to perform secondary chemical defluorination on concentrated defluorinated wet phosphoric acid. The defluorinating agent is ammonium bicarbonate, and the defluorination temperature is 60℃.

[0236] (9) Extraction 2: Butyl acetate was used as the extractant to extract wet phosphoric acid after secondary defluorination in defluorination 2. The extraction temperature was 60℃, the volume ratio of the composite extractant to the wet phosphoric acid after secondary defluorination was 1:3, and the number of extraction stages was 3.

[0237] (10) Phase separation 3: The mixed solution of composite extractant and wet phosphoric acid is allowed to stand and separate. The organic oil phase coupled with phosphoric acid is the extract, and the remaining aqueous phase is water and a small amount of phosphoric acid. Aqueous phase 3 is returned to wet phosphoric acid extraction 2 for the next extraction. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0238] (11) Back-extraction 2: The oil phase 3 obtained by phase separation 3 is back-extracted with ferric sulfate. The back-extraction temperature is 60℃, the volume ratio of back-extractant to organic phase is 1:1, and the molar ratio of iron ions to phosphoric acid carried in the extractant is 1:3.15.

[0239] (12) Phase separation 4: The mixture of extract and back-extraction solution is allowed to stand and separate to obtain oil phase 4 and aqueous phase 4. The main component of aqueous phase 4 is H2PO4. - Fe 3+ H2O and very small amounts of SO4 2-The main component of oil phase 4 is H2SO4 carried by a composite extractant. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0240] (13) Back-extraction 3: Ammonia water is added to the oil phase 4 obtained by phase separation 4 for secondary back-extraction. The secondary back-extraction temperature is 50℃, the volume ratio of back-extraction agent to organic phase is 1:1, and the molar ratio of back-extraction agent NH3 to acid carried in the extractant is 1.1:1.

[0241] (14) Phase separation 5: The substance obtained from back-extraction 3 is phase separated. The oil phase of the composite extractant after back-extraction is returned to the extraction 2 process, and the aqueous phase is used for subsequent operations.

[0242] (15) Evaporation and concentration 1: The aqueous ammonium sulfate solution obtained from phase separation 5 is evaporated and concentrated to a specific gravity of 1.35, the concentration temperature is 80℃, and the pressure is 30kPa;

[0243] (16) Cooling crystallization 2: The concentrated liquid is cooled and crystallized, with the final cooling temperature being 40℃;

[0244] (17) Filtration 2: The product obtained by cooling crystallization is filtered to obtain solid ammonium salt (NH4)2SO4, and the liquid crystallization mother liquor is returned to (15) Evaporation and Concentration 1 section.

[0245] (18) Evaporation and concentration 2: The concentrated mother liquor obtained in (12) is evaporated and concentrated to a specific gravity of 1.50, a concentration temperature of 80℃, and a pressure of 20kPa;

[0246] (19) Cooling crystallization 3: Cool the concentrated liquid to obtain relatively pure iron dihydrogen phosphate. The crystallization temperature is 30℃.

[0247] Example 8

[0248] Nickel dihydrogen phosphate

[0249] (1) Extraction 1: Extraction of metal cations in wet phosphoric acid with p-toluenesulfonic acid. The volume ratio of extractant to wet phosphoric acid is 3:1. The number of extraction stages is 3. The extraction temperature is 50℃ and the extraction time is 10min.

[0250] (2) Phase separation 1: The mixture of extractant and phosphoric acid after extraction is allowed to stand and separate to obtain oil phase 1 and aqueous phase 1. Oil phase 1 is organic extractant and a small amount of metal cations carried. The remaining aqueous phase 1 is phosphoric acid after metal cation removal. The oil phase is back-extracted and regenerated, and the aqueous phase enters the next stage. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0251] (3) Back-extraction 1: Ammonium sulfate is used as the back-extraction agent to back-extract oil phase 1. The volume ratio of back-extraction liquid to extractant is 1:1, the number of back-extraction stages is 1, the back-extraction temperature is 50℃, the back-extraction time is 60min, and the concentration of back-extraction agent is 30%.

[0252] (4) Phase separation 2: The extractant and the back-extractant solution after back-extraction are allowed to stand and separate to obtain oil phase 2 and water phase 2. Oil phase 2 is organic extractant, and the remaining water phase 2 is back-extractant and eluted metal ions. Oil phase 2 is returned to extraction 1 process, and water phase enters the next cooling crystallization 1 process. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0253] (5) Cooling crystallization 1: The liquid obtained after back-extraction is cooled and crystallized, with the final cooling temperature being 20℃.

[0254] (6) Filtration 1: The solid-liquid mixture obtained by cooling crystallization is filtered. The solid is ammonium aluminum sulfate and ammonium magnesium sulfate, and the liquid is crystallization mother liquor 1, which is returned to the cooling crystallization 1 section.

[0255] (7) Defluorination 1: The aqueous phase 1 obtained in (2) is concentrated and defluorinated. The concentration and defluorination temperature is 80℃, the pressure is 30kPa, the concentration endpoint is 40% phosphoric acid P2O5 content and specific gravity is 1.5.

[0256] (8) Defluorination 2: The concentrated defluorinated wet phosphoric acid is subjected to secondary chemical defluorination using a defluorinating agent, ammonium bicarbonate, at a temperature of 60°C.

[0257] (9) Extraction 2: Dioctyl sulfoxide was used as the extractant to extract wet phosphoric acid after secondary defluorination in defluorination 2. The extraction temperature was 60℃, the volume ratio of the composite extractant to the wet phosphoric acid after secondary defluorination was 1:3, and the number of extraction stages was 3.

[0258] (10) Phase separation 3: The mixed solution of composite extractant and wet phosphoric acid is allowed to stand and separate. The organic oil phase coupled with phosphoric acid is the extract, and the remaining aqueous phase is water and a small amount of phosphoric acid. Aqueous phase 3 is returned to wet phosphoric acid extraction 2 for the next extraction. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0259] (11) Back-extraction 2: Nickel sulfate is used to back-extract the oil phase 3 obtained from phase separation 3. The back-extraction temperature is 60℃, the volume ratio of the back-extractant to the organic phase is 1:1, and the molar ratio of nickel ions to phosphoric acid carried in the extractant is 1:1.6.

[0260] (12) Phase separation 4: The mixture of extract and back-extraction solution is allowed to stand and separate to obtain oil phase 4 and aqueous phase 4. The main component of aqueous phase 4 is H2PO4. - Ni 2+ H2O and very small amounts of SO4 2-The main component of oil phase 4 is H2SO4 carried by a composite extractant. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0261] (13) Back-extraction 3: Ammonia water is added to the oil phase 4 obtained by phase separation 4 for secondary back-extraction. The secondary back-extraction temperature is 50℃, the volume ratio of back-extraction agent to organic phase is 1:1, and the molar ratio of back-extraction agent NH3 to acid carried in the extractant is 1.1:1.

[0262] (14) Phase separation 5: The substance obtained from back-extraction 3 is phase separated. The oil phase of the composite extractant after back-extraction is returned to the extraction 2 process, and the aqueous phase is used for subsequent operations.

[0263] (15) Evaporation and concentration 1: The aqueous ammonium sulfate solution obtained from phase separation 5 is evaporated and concentrated to a specific gravity of 1.35, the concentration temperature is 80℃, and the pressure is 30kPa;

[0264] (16) Cooling crystallization 2: The concentrated liquid is cooled and crystallized, with the final cooling temperature being 40℃;

[0265] (17) Filtration 2: The product obtained by cooling crystallization is filtered to obtain solid ammonium salt (NH4)2SO4, and the liquid crystallization mother liquor is returned to (15) Evaporation and Concentration 1 section.

[0266] (18) Evaporation and concentration 2: The concentrated mother liquor obtained in (12) is evaporated and concentrated to a specific gravity of 1.50, a concentration temperature of 80℃, and a pressure of 20kPa;

[0267] (19) Cooling crystallization 3: Cooling crystallize the concentrated liquid to obtain relatively pure nickel dihydrogen phosphate at a crystallization temperature of 30℃.

[0268] Example 9

[0269] Zinc dihydrogen phosphate

[0270] (1) Extraction 1: Metal cations in wet phosphoric acid were extracted with benzenesulfonic acid. The volume ratio of extractant to wet phosphoric acid was 3:1. The number of extraction stages was 3. The extraction temperature was 50℃ and the extraction time was 10min.

[0271] (2) Phase separation 1: The mixture of extractant and phosphoric acid after extraction is allowed to stand and separate to obtain oil phase 1 and aqueous phase 1. Oil phase 1 is organic extractant and a small amount of metal cations carried. The remaining aqueous phase 1 is phosphoric acid after metal cation removal. The oil phase is back-extracted and regenerated, and the aqueous phase enters the next stage. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0272] (3) Back-extraction 1: Ammonium sulfate is used as the back-extraction agent to back-extract oil phase 1. The volume ratio of back-extraction liquid to extractant is 1:1, the number of back-extraction stages is 1, the back-extraction temperature is 50℃, the back-extraction time is 60min, and the concentration of back-extraction agent is 30%.

[0273] (4) Phase separation 2: The extractant and the back-extractant solution after back-extraction are allowed to stand and separate to obtain oil phase 2 and water phase 2. Oil phase 2 is organic extractant, and the remaining water phase 2 is back-extractant and eluted metal ions. Oil phase 2 is returned to extraction 1 process, and water phase enters the next cooling crystallization 1 process. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0274] (5) Cooling crystallization 1: The liquid obtained after back-extraction is cooled and crystallized, with the final cooling temperature being 20℃.

[0275] (6) Filtration 1: The solid-liquid mixture obtained by cooling crystallization is filtered. The solid is ammonium aluminum sulfate and ammonium magnesium sulfate, and the liquid is crystallization mother liquor 1, which is returned to the cooling crystallization 1 section.

[0276] (7) Defluorination 1: The aqueous phase 1 obtained in (2) is concentrated and defluorinated. The concentration and defluorination temperature is 80℃, the pressure is 30kPa, the concentration endpoint is 40% phosphoric acid P2O5 content and specific gravity is 1.5.

[0277] (8) Defluorination 2: The concentrated defluorinated wet phosphoric acid is subjected to secondary chemical defluorination using a defluorinating agent, ammonium bicarbonate, at a temperature of 60°C.

[0278] (9) Extraction 2: Diphenyl sulfoxide was used as the extractant to extract wet phosphoric acid after secondary defluorination in defluorination 2. The extraction temperature was 60℃, the volume ratio of the composite extractant to the wet phosphoric acid after secondary defluorination was 1:3, and the number of extraction stages was 3.

[0279] (10) Phase separation 3: The mixed solution of composite extractant and wet phosphoric acid is allowed to stand and separate. The organic oil phase coupled with phosphoric acid is the extract, and the remaining aqueous phase is water and a small amount of phosphoric acid. Aqueous phase 3 is returned to wet phosphoric acid extraction 2 for the next extraction. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0280] (11) Back-extraction 2: Zinc sulfate is used to back-extract the oil phase 3 obtained from phase separation 3. The back-extraction temperature is 60℃, the volume ratio of back-extractant to organic phase is 1:1, and the molar ratio of zinc ions to phosphoric acid carried in the extractant is 1:1.6.

[0281] (12) Phase separation 4: The mixture of extract and back-extraction solution is allowed to stand and separate to obtain oil phase 4 and aqueous phase 4. The main component of aqueous phase 4 is H2PO4. - Zn 2+ H2O and very small amounts of SO4 2-The main component of oil phase 4 is H2SO4 carried by a composite extractant. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0282] (13) Back-extraction 3: Ammonia water is added to the oil phase 4 obtained by phase separation 4 for secondary back-extraction. The secondary back-extraction temperature is 50℃, the volume ratio of back-extraction agent to organic phase is 1:1, and the molar ratio of back-extraction agent NH3 to acid carried in the extractant is 1.1:1.

[0283] (14) Phase separation 5: The substance obtained from back-extraction 3 is phase separated. The oil phase of the composite extractant after back-extraction is returned to the extraction 2 process, and the aqueous phase is used for subsequent operations.

[0284] (15) Evaporation and concentration 1: The aqueous ammonium sulfate solution obtained from phase separation 5 is evaporated and concentrated to a specific gravity of 1.35, the concentration temperature is 80℃, and the pressure is 30kPa;

[0285] (16) Cooling crystallization 2: The concentrated liquid is cooled and crystallized, with the final cooling temperature being 40℃;

[0286] (17) Filtration 2: The product obtained by cooling crystallization is filtered to obtain solid ammonium salt (NH4)2SO4, and the liquid crystallization mother liquor is returned to (15) Evaporation and Concentration 1 section.

[0287] (18) Evaporation and concentration 2: The concentrated mother liquor obtained in (12) is evaporated and concentrated to a specific gravity of 1.50, a concentration temperature of 80℃, and a pressure of 20kPa;

[0288] (19) Cooling crystallization 3: Cooling crystallize the concentrated liquid to obtain relatively pure nickel dihydrogen phosphate at a crystallization temperature of 30℃.

[0289] Example 10

[0290] Sodium hydrogen phosphate:

[0291] (1) Extraction 1: Use xylenesulfonic acid as a metal cation extractant to extract metal cations in wet phosphoric acid. The volume ratio of extractant to wet phosphoric acid is 3:1, the number of extraction stages is 3, the extraction temperature is 50℃, and the extraction time is 10min.

[0292] (2) Phase separation 1: The mixture of extractant and phosphoric acid after extraction is allowed to stand and separate to obtain oil phase 1 and aqueous phase 1. Oil phase 1 is organic extractant and a small amount of metal cations carried. The remaining aqueous phase 1 is phosphoric acid after metal cation removal. The oil phase is back-extracted and regenerated, and the aqueous phase enters the next stage. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0293] (3) Back-extraction 1: Ammonium sulfate is used to back-extract oil phase 1. The volume ratio of back-extraction liquid to extractant is 1:1, the number of back-extraction stages is 1, the back-extraction temperature is 50℃, the back-extraction time is 60min, and the concentration of back-extraction agent is 30%.

[0294] (4) Phase separation 2: The extractant and the back-extraction solution after back-extraction are allowed to stand and separate. The oil phase is the organic extractant, and the remaining aqueous phase is the back-extraction agent and the eluted metal ions. The oil phase is returned to the extraction 1 process, and the aqueous phase enters the next cooling crystallization 1 process. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0295] (5) Cooling crystallization 1: The liquid obtained after back-extraction is cooled and crystallized, with the final cooling temperature being 20℃.

[0296] (6) Filtration 1: The solid-liquid mixture obtained by cooling crystallization is filtered. The solid is ammonium aluminum sulfate and ammonium magnesium sulfate, and the liquid is the crystallization mother liquor, which is returned to the cooling crystallization section.

[0297] (7) Defluorination 1: The aqueous phase 1 obtained in (2) is concentrated and defluorinated. The concentration and defluorination temperature is 80℃, the pressure is 30kPa, the concentration endpoint is 40% phosphoric acid P2O5 content and specific gravity is 1.5.

[0298] (8) Defluorination 2: The concentrated defluorinated wet phosphoric acid is subjected to secondary chemical defluorination using a defluorinating agent, ammonium bicarbonate, at a temperature of 60°C.

[0299] (9) Extraction 2: Hydrocarbon sulfoxide is used to extract wet phosphoric acid after secondary defluorination 2. The extraction temperature is 60℃, the volume ratio of composite extractant to dilute phosphoric acid is 1:3, and the number of extraction stages is 3.

[0300] (10) Phase separation 3: The mixed solution of composite extractant and dilute phosphoric acid is allowed to stand and separate. The organic oil phase coupled with phosphoric acid is the extract, and the remaining aqueous phase is water and a small amount of phosphoric acid. After the phase separation 3 operation, the oil phase 3 and the aqueous phase 3 are obtained. The aqueous phase 3 is returned to the extraction stage 2 for the next extraction. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0301] (11) Back-extraction 2: NaCl is used for back-extraction at a temperature of 60°C. The volume ratio of the back-extractant to the organic phase is 1:1, and the molar ratio of sodium ions to phosphoric acid carried in the extractant is 1:0.55.

[0302] (12) Phase separation 4: The mixture of extract and back-extraction solution is allowed to stand and separate into phases. The oil phase is subjected to a second back-extraction. The main component of the aqueous phase is H2PO4. - Na + H2O and a very small amount of Cl -The main component of the oil phase is HCl carried by the composite extractant. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0303] (13) Back-extraction 3: Add ammonia water to the extractant obtained by phase separation for secondary back-extraction. The secondary back-extraction temperature is 50℃, the volume ratio of back-extractant to organic phase is 1:1, and the molar ratio of back-extractant NH3 to acid carried in the extractant is 1.1:1.

[0304] (14) Phase separation 5: The substance obtained from back-extraction 3 is phase separated. The oil phase of the composite extractant after back-extraction is returned to the extraction 2 process, and the aqueous phase is used for subsequent operations.

[0305] (15) Evaporation and concentration 1: The ammonium chloride solution obtained from the second back-extraction is evaporated and concentrated to a specific gravity of 1.31, the concentration temperature is 80℃, and the pressure is 30kPa;

[0306] (16) Cooling crystallization 2: The concentrated liquid is cooled and crystallized, with the final cooling temperature being 40℃;

[0307] (17) Filtration 2: Filter the product obtained by cooling crystallization to obtain solid ammonium salt NH4Cl, and return the liquid crystallization mother liquor to the concentration section.

[0308] (18) Neutralization: The aqueous phase obtained in (12) was neutralized with NaOH. The temperature of the neutralization process was 60℃ and the pH of the neutralization endpoint was 6.0.

[0309] (19) Evaporation and concentration 2: The concentrated mother liquor obtained in (18) is evaporated and concentrated to a specific gravity of 1.55, a concentration temperature of 80℃, and a pressure of 20kPa;

[0310] (20) Cooling crystallization 3: The concentrated liquid is cooled and crystallized to obtain relatively pure sodium hydrogen phosphate at a crystallization temperature of 30℃.

[0311] Example 11

[0312] Potassium hydrogen phosphate:

[0313] (1) Extraction 1: Dodecylbenzenesulfonic acid was used as a metal cation extractant to extract metal cations in wet phosphoric acid. The volume ratio of extractant to wet phosphoric acid was 3:1, the number of extraction stages was 3, the extraction temperature was 50℃, and the extraction time was 10min.

[0314] (2) Phase separation 1: The mixture of extractant and phosphoric acid after extraction is allowed to stand and separate to obtain oil phase 1 and aqueous phase 1. Oil phase 1 is organic extractant and a small amount of metal cations carried. The remaining aqueous phase 1 is phosphoric acid after metal cation removal. The oil phase is back-extracted and regenerated, and the aqueous phase enters the next stage. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0315] (3) Back-extraction 1: Ammonium sulfate is used to back-extract oil phase 1. The volume ratio of back-extraction liquid to extractant is 1:1, the number of back-extraction stages is 1, the back-extraction temperature is 50℃, the back-extraction time is 60min, and the concentration of back-extraction agent is 30%.

[0316] (4) Phase separation 2: The extractant and the back-extraction solution after back-extraction are allowed to stand and separate. The oil phase is the organic extractant, and the remaining aqueous phase is the back-extraction agent and the eluted metal ions. The oil phase is returned to the extraction 1 process, and the aqueous phase enters the next cooling crystallization 1 process. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0317] (5) Cooling crystallization 1: The liquid obtained after back-extraction is cooled and crystallized, with the final cooling temperature being 20℃.

[0318] (6) Filtration 1: The solid-liquid mixture obtained by cooling crystallization is filtered. The solid is ammonium aluminum sulfate and ammonium magnesium sulfate, and the liquid is the crystallization mother liquor, which is returned to the cooling crystallization section.

[0319] (7) Defluorination 1: The aqueous phase 1 obtained in (2) is concentrated and defluorinated. The concentration and defluorination temperature is 80℃, the pressure is 30kPa, the concentration endpoint is 40% phosphoric acid P2O5 content and specific gravity is 1.5.

[0320] (8) Defluorination 2: The concentrated defluorinated wet phosphoric acid is subjected to secondary chemical defluorination using a defluorinating agent, ammonium bicarbonate, at a temperature of 60°C.

[0321] (9) Extraction 2: The wet phosphoric acid after defluorination 2 was extracted with di(2-ethylhexyl) hexyl phosphate. The extraction temperature was 60℃, the volume ratio of the composite extractant to dilute phosphoric acid was 1:3, and the number of extraction stages was 3.

[0322] (10) Phase separation 3: The mixed solution of composite extractant and dilute phosphoric acid is allowed to stand and separate. The organic oil phase coupled with phosphoric acid is the extract, and the remaining aqueous phase is water and a small amount of phosphoric acid. After the phase separation 3 operation, the oil phase 3 and the aqueous phase 3 are obtained. The aqueous phase 3 is returned to the extraction stage 2 for the next extraction. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0323] (11) Back-extraction 2: KCl is used for back-extraction at a temperature of 60°C. The volume ratio of the back-extractant to the organic phase is 1:1, and the molar ratio of potassium ions to phosphoric acid carried in the extractant is 1:0.55.

[0324] (12) Phase separation 4: The mixture of extract and back-extraction solution is allowed to stand and separate into phases. The oil phase is subjected to a second back-extraction. The main component of the aqueous phase is H2PO4. - K + H2O and a very small amount of Cl -The main component of the oil phase is HCl carried by the composite extractant. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0325] (13) Back-extraction 3: Add ammonia water to the extractant obtained by phase separation for secondary back-extraction. The secondary back-extraction temperature is 50℃, the volume ratio of back-extractant to organic phase is 1:1, and the molar ratio of back-extractant NH3 to acid carried in the extractant is 1.1:1.

[0326] (14) Phase separation 5: The substance obtained from back-extraction 3 is phase separated. The oil phase of the composite extractant after back-extraction is returned to the extraction 2 process, and the aqueous phase is used for subsequent operations.

[0327] (15) Evaporation and concentration 1: The ammonium chloride solution obtained from the second back-extraction is evaporated and concentrated to a specific gravity of 1.31, the concentration temperature is 80℃, and the pressure is 30kPa;

[0328] (16) Cooling crystallization 2: The concentrated liquid is cooled and crystallized, with the final cooling temperature being 40℃;

[0329] (17) Filtration 2: Filter the product obtained by cooling crystallization to obtain solid ammonium salt NH4Cl, and return the liquid crystallization mother liquor to the concentration section.

[0330] (18) Neutralization: The aqueous phase obtained in (12) was neutralized with KOH. The temperature of the neutralization process was 60℃ and the pH of the neutralization endpoint was 6.0.

[0331] (19) Evaporation and concentration 2: The concentrated mother liquor obtained in (18) is evaporated and concentrated to a specific gravity of 1.55, a concentration temperature of 80℃, and a pressure of 20kPa;

[0332] (20) Cooling crystallization 3: Cool the concentrated liquid to obtain relatively pure potassium hydrogen phosphate. The crystallization temperature is 30℃.

[0333] Example 12

[0334] Calcium hydrogen phosphate:

[0335] (1) Extraction 1: Dinonylnaphthalenesulfonic acid was used as the metal cation extractant to extract the metal cations in wet phosphoric acid. The volume ratio of extractant to wet phosphoric acid was 3:1, the number of extraction stages was 3, the extraction temperature was 50℃, and the extraction time was 10min.

[0336] (2) Phase separation 1: The mixture of extractant and phosphoric acid after extraction is allowed to stand and separate to obtain oil phase 1 and aqueous phase 1. Oil phase 1 is organic extractant and a small amount of metal cations carried. The remaining aqueous phase 1 is phosphoric acid after metal cation removal. The oil phase is back-extracted and regenerated, and the aqueous phase enters the next stage. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0337] (3) Back-extraction 1: Ammonium sulfate is used to back-extract oil phase 1. The volume ratio of back-extraction liquid to extractant is 1:1, the number of back-extraction stages is 1, the back-extraction temperature is 50℃, the back-extraction time is 60min, and the concentration of back-extraction agent is 30%.

[0338] (4) Phase separation 2: The extractant and the back-extraction solution after back-extraction are allowed to stand and separate. The oil phase is the organic extractant, and the remaining aqueous phase is the back-extraction agent and the eluted metal ions. The oil phase is returned to the extraction 1 process, and the aqueous phase enters the next cooling crystallization 1 process. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0339] (5) Cooling crystallization 1: The liquid obtained after back-extraction is cooled and crystallized, with the final cooling temperature being 20℃.

[0340] (6) Filtration 1: The solid-liquid mixture obtained by cooling crystallization is filtered. The solid is ammonium aluminum sulfate and ammonium magnesium sulfate, and the liquid is the crystallization mother liquor, which is returned to the cooling crystallization section.

[0341] (7) Defluorination 1: The aqueous phase 1 obtained in (2) is concentrated and defluorinated. The concentration and defluorination temperature is 80℃, the pressure is 30kPa, the concentration endpoint is 40% phosphoric acid P2O5 content and specific gravity is 1.5.

[0342] (8) Defluorination 2: The concentrated defluorinated wet phosphoric acid is subjected to secondary chemical defluorination using a defluorinating agent, ammonium bicarbonate, at a temperature of 60°C.

[0343] (9) Extraction 2: The wet phosphoric acid after defluorination 2 was extracted with dioctyl octyl phosphate. The extraction temperature was 60℃, the volume ratio of the composite extractant to dilute phosphoric acid was 1:3, and the number of extraction stages was 3.

[0344] (10) Phase separation 3: The mixed solution of composite extractant and dilute phosphoric acid is allowed to stand and separate. The organic oil phase coupled with phosphoric acid is the extract, and the remaining aqueous phase is water and a small amount of phosphoric acid. After the phase separation 3 operation, the oil phase 3 and the aqueous phase 3 are obtained. The aqueous phase 3 is returned to the extraction stage 2 for the next extraction. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0345] (11) Back-extraction 2: CaCl2 is used for back-extraction, the back-extraction temperature is 60℃, the volume ratio of back-extractant to organic phase is 1:1, and the molar ratio of calcium ions to phosphoric acid carried in the extractant is 1:1.05.

[0346] (12) Phase separation 4: The mixture of extract and back-extraction solution is allowed to stand and separate into phases. The oil phase is subjected to a second back-extraction. The main component of the aqueous phase is H2PO4. - Ca 2+ H2O and a very small amount of Cl -The main component of the oil phase is HCl carried by the composite extractant. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0347] (13) Back-extraction 3: Add ammonia water to the extractant obtained by phase separation for secondary back-extraction. The secondary back-extraction temperature is 50℃, the volume ratio of back-extractant to organic phase is 1:1, and the molar ratio of back-extractant NH3 to acid carried in the extractant is 1.1:1.

[0348] (14) Phase separation 5: The substance obtained from back-extraction 3 is phase separated. The oil phase of the composite extractant after back-extraction is returned to the extraction 2 process, and the aqueous phase is used for subsequent operations.

[0349] (15) Evaporation and concentration 1: The ammonium chloride solution obtained from the second back-extraction is evaporated and concentrated to a specific gravity of 1.31, the concentration temperature is 80℃, and the pressure is 30kPa;

[0350] (16) Cooling crystallization 2: The concentrated liquid is cooled and crystallized, with the final cooling temperature being 40℃;

[0351] (17) Filtration 2: Filter the product obtained by cooling crystallization to obtain solid ammonium salt NH4Cl, and return the liquid crystallization mother liquor to the concentration section.

[0352] (18) Neutralization: The aqueous phase obtained in (12) was neutralized using Ca(HCO3)2. The temperature of the neutralization process was 60℃, and the pH at the neutralization endpoint was 6.0.

[0353] (19) Evaporation and concentration 2: The concentrated mother liquor obtained in (18) is evaporated and concentrated to a specific gravity of 1.55, a concentration temperature of 80℃, and a pressure of 20kPa;

[0354] (20) Cooling crystallization 3: Cool the concentrated liquid to obtain relatively pure calcium hydrogen phosphate. The crystallization temperature is 30℃.

[0355] Example 13

[0356] Sodium phosphate:

[0357] (1) Extraction 1: Cinnamic acid was used as a metal cation extractant to extract metal cations in wet phosphoric acid. The volume ratio of extractant to wet phosphoric acid was 3:1, the number of extraction stages was 3, the extraction temperature was 50℃, and the extraction time was 10min.

[0358] (2) Phase separation 1: The mixture of extractant and phosphoric acid after extraction is allowed to stand and separate to obtain oil phase 1 and aqueous phase 1. Oil phase 1 is organic extractant and a small amount of metal cations carried. The remaining aqueous phase 1 is phosphoric acid after metal cation removal. The oil phase is back-extracted and regenerated, and the aqueous phase enters the next stage. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0359] (3) Back-extraction 1: Ammonium sulfate is used to back-extract oil phase 1. The volume ratio of back-extraction liquid to extractant is 1:1, the number of back-extraction stages is 1, the back-extraction temperature is 50℃, the back-extraction time is 60min, and the concentration of back-extraction agent is 30%.

[0360] (4) Phase separation 2: The extractant and the back-extraction solution after back-extraction are allowed to stand and separate. The oil phase is the organic extractant, and the remaining aqueous phase is the back-extraction agent and the eluted metal ions. The oil phase is returned to the extraction 1 process, and the aqueous phase enters the next cooling crystallization 1 process. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0361] (5) Cooling crystallization 1: The liquid obtained after back-extraction is cooled and crystallized, with the final cooling temperature being 20℃.

[0362] (6) Filtration 1: The solid-liquid mixture obtained by cooling crystallization is filtered. The solid is ammonium aluminum sulfate and ammonium magnesium sulfate, and the liquid is the crystallization mother liquor, which is returned to the cooling crystallization section.

[0363] (7) Defluorination 1: The aqueous phase 1 obtained in (2) is concentrated and defluorinated. The concentration and defluorination temperature is 80℃, the pressure is 30kPa, the concentration endpoint is 40% phosphoric acid P2O5 content and specific gravity is 1.5.

[0364] (8) Defluorination 2: The concentrated defluorinated wet phosphoric acid is subjected to secondary chemical defluorination using a defluorinating agent, ammonium bicarbonate, at a temperature of 60°C.

[0365] (9) Extraction 2: Tributyl phosphate was used to extract wet phosphoric acid after defluorination 2. The extraction temperature was 60℃, the volume ratio of the composite extractant to dilute phosphoric acid was 1:3, and the number of extraction stages was 3.

[0366] (10) Phase separation 3: The mixed solution of composite extractant and dilute phosphoric acid is allowed to stand and separate. The organic oil phase coupled with phosphoric acid is the extract, and the remaining aqueous phase is water and a small amount of phosphoric acid. After the phase separation 3 operation, the oil phase 3 and the aqueous phase 3 are obtained. The aqueous phase 3 is returned to the extraction stage 2 for the next extraction. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0367] (11) Back-extraction 2: NaCl is used for back-extraction at a temperature of 60°C. The volume ratio of the back-extractant to the organic phase is 1:1, and the molar ratio of sodium ions to phosphoric acid carried in the extractant is 1:0.35.

[0368] (12) Phase separation 4: The mixture of extract and back-extraction solution is allowed to stand and separate into phases. The oil phase is subjected to a second back-extraction. The main component of the aqueous phase is H2PO4. - Na + H2O and a very small amount of Cl -The main component of the oil phase is HCl carried by the composite extractant. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0369] (13) Back-extraction 3: Add ammonia water to the extractant obtained by phase separation for secondary back-extraction. The secondary back-extraction temperature is 50℃, the volume ratio of back-extractant to organic phase is 1:1, and the molar ratio of back-extractant NH3 to acid carried in the extractant is 1.1:1.

[0370] (14) Phase separation 5: The substance obtained from back-extraction 3 is phase separated. The oil phase of the composite extractant after back-extraction is returned to the extraction 2 process, and the aqueous phase is used for subsequent operations.

[0371] (15) Evaporation and concentration 1: The ammonium chloride solution obtained from the second back-extraction is evaporated and concentrated to a specific gravity of 1.31, the concentration temperature is 80℃, and the pressure is 30kPa;

[0372] (16) Cooling crystallization 2: The concentrated liquid is cooled and crystallized, with the final cooling temperature being 40℃;

[0373] (17) Filtration 2: Filter the product obtained by cooling crystallization to obtain solid ammonium salt NH4Cl, and return the liquid crystallization mother liquor to the concentration section.

[0374] (18) Neutralization: The aqueous phase obtained in (12) was neutralized with NaOH. The temperature of the neutralization process was 60℃ and the pH of the neutralization endpoint was 9.0.

[0375] (19) Evaporation and concentration 2: The concentrated mother liquor obtained in (18) is evaporated and concentrated to a specific gravity of 1.55, a concentration temperature of 80℃, and a pressure of 20kPa;

[0376] (20) Cooling crystallization 3: The concentrated liquid is cooled and crystallized to obtain relatively pure sodium phosphate at a crystallization temperature of 30℃.

[0377] Example 14

[0378] Ferric phosphate:

[0379] (1) Extraction 1: Fatty acid was used as a metal cation extractant to extract metal cations in wet phosphoric acid. The volume ratio of extractant to wet phosphoric acid was 3:1, the number of extraction stages was 3, the extraction temperature was 50℃, and the extraction time was 10min.

[0380] (2) Phase separation 1: The mixture of extractant and phosphoric acid after extraction is allowed to stand and separate to obtain oil phase 1 and aqueous phase 1. Oil phase 1 is organic extractant and a small amount of metal cations carried. The remaining aqueous phase 1 is phosphoric acid after metal cation removal. The oil phase is back-extracted and regenerated, and the aqueous phase enters the next stage. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0381] (3) Back-extraction 1: Ammonium sulfate is used to back-extract oil phase 1. The volume ratio of back-extraction liquid to extractant is 1:1, the number of back-extraction stages is 1, the back-extraction temperature is 50℃, the back-extraction time is 60min, and the concentration of back-extraction agent is 30%.

[0382] (4) Phase separation 2: The extractant and the back-extraction solution after back-extraction are allowed to stand and separate. The oil phase is the organic extractant, and the remaining aqueous phase is the back-extraction agent and the eluted metal ions. The oil phase is returned to the extraction 1 process, and the aqueous phase enters the next cooling crystallization 1 process. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0383] (5) Cooling crystallization 1: The liquid obtained after back-extraction is cooled and crystallized, with the final cooling temperature being 20℃.

[0384] (6) Filtration 1: The solid-liquid mixture obtained by cooling crystallization is filtered. The solid is ammonium aluminum sulfate and ammonium magnesium sulfate, and the liquid is the crystallization mother liquor, which is returned to the cooling crystallization section.

[0385] (7) Defluorination 1: The aqueous phase 1 obtained in (2) is concentrated and defluorinated. The concentration and defluorination temperature is 80℃, the pressure is 30kPa, the concentration endpoint is 40% phosphoric acid P2O5 content and specific gravity is 1.5.

[0386] (8) Defluorination 2: The concentrated defluorinated wet phosphoric acid is subjected to secondary chemical defluorination using a defluorinating agent. The preferred defluorinating agent is ammonium bicarbonate, and the defluorination temperature is 60℃.

[0387] (9) Extraction 2: Methyl isobutyl ketone was used as the extractant to extract the above dilute phosphoric acid. The extraction temperature was 60°C, the volume ratio of the composite extractant to the dilute phosphoric acid was 1:3, and the number of extraction stages was 3.

[0388] (10) Phase separation 3: The mixed solution of composite extractant and dilute phosphoric acid is allowed to stand and separate. The organic oil phase coupled with phosphoric acid is the extract, and the remaining aqueous phase is water and a small amount of phosphoric acid. After the phase separation 3 operation, the oil phase 3 and the aqueous phase 3 are obtained. The aqueous phase 3 is returned to the extraction stage 2 for the next extraction. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0389] (11) Back-extraction 2: Ferric sulfate is used for back-extraction at a temperature of 60°C. The volume ratio of the back-extractant to the organic phase is 1:1, and the molar ratio of iron ions to phosphoric acid carried in the extractant is 1:1.05.

[0390] (12) Phase separation 4: The mixture of extract and back-extraction solution is allowed to stand and separate into phases. The oil phase is subjected to a second back-extraction. The main component of the aqueous phase is H2PO4. - Fe 3+ H2O and very small amounts of SO4 2-The main component in the oil phase is H2SO4 carried by the composite extractant. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0391] (13) Back-extraction 3: Add ammonia water to the extractant obtained by phase separation for secondary back-extraction. The secondary back-extraction temperature is 50℃, the volume ratio of back-extractant to organic phase is 1:1, and the molar ratio of back-extractant NH3 to acid carried in the extractant is 1.1:1.

[0392] (14) Phase separation 5: The substance obtained from back-extraction 3 is separated into phases. The oil phase of the composite extractant after back-extraction is returned to the extraction process, and the aqueous phase is used for subsequent operations.

[0393] (15) Evaporation and concentration 1: The ammonium sulfate solution obtained from the second back-extraction is evaporated and concentrated to a specific gravity of 1.35, the concentration temperature is 80℃, and the pressure is 30kPa;

[0394] (16) Cooling crystallization 2: The concentrated liquid is cooled and crystallized, with the final cooling temperature being 40℃;

[0395] (17) Filtration 2: The product obtained by cooling crystallization is filtered to obtain solid ammonium salt (NH4)2SO4, and the liquid crystallization mother liquor is returned to the concentration section.

[0396] (18) Neutralization: Fe(OH)3 was used to neutralize the aqueous phase obtained in (18). The temperature of the neutralization process was 60℃ and the pH of the neutralization endpoint was 4.0.

[0397] (19) Cooling crystallization 3: The liquid obtained by neutralization in (19) is cooled and crystallized to obtain relatively pure iron phosphate at a crystallization temperature of 50℃.

[0398] Example 15

[0399] Cobalt phosphate:

[0400] (1) Extraction 1: Lauric acid was used as the metal cation extractant to extract the metal cations in wet phosphoric acid. The volume ratio of extractant to wet phosphoric acid was 3:1, the number of extraction stages was 3, the extraction temperature was 50℃, and the extraction time was 10min.

[0401] (2) Phase separation 1: The mixture of extractant and phosphoric acid after extraction is allowed to stand and separate to obtain oil phase 1 and aqueous phase 1. Oil phase 1 is organic extractant and a small amount of metal cations carried. The remaining aqueous phase 1 is phosphoric acid after metal cation removal. The oil phase is back-extracted and regenerated, and the aqueous phase enters the next stage. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0402] (3) Back-extraction 1: Ammonium sulfate is used to back-extract oil phase 1. The volume ratio of back-extraction liquid to extractant is 1:1, the number of back-extraction stages is 1, the back-extraction temperature is 50℃, the back-extraction time is 60min, and the concentration of back-extraction agent is 30%.

[0403] (4) Phase separation 2: The extractant and the back-extraction solution after back-extraction are allowed to stand and separate. The oil phase is the organic extractant, and the remaining aqueous phase is the back-extraction agent and the eluted metal ions. The oil phase is returned to the extraction 1 process, and the aqueous phase enters the next cooling crystallization 1 process. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0404] (5) Cooling crystallization 1: The liquid obtained after back-extraction is cooled and crystallized, with the final cooling temperature being 20℃.

[0405] (6) Filtration 1: The solid-liquid mixture obtained by cooling crystallization is filtered. The solid is ammonium aluminum sulfate and ammonium magnesium sulfate, and the liquid is the crystallization mother liquor, which is returned to the cooling crystallization section.

[0406] (7) Defluorination 1: The aqueous phase 1 obtained in (2) is concentrated and defluorinated. The concentration and defluorination temperature is 80℃, the pressure is 30kPa, the concentration endpoint is 40% phosphoric acid P2O5 content and specific gravity is 1.5.

[0407] (8) Defluorination 2: The concentrated defluorinated wet phosphoric acid is subjected to secondary chemical defluorination using a defluorinating agent, ammonium bicarbonate, at a temperature of 60°C.

[0408] (9) Extraction 2: The wet phosphoric acid after defluorination 2 was extracted with cyclohexanone isoamyl alcohol. The extraction temperature was 60℃, the volume ratio of the composite extractant to dilute phosphoric acid was 1:3, and the number of extraction stages was 3.

[0409] (10) Phase separation 3: The mixed solution of composite extractant and dilute phosphoric acid is allowed to stand and separate. The organic oil phase coupled with phosphoric acid is the extract, and the remaining aqueous phase is water and a small amount of phosphoric acid. After the phase separation 3 operation, the oil phase 3 and the aqueous phase 3 are obtained. The aqueous phase 3 is returned to the extraction stage 2 for the next extraction. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0410] (11) Back-extraction 2: Back-extraction is carried out using CoCl2 at a temperature of 60°C. The volume ratio of the back-extractant to the organic phase is 1:1, and the molar ratio of cobalt ions to phosphoric acid carried in the extractant is 1:1.6.

[0411] (12) Phase separation 4: The mixture of extract and back-extraction solution is allowed to stand and separate into phases. The oil phase is subjected to a second back-extraction. The main component of the aqueous phase is H2PO4. - Co 2+ H2O and a very small amount of Cl -The main component of the oil phase is HCl carried by the composite extractant. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0412] (13) Back-extraction 3: Add ammonia water to the extractant obtained by phase separation for secondary back-extraction. The secondary back-extraction temperature is 50℃, the volume ratio of back-extractant to organic phase is 1:1, and the molar ratio of back-extractant NH3 to acid carried in the extractant is 1.1:1.

[0413] (14) Phase separation 5: The substance obtained from back-extraction 3 is phase separated. The oil phase of the composite extractant after back-extraction is returned to the extraction 2 process, and the aqueous phase is used for subsequent operations.

[0414] (15) Evaporation and concentration 1: The ammonium chloride solution obtained from the second back-extraction is evaporated and concentrated to a specific gravity of 1.31, the concentration temperature is 80℃, and the pressure is 30kPa;

[0415] (16) Cooling crystallization 2: The concentrated liquid is cooled and crystallized, with the final cooling temperature being 40℃;

[0416] (17) Filtration 2: Filter the product obtained by cooling crystallization to obtain solid ammonium salt NH4Cl, and return the liquid crystallization mother liquor to the concentration section.

[0417] (18) Neutralization: The aqueous phase obtained in (12) was neutralized using cobalt carbonate. The temperature of the neutralization process was 60°C and the pH of the neutralization endpoint was 9.0.

[0418] (19) Evaporation and concentration 2: The concentrated mother liquor obtained in (18) is evaporated and concentrated to a specific gravity of 1.55, a concentration temperature of 80℃, and a pressure of 20kPa;

[0419] (20) Cooling crystallization 3: Cool the concentrated liquid to obtain relatively pure cobalt phosphate at a crystallization temperature of 30℃.

[0420] Example 16

[0421] Copper phosphate:

[0422] (1) Extraction 1: Naphthenic acid was used as a metal cation extractant to extract metal cations in wet phosphoric acid. The volume ratio of extractant to wet phosphoric acid was 3:1, the number of extraction stages was 3, the extraction temperature was 50℃, and the extraction time was 10min.

[0423] (2) Phase separation 1: The mixture of extractant and phosphoric acid after extraction is allowed to stand and separate to obtain oil phase 1 and aqueous phase 1. Oil phase 1 is organic extractant and a small amount of metal cations carried. The remaining aqueous phase 1 is phosphoric acid after metal cation removal. The oil phase is back-extracted and regenerated, and the aqueous phase enters the next stage. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0424] (3) Back-extraction 1: Ammonium sulfate is used to back-extract oil phase 1. The volume ratio of back-extraction liquid to extractant is 1:1, the number of back-extraction stages is 1, the back-extraction temperature is 50℃, the back-extraction time is 60min, and the concentration of back-extraction agent is 30%.

[0425] (4) Phase separation 2: The extractant and the back-extraction solution after back-extraction are allowed to stand and separate. The oil phase is the organic extractant, and the remaining aqueous phase is the back-extraction agent and the eluted metal ions. The oil phase is returned to the extraction 1 process, and the aqueous phase enters the next cooling crystallization 1 process. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0426] (5) Cooling crystallization 1: The liquid obtained after back-extraction is cooled and crystallized, with the final cooling temperature being 20℃.

[0427] (6) Filtration 1: The solid-liquid mixture obtained by cooling crystallization is filtered. The solid is ammonium aluminum sulfate and ammonium magnesium sulfate, and the liquid is the crystallization mother liquor, which is returned to the cooling crystallization section.

[0428] (7) Defluorination 1: The aqueous phase 1 obtained in (2) is concentrated and defluorinated. The concentration and defluorination temperature is 80℃, the pressure is 30kPa, the concentration endpoint is 40% phosphoric acid P2O5 content and specific gravity is 1.5.

[0429] (8) Defluorination 2: Use a defluorinating agent to perform secondary chemical defluorination on concentrated defluorinated wet phosphoric acid. The defluorinating agent is ammonium bicarbonate, and the defluorination temperature is 60℃.

[0430] (9) Extraction 2: Use 2-octanol to extract wet phosphoric acid after defluorination 2. The extraction temperature is 60℃, the volume ratio of composite extractant to dilute phosphoric acid is 1:3, and the number of extraction stages is 3.

[0431] (10) Phase separation 3: The mixed solution of composite extractant and dilute phosphoric acid is allowed to stand and separate. The organic oil phase coupled with phosphoric acid is the extract, and the remaining aqueous phase is water and a small amount of phosphoric acid. After the phase separation 3 operation, the oil phase 3 and the aqueous phase 3 are obtained. The aqueous phase 3 is returned to the extraction stage 2 for the next extraction. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0432] (11) Back-extraction 2: CuCl2 is used for back-extraction at a temperature of 60°C. The volume ratio of the back-extractant to the organic phase is 1:1, and the molar ratio of copper ions to phosphoric acid carried in the extractant is 1:1.6.

[0433] (12) Phase separation 4: The mixture of extract and back-extraction solution is allowed to stand and separate into phases. The oil phase is subjected to a second back-extraction. The main component of the aqueous phase is H2PO4. - Cu 2+ H2O and a very small amount of Cl -The main component of the oil phase is HCl carried by the composite extractant. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0434] (13) Back-extraction 3: Add ammonia water to the extractant obtained by phase separation for secondary back-extraction. The secondary back-extraction temperature is 50℃, the volume ratio of back-extractant to organic phase is 1:1, and the molar ratio of back-extractant NH3 to acid carried in the extractant is 1.1:1.

[0435] (14) Phase separation 5: The substance obtained from back-extraction 3 is phase separated. The oil phase of the composite extractant after back-extraction is returned to the extraction 2 process, and the aqueous phase is used for subsequent operations.

[0436] (15) Evaporation and concentration 1: The ammonium chloride solution obtained from the second back-extraction is evaporated and concentrated to a specific gravity of 1.31, the concentration temperature is 80℃, and the pressure is 30kPa;

[0437] (16) Cooling crystallization 2: The concentrated liquid is cooled and crystallized, with the final cooling temperature being 40℃;

[0438] (17) Filtration 2: Filter the product obtained by cooling crystallization to obtain solid ammonium salt NH4Cl, and return the liquid crystallization mother liquor to the concentration section.

[0439] (18) Neutralization: The aqueous phase obtained in (12) was neutralized using copper hydroxide. The temperature of the neutralization process was 60°C and the pH of the neutralization endpoint was 9.0.

[0440] (19) Evaporation and concentration 2: The concentrated mother liquor obtained in (18) is evaporated and concentrated to a specific gravity of 1.55, a concentration temperature of 80℃, and a pressure of 20kPa;

[0441] (20) Cooling crystallization 3: Cool the concentrated liquid to obtain relatively pure copper phosphate. The crystallization temperature is 30℃.

[0442] Example 17

[0443] Nickel phosphate:

[0444] (1) Extraction 1: Di(2-ethylhexyl) phosphate was used as a metal cation extractant to extract metal cations in wet phosphoric acid. The volume ratio of extractant to wet phosphoric acid was 3:1, the number of extraction stages was 3, the extraction temperature was 50℃, and the extraction time was 10min.

[0445] (2) Phase separation 1: The mixture of extractant and phosphoric acid after extraction is allowed to stand and separate to obtain oil phase 1 and aqueous phase 1. Oil phase 1 is organic extractant and a small amount of metal cations carried. The remaining aqueous phase 1 is phosphoric acid after metal cation removal. The oil phase is back-extracted and regenerated, and the aqueous phase enters the next stage. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0446] (3) Back-extraction 1: Ammonium sulfate is used to back-extract oil phase 1. The volume ratio of back-extraction liquid to extractant is 1:1, the number of back-extraction stages is 1, the back-extraction temperature is 50℃, the back-extraction time is 60min, and the concentration of back-extraction agent is 30%.

[0447] (4) Phase separation 2: The extractant and the back-extraction solution after back-extraction are allowed to stand and separate. The oil phase is the organic extractant, and the remaining aqueous phase is the back-extraction agent and the eluted metal ions. The oil phase is returned to the extraction 1 process, and the aqueous phase enters the next cooling crystallization 1 process. The phase separation temperature is 50℃ and the phase separation time is 3min.

[0448] (5) Cooling crystallization 1: The liquid obtained after back-extraction is cooled and crystallized, with the final cooling temperature being 20℃.

[0449] (6) Filtration 1: The solid-liquid mixture obtained by cooling crystallization is filtered. The solid is ammonium aluminum sulfate and ammonium magnesium sulfate, and the liquid is the crystallization mother liquor, which is returned to the cooling crystallization section.

[0450] (7) Defluorination 1: The aqueous phase 1 obtained in (2) is concentrated and defluorinated. The concentration and defluorination temperature is 80℃, the pressure is 30kPa, the concentration endpoint is 40% phosphoric acid P2O5 content and specific gravity is 1.5.

[0451] (8) Defluorination 2: Use a defluorinating agent to perform secondary chemical defluorination on concentrated defluorinated wet phosphoric acid. The defluorinating agent is ammonium bicarbonate, and the defluorination temperature is 60℃.

[0452] (9) Extraction 2: Wet phosphoric acid after defluorination 2 is extracted with substituted primary alcohol. The extraction temperature is 60℃, the volume ratio of composite extractant to dilute phosphoric acid is 1:3, and the number of extraction stages is 3.

[0453] (10) Phase separation 3: The mixed solution of composite extractant and dilute phosphoric acid is allowed to stand and separate. The organic oil phase coupled with phosphoric acid is the extract, and the remaining aqueous phase is water and a small amount of phosphoric acid. After the phase separation 3 operation, the oil phase 3 and the aqueous phase 3 are obtained. The aqueous phase 3 is returned to the extraction stage 2 for the next extraction. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0454] (11) Back-extraction 2: Back-extraction is performed using NiCl2 at a temperature of 60℃. The volume ratio of the back-extractant to the organic phase is 1:1, and NiCl2 is used in the back-extraction process. 2+ The molar ratio of phosphoric acid to that carried in the extractant is 1:1.6;

[0455] (12) Phase separation 4: The mixture of extract and back-extraction solution is allowed to stand and separate into phases. The oil phase is subjected to a second back-extraction. The main component of the aqueous phase is H2PO4. - Ni 2+ H2O and a very small amount of Cl -The main component of the oil phase is HCl carried by the composite extractant. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0456] (13) Back-extraction 3: Add ammonia water to the extractant obtained by phase separation for secondary back-extraction. The secondary back-extraction temperature is 50℃, the volume ratio of back-extractant to organic phase is 1:1, and the molar ratio of back-extractant NH3 to acid carried in the extractant is 1.1:1.

[0457] (14) Phase separation 5: The substance obtained from back-extraction 3 is phase separated. The oil phase of the composite extractant after back-extraction is returned to the extraction 2 process, and the aqueous phase is used for subsequent operations.

[0458] (15) Evaporation and concentration 1: The ammonium chloride solution obtained from the second back-extraction is evaporated and concentrated to a specific gravity of 1.31, the concentration temperature is 80℃, and the pressure is 30kPa;

[0459] (16) Cooling crystallization 2: The concentrated liquid is cooled and crystallized, with the final cooling temperature being 40℃;

[0460] (17) Filtration 2: Filter the product obtained by cooling crystallization to obtain solid ammonium salt NH4Cl, and return the liquid crystallization mother liquor to the concentration section.

[0461] (18) Neutralization: The aqueous phase obtained in (12) was neutralized with nickel hydroxide. The temperature of the neutralization process was 60°C and the pH of the neutralization endpoint was 9.0.

[0462] (19) Evaporation and concentration 2: The concentrated mother liquor obtained in (18) is evaporated and concentrated to a specific gravity of 1.55, a concentration temperature of 80℃, and a pressure of 20kPa;

[0463] (20) Cooling crystallization 3: The concentrated liquid is cooled and crystallized to obtain relatively pure nickel phosphate at a crystallization temperature of 30℃.

[0464] Comparative Example 1

[0465] Phosphate is produced directly using a second-step extraction process instead of a first-step extraction with a metal cation extractant.

[0466] (1) Defluorination 1: The raw material wet phosphoric acid is concentrated and defluorinated. The concentration and defluorination temperature is 80℃, the pressure is 30kPa, and the concentration endpoint phosphoric acid P2O5 content is 40% and the specific gravity is 1.55.

[0467] (2) Defluorination 2: The concentrated defluorinated wet phosphoric acid is subjected to secondary chemical defluorination using a defluorinating agent. The preferred defluorinating agent is ammonium bicarbonate, and the defluorination temperature is 60℃.

[0468] (3) Extraction: Tributyl phosphate was used as the extractant to extract the above dilute phosphoric acid. The extraction temperature was 60℃, the volume ratio of the composite extractant to the dilute phosphoric acid was 1:3, and the number of extraction stages was 3.

[0469] (4) Phase separation 1: The mixture of extractant and dilute phosphoric acid is allowed to stand and separate. The organic oil phase coupled with phosphoric acid is the extract, and the remaining aqueous phase is water and a small amount of phosphoric acid. After the phase separation 1 operation, the oil phase 1 and the aqueous phase 1 are obtained. The aqueous phase 1 is returned to the extraction section for the next extraction. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0470] (5) Back-extraction 1: Ferric sulfate is used for back-extraction at a temperature of 60°C. The volume ratio of the back-extractant to the organic phase is 1:1, and the molar ratio of iron ions to phosphoric acid carried in the extractant is 1.05:1.

[0471] (6) Phase separation 2: The mixture of extract and back-extraction solution is allowed to stand and separate into phases. The oil phase is subjected to a second back-extraction. The main component of the aqueous phase is H2PO4. - Fe 2+ H2O and very small amounts of SO4 2- The main component in the oil phase is H2SO4 carried by the composite extractant. The phase separation temperature is 60℃ and the phase separation time is 3min.

[0472] (7) Back-extraction 2: Add ammonia water to the extractant obtained by phase separation for secondary back-extraction. The secondary back-extraction temperature is 50℃, the volume ratio of back-extractant to organic phase is 1:1, and the molar ratio of back-extractant NH3 to acid carried in the extractant is 1.1:1.

[0473] (8) Phase separation 3: The substance obtained from back-extraction 2 is separated into phases. The oil phase of the composite extractant after back-extraction is returned to the extraction process, and the aqueous phase is used for subsequent operations.

[0474] (9) Evaporation and concentration 1: The ammonium sulfate solution obtained from the second back-extraction is evaporated and concentrated to a specific gravity of 1.35, the concentration temperature is 80℃, and the pressure is 30kPa;

[0475] (10) Cooling crystallization 2: The concentrated liquid is cooled and crystallized, with the final cooling temperature being 40℃;

[0476] (11) Filtration 2: The product obtained by cooling crystallization is filtered to obtain solid ammonium salt (NH4)2SO4, and the liquid crystallization mother liquor is returned to the concentration section.

[0477] (12) Neutralization: Fe(OH)3 was used to neutralize the aqueous phase obtained in (7). The temperature of the neutralization process was 60℃ and the pH of the neutralization endpoint was 4.0.

[0478] (13) Cooling crystallization 3: The liquid obtained by neutralization in (19) is cooled and crystallized to obtain relatively pure iron phosphate at a crystallization temperature of 50℃.

[0479] Table 1 shows the extractants used and product purity in the examples.

[0480]

[0481]

Claims

1. A method for preparing phosphate, characterized in that, The method includes: A. Extraction 1: Extract metal cations from wet-process phosphoric acid using extractant 1. After extraction, phase separation is performed to obtain oil phase 1 and aqueous phase 1. The extractant 1 is an extractant for acidic metal ions. B. Defluorination 1: The aqueous phase 1 is concentrated and defluorinated. The concentration and defluorination temperature is 40-100℃, the pressure is 10-90kPa, the concentration endpoint has a phosphoric acid P2O5 content of 30-50%, and a specific gravity of 1.4-2.

2. C. Defluorination 2: The wet-process phosphoric acid after defluorination 1 is chemically defluorinated a second time using a defluorinating agent; D. Extraction 2: Extract the wet phosphoric acid after the secondary defluorination in step C with extractant 2. After extraction, separate the phases to obtain oil phase 2 and aqueous phase 2. E. Back-extraction 1: Back-extracting the oil phase 2 using back-extracting agent 2 to separate the phases into oil phase 3 and aqueous phase 3. The back-extracting agent 2 is a solution of at least one of the soluble sulfate, hydrogen sulfate, hydrochloride, and nitrate of metal M. M is Li, Mg, Al, Mn, Fe, Co, Ni, Cu, Zn, Cr, Ge, Zr, Sn, Sb, or Ba, and the molar ratio of M ions to phosphate carried in oil phase 2 is 2:1 to 1:

2. F. pH adjustment: The pH of the aqueous phase 3 obtained in step E is adjusted to 2.0–11.0 using the carbonate, bicarbonate, and hydroxide of metal M; G. Evaporate and concentrate to a specific gravity of 1.45–1.75, cool to crystallize, and separate the solid and liquid to obtain phosphate with a purity of over 97%; The extractant 1 is at least one of the following: di(2-ethylhexyl) phosphate, 2-ethylhexyl phosphate, di(2,4,4-trimethylpentyl)phosphine, 2-ethylhexylphosphonic acid mono(2-ethylhexyl) ester, N,N-N-n-octylaminedimethylenephenylphosphonic acid, N,N-N-n-hexylaminedimethylenephenylphosphonic acid, toluenesulfonic acid, p-toluenesulfonic acid, benzenesulfonic acid, xylenesulfonic acid, dodecylbenzenesulfonic acid, dinonylnaphthalenesulfonic acid, cinnamic acid, fatty acids, lauric acid, and naphthenic acid. The extractant 2 is at least one selected from ethyl acetate, amyl acetate, butyl acetate, dioctyl sulfoxide, diphenyl sulfoxide, hydrocarbon sulfoxide, di(2-ethylhexyl) hexyl phosphate, dioctyl phosphate, tributyl phosphate, methyl isobutyl ketone, cyclohexanone, isoamyl alcohol, sec-octanol, substituted primary alcohols, trialkylmethylamine, didecylamine, trioctylamine, and trinonylamine. The method further includes the treatment of oil phase 1: back-extracting oil phase 1 with back-extracting agent 1, and then separating the phases to obtain oil phase 4 and aqueous phase 4; The oil phase 4 is returned and recycled as extractant 1; The aqueous phase 4 is cooled and crystallized to obtain aluminum ammonium sulfate, magnesium ammonium sulfate solids and crystallization mother liquor 1. The crystallization mother liquor 1 is returned to the aqueous phase 4 for recycling. The back-extraction agent 1 is at least one of sulfuric acid and soluble sulfate solution.

2. The method for preparing phosphate according to claim 1, characterized in that, In step A, the volume ratio of extractant 1 to wet phosphoric acid is 1:3 to 3:1, and the extraction temperature is 10 to 80°C.

3. The method for preparing phosphate according to claim 2, characterized in that, The number of extraction stages in step A is 1 to 6.

4. The method for preparing phosphate according to claim 2, characterized in that, The extraction time in step A is 5–60 min.

5. The method for preparing phosphate according to claim 2, characterized in that, The phase separation temperature in step A is 10–80°C, and the phase separation time is 0.5–10 min.

6. The method for preparing phosphate according to claim 1, characterized in that, The volume ratio of the back-extraction agent 1 to the oil phase 1 is 1:3 to 3:1, the number of back-extraction stages is 1 to 6, the back-extraction temperature is 10 to 80°C, and the back-extraction time is 5 to 60 min.

7. The method for preparing phosphate according to claim 1, characterized in that, The stripping agent 1 is at least one of sulfuric acid and ammonium sulfate solution, and the concentration of the stripping agent 1 is 1-50%.

8. The method for preparing phosphate according to claim 1, characterized in that, The phase separation temperature for obtaining oil phase 4 and water phase 4 is 10–80℃, and the phase separation time is 0.5–10 min.

9. The method for preparing phosphate according to claim 1, characterized in that, The cooling endpoint temperature for the aqueous phase 4 cooling crystallization is 10–60°C.

10. The method for preparing phosphate according to claim 1 or 2, characterized in that, The defluorinating agent is at least one of sodium carbonate and ammonium bicarbonate, calcium carbonate, sodium phosphate, sodium hydroxide, calcium hydroxide, calcium oxide, and sodium bicarbonate.

11. The method for preparing phosphate according to claim 1 or 2, characterized in that, The temperature for secondary defluorination is 10–80℃.

12. The method for preparing phosphate according to claim 1 or 2, characterized in that, The extraction temperature in step D is 10–80°C, the volume ratio of extractant 2 to wet phosphoric acid after secondary defluorination is 3:1–1:3, and the number of extraction stages is 1–6. The phase separation temperature in step D is 10–80°C, and the phase separation time is 0.5–10 min.

13. The method for preparing phosphate according to claim 1 or 2, characterized in that, The aqueous phase 2 is returned to extraction 2 in step D.

14. The method for preparing phosphate according to claim 1 or 2, characterized in that, The back-extraction temperature in step E is 10–80°C, and the volume ratio of back-extraction agent 2 to oil phase 2 is 3:1–1:

3.

15. The method for preparing phosphate according to claim 1 or 2, characterized in that, The phase separation temperature in step E is 10–80°C, and the phase separation time is 0.5–10 min.

16. The method for preparing phosphate according to claim 1 or 2, characterized in that, The method also includes the treatment of oil phase 3: back-extracting oil phase 3 with ammonia or water to separate oil phase 5 and water phase 5; The aqueous phase 5 was evaporated and concentrated, cooled and crystallized, and then separated into solid and liquid phases to obtain solid ammonium salt and liquid crystallization mother liquor. The oil phase 5 is returned to step D and recycled as extractant 2.

17. The method for preparing phosphate according to claim 16, characterized in that, The back-extraction temperature of oil phase 3 is 10-80℃, the volume ratio of ammonia water to organic phase is 3:1-1:3, and the molar ratio of NH3 in ammonia water to acid carried in oil phase 3 is 3:1-1:

3.

18. The method for preparing phosphate according to claim 16, characterized in that, The aqueous phase 5 is evaporated and concentrated to a specific gravity of 1.1–1.7, at a concentration temperature of 10–100°C, at a pressure of 10–80 kPa, and at a cooling point of 10–60°C for the final crystallization. The liquid crystallization mother liquor obtained by evaporation, concentration, cooling, crystallization, and solid-liquid separation in aqueous phase 5 is returned to aqueous phase 5 for recycling.

19. The method for preparing phosphate according to claim 16, characterized in that, Ammonia was used to back-extract oil phase 3.

20. The method for preparing phosphate according to claim 1 or 2, characterized in that, When the M ion in the wet-process phosphoric acid is in a low valence state and a phosphate with a high valence state of M ion needs to be prepared, the method further includes oxidizing the M ion before neutralization; the oxidant is at least one of oxygen, ozone, hydrogen peroxide, sodium peroxide, barium peroxide, chlorine, potassium permanganate, potassium dichromate, sodium hypochlorite, perchloric acid, and nitric acid.

21. The method for preparing phosphate according to claim 1 or 2, characterized in that, The oxidation temperature is 20–80℃.

22. The method for preparing phosphate according to claim 1 or 2, characterized in that, The pH adjustment temperature in step F is 10–100°C.

23. The method for preparing phosphate according to claim 1 or 2, characterized in that, The evaporation and concentration in step G is carried out at a temperature of 10–90°C and a pressure of 10–50 kPa.

24. The method for preparing phosphate according to claim 23, characterized in that, The crystallization temperature in step G is 10–60°C.

25. The method for preparing phosphate according to claim 23, characterized in that, The mother liquor from the solid-liquid separation in step G is returned to the evaporation and concentration process in step G.

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