Phosphorite wet processing method

Abstract

The invention relates to a phosphorite wet processing method, and belongs to the technical field of chemical engineering. The phosphorite wet processing method comprises the following steps: A, phosphorite acid leaching: carrying out acidolysis leaching on a phosphorite raw material, and carrying out solid-liquid separation to obtain wet-process phosphoric acid; b, removing metal ions from wet-process phosphoric acid; and C, wet-process phosphoric acid defluorination: extracting fluosilicic acid in wet-process phosphoric acid at 10-70 DEG C by adopting a wet-process phosphoric acid defluorination extraction agent, and carrying out phase splitting to obtain defluorinated purified phosphoric acid and an organic phase loaded with fluosilicate. According to the invention, high-value utilization of low-grade phosphorite resources is realized. According to the method, impurities can be fully removed, the loss rate of phosphoric acid is low, the purity of the finally obtained phosphate product reaches 97% or above, the product quality is high, and the requirements of different fields for high-quality phosphate can be met. The method is simple in process, low in energy consumption and low in cost.

Description

Technical Field

[0001] The invention relates to a method for wet processing of phosphate rock, belonging to the technical field of chemical industry. Background Art

[0002] Phosphorus chemical industry is a vital component of the chemical industry. Wet-process phosphoric acid, a fundamental intermediate product in the phosphate fertilizer and compound fertilizer industries, plays a vital role and enjoys high demand. Currently, phosphoric acid production relies on high-grade phosphate rock (there are no strict standards for phosphate rock grade; industry insiders typically categorize it as follows: high-grade: P2O5 ≥ 30%, medium-grade: 24% ≤ P2O5 < 30%, and low-grade: P2O5 < 24%. Enterprise standards stipulate that phosphate rock used for wet-process processing should have a P2O5 content of ≥ 28%.). However, as rich phosphate deposits gradually become depleted and production costs rise, the utilization of medium- and low-grade phosphate rock has attracted attention.

[0003] The wet processing of low- and medium-grade phosphate rock produces a large amount of impurities, which can affect product quality and subsequent applications. Obtaining high-purity products through wet processing of phosphate rock is of great significance.

[0004] CN201410035054.1 discloses a method for preparing sodium dihydrogen phosphate using wet-process phosphoric acid. This method involves reacting carbonamide urea (CO(NH2)2) with wet-process phosphoric acid to produce an intermediate, which is then reacted with sodium hydroxide to produce the sodium dihydrogen phosphate product. The method features a short process route, low energy consumption, stable product quality, low production costs, convenient operation, and safe production. The byproduct slurry can also be fully recycled and reused. The entire production process is environmentally friendly, clean, and pollution-free, with no waste gas, wastewater, or waste residue emitted. This method addresses the policy call for energy conservation, emission reduction, and clean production, and overcomes the problems of complex processes, unstable product quality, high energy consumption, and environmental pollution in existing technologies. The resulting sodium dihydrogen phosphate product has a purity of ≥98%. However, its purity is insufficient, and sample diversity is poor.

[0005] CN201710509626.9 Method and device for producing potassium dihydrogen phosphate from wet-process phosphoric acid discloses a method and device 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; double decomposition; filter press filtration: index adjustment; crystallization: separation and drying. The solid is sent to a fluidized bed dryer for drying to obtain a finished industrial-grade potassium dihydrogen phosphate, and the mother liquor is recycled. The present invention uses wet-process phosphoric acid extraction instead of thermal phosphoric acid to prepare industrial-grade potassium dihydrogen phosphate, achieving the purpose of reducing production costs and improving product market competitiveness. It also has the problems of insufficient purity and poor sample diversity.

[0006] CN201010543993.9 discloses a method for producing industrial potassium dihydrogen phosphate and nitrogen-phosphorus-potassium compound fertilizer by an extraction process using wet-process purified phosphoric acid as the raw material. The method uses wet-process purified phosphoric acid and potassium chloride as the raw materials, and uses octanol and octyl ammonium as the extractants for extraction. The aqueous phase is cooled and crystallized to obtain the potassium dihydrogen phosphate product. The organic phase is stirred with a saturated ammonium chloride solution for back extraction, and the mother liquor after back extraction is cooled and crystallized to obtain the nitrogen-phosphorus-potassium compound fertilizer as a byproduct. The product obtained by this method is highly pure. By using octanol and octyl ammonium as the extractants to modify the chemical composition and H value of the aqueous phase, it can achieve both selective separation of specific components and broad-spectrum separation of multiple components. The method provides high extraction efficiency and good selectivity, and the extractant is easily regenerated by back extraction. However, the method also suffers from insufficient purity and poor sample diversity.

[0007] It has the problems of low purity and poor sample diversity.

[0008] CN201610988560.1 discloses a method for producing potassium dihydrogen phosphate using wet-process phosphoric acid. Phosphorus and pulp and barium carbonate are added to the wet-process phosphoric acid in sequence to reduce the concentration of sulfate ions in the phosphoric acid to 0.5 g / L. After standing and settling, the supernatant acid is taken as refined desulfurized dilute phosphoric acid, and the refined desulfurized dilute phosphoric acid is extracted with a composite extractant prepared by mixing octylamine and octanol in any proportion. The mixture is allowed to stand for phase separation, and the organic phase of the coupled phosphoric acid is the extract, and the remaining phosphoric acid is the raffinate. Potassium chloride is added to the extract to react, and the mixture is allowed to stand for phase separation, resulting in a lower layer of a crude potassium dihydrogen phosphate solution and an upper layer of an organic phase rich in hydrochloric acid, a small amount of phosphoric acid and potassium chloride. After filtering the crude potassium dihydrogen phosphate solution, potassium hydroxide is added to the filtrate to adjust the solution H to about 6. After standing and settling, impurities are filtered out to obtain a refined potassium dihydrogen phosphate solution. Purified phosphoric acid is added to the refined potassium dihydrogen phosphate solution to adjust the pH to 4-5. The obtained solution is concentrated, crystallized, filtered and dried to obtain a potassium dihydrogen phosphate product. It has the problems of low purity and poor sample diversity.

[0009] CN202210262414.6 discloses a feed-grade monocalcium phosphate preparation process, comprising the following steps: (1) using phosphate rock with a grade of 14-25%, extracting and purifying it with dihydrate to produce I phosphate, then neutralizing I phosphate with lime milk to produce a dicalcium phosphate slurry, subjecting the calcium phosphate slurry to solid-liquid separation to obtain a semi-finished product of dicalcium phosphate, which is first dried to obtain a dicalcium phosphate with a moisture content of less than 4%; (2) using phosphate rock with a grade higher than 29%, extracting it with half water and dihydrate, and then concentrating, defluorinating, and dearsenicizing to obtain a phosphoric acid II with a concentration greater than 50%; (3) reacting the phosphoric acid II obtained in step (2) with the calcium phosphate obtained in step (1) to produce a monocalcium phosphate semi-finished product, drying it after aging, and then screening it to obtain a monocalcium phosphate finished product. The present invention can make extensive use of low-grade phosphate rock that is difficult to apply in traditional processes, reduce the proportion of high-grade phosphate rock used, and save about 45% of coal consumption per unit product. It has the problems of insufficient purity and poor sample diversity.

[0010] CN200910272379.0 discloses a method for removing metal ions from wet-process phosphoric acid, comprising the following steps: 1) adding alkaline partially hydrolyzed polyacrylonitrile fibers to deionized water and stirring, followed by natural filtration, and then rinsing the filter cake until the pH of the liquid flowing out from the bottom of the Büchner funnel reaches 8. When no liquid flows out from the bottom, wet polyacrylonitrile ion exchange fibers are produced; 2) filling the wet polyacrylonitrile ion exchange fibers between the polypropylene cotton of a fiber filter and a cylindrical shell, connecting the phosphoric acid outlet pipe of an electrodialysis device to the inlet pipe of the fiber filter, and the outlet pipe of the fiber filter is connected to a phosphoric acid storage tank. The outlet pipe is then pumped into the electrodialysis device for circulation, passing the wet-process phosphoric acid through the electrodialysis device, while the fiber filter and the electrodialysis device are operated simultaneously. However, the electrodialysis device used in this method, while simple and easy to operate, has a small capacity and cannot be applied to large-scale industrial production.

[0011] CN116022752A discloses a process for purifying raffinate acid using a composite extractant, and specifically discloses the following steps: diluting the raffinate acid having an original mass concentration of not less than 40% to a mass concentration of 20-25%, then filtering to remove obvious particulate matter; mixing the diluted and filtered raffinate acid with an extracting liquid, and extracting to obtain an organic phase 1 and an aqueous phase 1; the volume ratio of the extracting liquid to the raffinate acid is 1:1-4:1; the extracting liquid is a mixture of an acidic extractant and a diluent; back-extracting the organic phase 1 at least three times with a stripping liquid, and after back-extraction, standing and separating the phases to obtain an organic phase 2 and an aqueous phase 2, wherein the obtained organic phase 2 is the purified regenerated extract; the stripping liquid is a strong acid or a medium-strong acid, and the volume ratio of the stripping liquid to the organic phase 1 is 3:1-1:1. An extracting solution prepared by using an acidic extractant and a diluent is used to directly extract impure metal cations such as iron ions, magnesium ions, aluminum ions, and manganese ions from the raffinate acid. The acidic extractant has good selectivity for extracting metal ions, carries little phosphoric acid during the extraction process, and can achieve a phosphoric acid recovery rate of over 90%. The iron ions, aluminum ions, magnesium ions, and manganese ions in the raffinate acid purified by the method of the present invention are all adsorbed into the extractant, thereby effectively purifying the raffinate acid. The iron ion recovery rate can reach 90%, the aluminum ion recovery rate can reach 50%, the magnesium ion recovery rate can reach 50%, and the manganese ion recovery rate can reach 50%. However, the phosphoric acid recovery rate of this raffinate acid purification process is still low, and the metal ions cannot be separated, making them difficult to recycle and reuse, resulting in low added value.

[0012] CN 202011029531.5 discloses a method for preparing anhydrous silicon tetrafluoride and hydrogen fluoride gas mixture from fluorine associated with phosphate rock. This method involves reacting aqueous fluorosilicic acid with washed sulfuric acid to produce fluorine-containing sulfuric acid and a mixed gas of aqueous silicon tetrafluoride and hydrogen fluoride. The mixed gas is then washed with hot concentrated sulfuric acid to obtain anhydrous silicon tetrafluoride and hydrogen fluoride gas mixture. Although this method is simple to operate, it requires a large amount of concentrated sulfuric acid to absorb water, resulting in the production of a large amount of 70-80wt% sulfuric acid. Summary of the Invention

[0013] The object of the present invention is to provide a method for wet processing of phosphate rock.

[0014] To achieve the purpose of the present invention, the method for wet processing of phosphate rock comprises:

[0015] A. Acid leaching of phosphate rock: The phosphate rock raw material is acid-lyzed and leached, and the solid-liquid separation is carried out to obtain wet-process phosphoric acid;

[0016] B. Removing metal ions from wet-process phosphoric acid;

[0017] C. Wet-process phosphoric acid defluorination: using a wet-process phosphoric acid defluorination extractant to extract fluorosilicic acid in wet-process phosphoric acid at 10-70° C., and separating the phases to obtain defluorinated purified phosphoric acid and an organic phase loaded with fluorosilicate; the wet-process phosphoric acid defluorination extractant is a compound of an organic amine, a cosolvent, and a diluent in a volume ratio of 10-100:0-30:0-90, and the content of the cosolvent and the diluent is not 0; the organic amine includes a primary amine RNH2, a secondary amine R2NH, a tertiary amine R3N, and a quaternary ammonium salt R4N + OH - At least one of, R is a hydrocarbon group; the cosolvent includes at least one of n-butanol, n-hexanol, cyclohexanol, 2-ethyl-1-hexanol, and n-octanol; the diluent includes at least one of sulfonated kerosene, n-dodecane, isomeric dodecane, n-tetradecane, n-hexadecane, and 260# solvent oil.

[0018] The phosphate rock raw material in step A comprises at least one of phosphate rock, white fertilizer, and slag acid, preferably, the phosphate rock raw material contains 5wt% to 35wt% of P2O5;

[0019] Preferably, the acid used for the acidolysis is sulfuric acid or dilute phosphoric acid, and the acid concentration of the acidolysis is ≥5%; the solid-liquid ratio of the acidolysis is 0.1-15, the reaction temperature is 0-90° C., the pH value of the reaction material is controlled within 0.5-1.5, and the stirring reaction is carried out for 0.5-12 hours.

[0020] The method for removing metal ions from wet-process phosphoric acid in step B is extraction or ion exchange;

[0021] Preferably, a decalcification pretreatment is performed before the extraction and removal of metal ions, the amount of the decalcifying agent is 0.8 to 1.2 times the theoretical amount, the reaction temperature is 40 to 60° C., the reaction time is 20 to 40 minutes, the aging time is 0.5 to 2 hours, and the extraction and removal of metal ions are performed after solid-liquid separation; the decalcifying agent is preferably at least one of sulfuric acid, ammonium sulfate, sodium sulfate, and potassium sulfate;

[0022] The method for extracting and removing metal ions preferably adopts the method disclosed in the invention patent publication number CN119265410A;

[0023] Or the method for extracting and removing metal ions preferably adopts the method disclosed in the invention patent publication number CN119503888A, wherein the temperature of step (3) of CN119503888A is 25-85°C, and the volume ratio of O / A is 1:5-10:1; step (4) adopts method 3, and in method 3 of step (4), ammonia is passed through to adjust the pH to 7-12; the temperature of ammonia evaporation is 30-80°C;

[0024] Alternatively, a mixed solution of oxalic acid and ammonium oxalate is used as a stripping agent instead of the stripping agent in step (3) of CN119503888A, wherein the ammonium oxalate concentration in step (3) is 1 wt.% to 20 wt.%, the oxalic acid concentration is 0 wt.% to 20 wt.%, the temperature is 25 to 85°C, and the volume ratio O / A is 1:5 to 10:1; step (4) adopts method 3, wherein ammonia is passed through method 3 to adjust the pH to 7 to 12; and the temperature of ammonia evaporation is 30 to 80°C;

[0025] The method for removing metal ions by ion exchange preferably adopts the method disclosed in the invention patent publication number CN116786096A.

[0026] In a specific embodiment, the wet-process phosphoric acid defluorination extractant is prepared by mixing an organic amine, a co-solvent, and a diluent in a volume ratio of 10-15:10-20:65-75.

[0027] In a specific embodiment, the step C further comprises:

[0028] Stripping: reacting the fluorosilicate-loaded organic phase with a stripping agent, separating the phases to obtain a fluorosilicate solution and a regenerated organic phase, wherein the stripping agent is an inorganic alkali aqueous solution or an inorganic salt aqueous solution;

[0029] Extractant regeneration: adding ammonia water to the regenerated organic phase to obtain an ammonium salt solution and a regenerated extractant, or heating the regenerated organic phase to obtain an aqueous solution and a regenerated extractant. The regenerated extractant is returned to step C for recycling, and the aqueous solution is returned to prepare a stripping agent.

[0030] In a specific embodiment, the method further comprises performing silicon replenishment before extraction in step C: adding a silicon replenishing agent to the wet-process phosphoric acid raw material, with the silicon replenishing amount being 1.1 to 1.2 times the theoretical value, reacting at 10 to 90° C. for 20 to 40 minutes, aging for 1 hour, and filtering;

[0031] The silicon supplementing agent is preferably silicon dioxide, diatomaceous earth or sodium silicate.

[0032] In one embodiment, the wet-process phosphoric acid in step C has a mass concentration of 20% to 48% in terms of P2O5;

[0033] The mass concentration of fluorine in the wet-process phosphoric acid before defluorination in step C is 0.1 wt.% to 5.0 wt.%, preferably 0.1 wt.% to 1.5 wt.%; the extraction time in step C is preferably 10 to 90 minutes, and the volume ratio of the extractant to the wet-process phosphoric acid raw material is preferably 1 to 2; more preferably, the volume ratio of the extractant to the wet-process phosphoric acid raw material is 1.5 to 2.

[0034] In a specific embodiment, the fluorine extraction rate of the extraction in step C is above 60%, and the phosphoric acid extraction rate is below 30%; preferably, the fluorine extraction rate is above 82%.

[0035] In a specific embodiment, the volume ratio of the organic phase in step C to the stripping agent is 2:1 to 1:5, the stripping temperature is 10 to 70° C., and the inorganic base is preferably KOH or NaOH;

[0036] Step C also includes fluorine resource recovery:

[0037] The fluorosilicate solution of step C is subjected to solid-liquid separation at 70-90° C. to obtain a fluorosilicate ointment and a filtrate, the fluorosilicate ointment is dried to obtain a solid product fluorosilicate, and the solid product fluorosilicate is pyrolyzed at 200-800° C. to generate SiF4 and a fluoride salt, or acidolyzed at 95-105% sulfuric acid to generate SiF4 and HF;

[0038] Preferably, the pyrolysis is carried out under a nitrogen atmosphere, the pyrolysis temperature of fluorosilicate K2SiF6 is 400-600°C, and the pyrolysis temperature of Na2SiF6 is 300-800°C;

[0039] Phosphoric acid recovery: the filtrate is cooled and crystallized, and solid-liquid separation is performed to obtain phosphate solid and mother liquor, and the mother liquor is recycled for the stripping step.

[0040] In a specific embodiment, the method further comprises the step D. preparing phosphate using the defluorinated purified phosphoric acid;

[0041] Preferably, the step D comprises: neutralizing the defluorinated purified phosphoric acid with a neutralizing agent, separating the solid and liquid, washing and drying to obtain phosphate, wherein the neutralizing agent is at least one of calcium oxide, calcium hydroxide, and calcium carbonate, and controlling the pH range to be 2.0 to 6.0 to obtain calcium dihydrogen phosphate and / or calcium phosphate;

[0042] Or step D adopts steps D to G of the invention patent with publication number CN 118419875 A to prepare phosphate.

[0043] Beneficial effects:

[0044] 1. High-value utilization of resources: The raw materials of the present invention include phosphate rock (grade range 5%-36%) or white fertilizer (industrial MAP white fertilizer, calcium feed by-product white fertilizer), slag acid, etc., and acid hydrolysis is carried out through a series of treatments to finally obtain phosphate products. This realizes the utilization of associated resources in medium- and low-grade phosphate rock, improves the utilization value of resources, reduces dependence on high-grade phosphate rock, and helps to alleviate the pressure on phosphate rock resources.

[0045] 2. The wet processing method of phosphate rock of the present invention can fully remove impurities, with a low loss rate of phosphoric acid. The purity of the final phosphate product reaches more than 97%, and the product quality is high, which can meet the demand for high-quality phosphate in different fields.

[0046] 3. The process of the present invention is simple, has low energy consumption and low cost. The temperature and time required for extraction are low, and there is no concentration step after extraction, which achieves energy saving, carbon reduction and consumption reduction. At the same time, through reasonable process design, the recycling of materials is achieved, further reducing production costs.

[0047] 4. Coupling and coordinated development with the fluorine and silicon industries: Producing three types of high-value products, namely SiF4, HF and phosphate, to achieve the recycling and utilization of phosphorus-fluorine-silicon resources, thereby improving the economic benefits and resource utilization efficiency of the entire industrial chain.

[0048] 5. Short process flow and high sample diversity: Compared with traditional phosphate production methods, the process flow of the present invention is significantly shortened, reducing complex operations and equipment investment in the production process, and lowering production risks. Furthermore, by selecting different stripping agents, different phosphate products can be obtained, providing high sample diversity and adapting to changing market demand for different phosphate products. BRIEF DESCRIPTION OF THE DRAWINGS

[0049] Figure 1 This is a process flow chart of a specific embodiment of the wet-process phosphoric acid defluorination in step C of the present invention.

[0050] Figure 2 The present invention is a process flow chart of a specific embodiment of the novel wet-process phosphoric acid processing technology. DETAILED DESCRIPTION

[0051] To achieve the purpose of the present invention, the method for wet processing of phosphate rock comprises:

[0052] A. Acid leaching of phosphate rock: The phosphate rock raw material is acid-lyzed and leached, and the solid-liquid separation is carried out to obtain wet-process phosphoric acid;

[0053] B. Removing metal ions from wet-process phosphoric acid;

[0054] C. Wet-process phosphoric acid defluorination: using a wet-process phosphoric acid defluorination extractant to extract fluorosilicic acid in wet-process phosphoric acid at 10-70° C., and separating the phases to obtain defluorinated purified phosphoric acid and an organic phase loaded with fluorosilicate; the wet-process phosphoric acid defluorination extractant is a compound of an organic amine, a cosolvent, and a diluent in a volume ratio of 10-100:0-30:0-90, and the content of the cosolvent and the diluent is not 0; the organic amine includes a primary amine RNH2, a secondary amine R2NH, a tertiary amine R3N, and a quaternary ammonium salt R4N + OH -At least one of, R is a hydrocarbon group; the cosolvent includes at least one of n-butanol, n-hexanol, cyclohexanol, 2-ethyl-1-hexanol, and n-octanol; the diluent includes at least one of sulfonated kerosene, n-dodecane, isomeric dodecane, n-tetradecane, n-hexadecane, and 260# solvent oil.

[0055] The phosphate rock raw material in step A comprises at least one of phosphate rock, white fertilizer, and slag acid, preferably, the phosphate rock raw material contains 5wt% to 35wt% of P2O5;

[0056] Preferably, the acid used for the acidolysis is sulfuric acid or dilute phosphoric acid, and the acid concentration of the acidolysis is ≥5%; the solid-liquid ratio of the acidolysis is 0.1-15, the reaction temperature is 0-90° C., the pH value of the reaction material is controlled within 0.5-1.5, and the stirring reaction is carried out for 0.5-12 hours.

[0057] The method for removing metal ions from wet-process phosphoric acid in step B is extraction or ion exchange;

[0058] Preferably, a decalcification pretreatment is performed before the extraction and removal of metal ions, the amount of the decalcifying agent is 0.8 to 1.2 times the theoretical amount, the reaction temperature is 40 to 60° C., the reaction time is 20 to 40 minutes, the aging time is 0.5 to 2 hours, and the extraction and removal of metal ions are performed after solid-liquid separation; the decalcifying agent is preferably at least one of sulfuric acid, ammonium sulfate, sodium sulfate, and potassium sulfate;

[0059] The method for extracting and removing metal ions preferably adopts the method disclosed in the invention patent publication number CN119265410A;

[0060] Or the method for extracting and removing metal ions preferably adopts the method disclosed in the invention patent publication number CN119503888A, wherein the temperature of step (3) of CN119503888A is 25-85°C, and the volume ratio of O / A is 1:5-10:1; step (4) adopts method 3, and in method 3 of step (4), ammonia is passed through to adjust the pH to 7-12; the temperature of ammonia evaporation is 30-80°C;

[0061] Alternatively, a mixed solution of oxalic acid and ammonium oxalate is used as a stripping agent instead of the stripping agent in step (3) of CN119503888A, wherein the ammonium oxalate concentration in step (3) is 1 wt.% to 20 wt.%, the oxalic acid concentration is 0 wt.% to 20 wt.%, the temperature is 25 to 85°C, and the volume ratio O / A is 1:5 to 10:1; step (4) adopts method 3, wherein ammonia is passed through method 3 to adjust the pH to 7 to 12; and the temperature of ammonia evaporation is 30 to 80°C;

[0062] The method for removing metal ions by ion exchange preferably adopts the method disclosed in the invention patent publication number CN116786096A.

[0063] In a specific embodiment, the wet-process phosphoric acid defluorination extractant is prepared by mixing an organic amine, a co-solvent, and a diluent in a volume ratio of 10-15:10-20:65-75.

[0064] In a specific embodiment, the step C further comprises:

[0065] Stripping: reacting the fluorosilicate-loaded organic phase with a stripping agent, separating the phases to obtain a fluorosilicate solution and a regenerated organic phase, wherein the stripping agent is an inorganic alkali aqueous solution or an inorganic salt aqueous solution;

[0066] Extractant regeneration: adding ammonia water to the regenerated organic phase to obtain an ammonium salt solution and a regenerated extractant, or heating the regenerated organic phase to obtain an aqueous solution and a regenerated extractant. The regenerated extractant is returned to step C for recycling, and the aqueous solution is returned to prepare a stripping agent.

[0067] The inorganic salt aqueous solution for stripping can be at least one of soluble potassium salts K2SO4, KCl, KNO3, and sodium salt solutions Na2SO4, NaCl, and NaCl.

[0068] In a specific embodiment, wet-process phosphoric acid is pretreated before extraction in step C: a silicon replenisher is added to the wet-process phosphoric acid, the silicon replenisher amount being 1.1 to 1.2 times the theoretical value, reacted at 10 to 90° C. for 20 to 40 minutes, aged for 1 hour, and filtered to obtain the pretreated wet-process phosphoric acid raw material;

[0069] The silicon supplementing agent is preferably silicon dioxide, diatomaceous earth or sodium silicate.

[0070] The theoretical value of silicon replenishment is that the F:Si ratio in phosphoric acid is 1:6, that is, the fluorine is completely converted into fluorosilicic acid.

[0071] In one embodiment, the wet-process phosphoric acid in step C has a mass concentration of 20% to 48% in terms of P2O5;

[0072] The mass concentration of fluorine in the wet-process phosphoric acid before defluorination in step C is 0.1 wt.% to 5.0 wt.%, preferably 0.1 wt.% to 1.5 wt.%; the extraction time in step C is preferably 10 to 90 minutes, and the volume ratio of the extractant to the wet-process phosphoric acid raw material is preferably 1 to 2; more preferably, the volume ratio of the extractant to the wet-process phosphoric acid raw material is 1.5 to 2.

[0073] In a specific embodiment, the fluorine extraction rate of the extraction in step C is above 60%, and the phosphoric acid extraction rate is below 30%; preferably, the fluorine extraction rate is above 82%.

[0074] In a specific embodiment, the volume ratio of the organic phase in step C to the stripping agent is 2:1 to 1:5, the stripping temperature is 10 to 70° C., and the inorganic base is preferably KOH or NaOH;

[0075] Step C also includes fluorine resource recovery:

[0076] The fluorosilicate solution of step C is subjected to solid-liquid separation at 70-90° C. to obtain a fluorosilicate ointment and a filtrate, the fluorosilicate ointment is dried to obtain a solid product fluorosilicate, and the solid product fluorosilicate is pyrolyzed at 200-800° C. to generate SiF4 and a fluoride salt, or acidolyzed at 95-105% sulfuric acid to generate SiF4 and HF;

[0077] Preferably, the pyrolysis is carried out under a nitrogen atmosphere, the pyrolysis temperature of fluorosilicate K2SiF6 is 400-600°C, and the pyrolysis temperature of Na2SiF6 is 300-800°C;

[0078] Phosphoric acid recovery: the filtrate is cooled and crystallized, and solid-liquid separation is performed to obtain phosphate solid and mother liquor, and the mother liquor is recycled for the stripping step.

[0079] Solid-liquid separation can be achieved by plate and frame filter press, belt vacuum filtration, fully automatic vertical filter press, etc. After filtration, the filter residue obtained is K2SiF6 and Na2SiF6 fluorosilicates, depending on the alkali and salt used.

[0080] The pyrolysis is kept at high temperature for a period of time to remove the residual organic solvent, and the silicon tetrafluoride gas obtained after the pyrolysis is separated and purified, and the fluorides KF and NaF are recovered and sold.

[0081] In a specific embodiment, the process further includes purifying the silicon tetrafluoride gas obtained after pyrolysis, wherein the method for purifying the silicon tetrafluoride gas includes at least one of condensation, dust removal, and purification, and finally separating and recovering the fluoride products KF and NaF.

[0082] In a specific embodiment, the acidolysis includes high-temperature acidolysis or low-temperature acidolysis, the fluorosilicate solid for the high-temperature acidolysis is at least one of K2SiF6 and Na2SiF6, the temperature for the high-temperature acidolysis is 19°C to 450°C, and the reaction time for the high-temperature acidolysis is preferably 30 to 90 minutes; the fluorosilicate solid for the low-temperature acidolysis is at least one of K2SiF6 and Na2SiF6, the temperature for the low-temperature acidolysis is less than 19°C, and the reaction time for the low-temperature acidolysis is preferably 30 to 180 minutes; the sulfuric acid concentration of the acidolysis is preferably 95 to 105 wt%.

[0083] The sulfuric acid used in the present invention is fuming sulfuric acid, and its concentration is based on the sulfate SO4 2- Determined, because there is gaseous sulfate ion SO4 in fuming sulfuric acid 2-, so there is a case where it may be greater than 100wt%.

[0084] In a specific embodiment, the separation method includes: performing high-temperature acid hydrolysis to obtain a mixed gas of anhydrous hydrogen fluoride and silicon tetrafluoride, and performing rectification to obtain high-purity anhydrous hydrogen fluoride and silicon tetrafluoride respectively; performing low-temperature acid hydrolysis to directly collect anhydrous silicon tetrafluoride gas, and then heating the liquid after low-temperature acid hydrolysis to 20°C to 60°C and reacting for 20 to 55 minutes to separate the anhydrous hydrogen fluoride gas;

[0085] Preferably, the separated sulfate is returned to the recycling stripping process;

[0086] The methods used for the rectification and purification include but are not limited to adsorption method, freezing method, fluorine gas method and fluorinating agent method.

[0087] In a specific embodiment, the method further comprises the step D. preparing phosphate using the defluorinated purified phosphoric acid;

[0088] Preferably, the step D comprises: neutralizing the defluorinated purified phosphoric acid with a neutralizing agent, separating the solid and liquid, washing and drying to obtain phosphate, wherein the neutralizing agent is at least one of calcium oxide, calcium hydroxide, and calcium carbonate, and controlling the pH range to be 2.0 to 6.0 to obtain calcium dihydrogen phosphate and / or calcium phosphate;

[0089] Or step D adopts steps D to G of the invention patent with publication number CN 118419875 A to prepare phosphate.

[0090] The specific embodiments of the present invention are further described below in conjunction with examples, but the present invention is not limited to the scope of the examples.

[0091] Example 1

[0092] Wet-process phosphoric acid is produced from low-grade phosphate rock with a 10% P2O5 content. Phosphate rock, 98% sulfuric acid, and 20% dilute phosphoric acid are added to a decomposition tank in a specific ratio. The solid-liquid ratio is set at 5:1, the reaction temperature is controlled at 80°C, and the pH of the reaction mixture is adjusted to 1.0. The agitator is turned on and the reaction is stirred for 6 hours. The dilute acid concentration is maintained above 10%. After the reaction is completed, the phosphoric acid slurry is passed through a filtration device, where it is filtered to produce wet-process phosphoric acid with a 25% P2O5 content and phosphogypsum.

[0093] Before the metal ions are removed by solvent extraction, the wet-process phosphoric acid is pretreated with decalcification. The decalcification agent is ammonium sulfate, the amount of the decalcification agent is 1.2 times the theoretical amount, the reaction temperature is 60° C., the reaction time is 40 minutes, the aging time is 1 hour, and the pretreated phosphoric acid is obtained after filtration.

[0094] Then extract wet-process phosphoric acid to remove metal ions:

[0095] (1) Extraction: Wet-process phosphoric acid was subjected to three-stage three-stage countercurrent extraction. The extraction temperature of each stage was 60°C, and the O / A ratio was 4:1 (volume ratio).

[0096] (2) Washing: The metal ion-loaded extractant is first subjected to two-stage countercurrent washing in dilute phosphoric acid with a concentration of 1 wt.% P2O5 at a temperature of 50°C and an O / A ratio of 3:1 (volume ratio) to recover the phosphoric acid entrained in the extractant and improve the P2O5 yield in the phosphoric acid purification process.

[0097] (3) Stripping 1: Using 35 wt.% ammonium sulfate solution and 1 wt.% sulfuric acid solution as stripping agent 1, the metal ion-loaded extractant is subjected to two-stage three-stage countercurrent stripping. The stripping temperature of each stage is 30°C, and the O / A ratio is 2:1 (volume ratio), to obtain stripping solution 1 and organic phase 1. Stripping solution 1 is evaporated and concentrated, filtered to obtain calcium sulfate; the filtrate is continuously cooled to 80°C to obtain ammonium manganese sulfate; cooled to 50°C to obtain ammonium magnesium sulfate; cooled to 10°C to obtain mixed crystals of potassium magnesium sulfate and sodium magnesium sulfate; this step classifies and recovers calcium, manganese, magnesium, potassium, and sodium ions.

[0098] (4) Washing 1: The organic phase 1 was washed with a dilute ammonium sulfate solution having a concentration of 0.1 wt.% in two stages in countercurrent at a temperature of 30° C. and a phase O / A ratio of 3:1 (volume ratio) to recover the ammonium sulfate entrained in the organic phase 1.

[0099] (5) Stripping 2: Using a sulfuric acid solution with a concentration of 30 wt.% as stripping agent 2, stripping the metal ions in the organic phase 1 at a temperature of 30°C and a phase O / A ratio of 1:1 (volume ratio) to obtain stripping solution 2 and organic phase 2; adding ammonium sulfate solid to the stripping solution 2, cooling it to 80°C to obtain cesium aluminum sulfate; cooling it to 55°C to obtain potassium aluminum sulfate; cooling it to 20°C to obtain ammonium aluminum sulfate; the crystallization mother liquor is recycled as stripping agent 2.

[0100] (6) Washing 2: The organic phase 2 is washed with dilute sulfuric acid having a concentration of 5 wt.% in two-stage countercurrent at a temperature of 30°C and an O / A ratio of 3:1 (volume ratio). The sulfuric acid entrained in the organic phase 2 is recovered and returned to the stripping agent 1 for recycling.

[0101] (7) Stripping 3: Using 10 wt.% oxalic acid solution as stripping agent 3, the organic phase 2 is subjected to two-stage countercurrent stripping at a temperature of 50°C and an O / A ratio of 1:1 (volume ratio) to obtain stripping solution 3 and organic phase 3. The stripping solution 3 is treated with ultraviolet light to obtain ferrous oxalate precipitate.

[0102] (8) Washing 3: The organic phase 3 is washed in two stages in countercurrent with dilute oxalic acid having a concentration of 1 wt.% at a temperature of 50°C and an O / A ratio of 3:1 (volume ratio). The oxalic acid entrained in the organic phase 3 is recovered to obtain a regenerated extractant that can be recycled.

[0103] The extractant is returned to the wet-process phosphoric acid for recycling.

[0104] After extraction, the metal ion content in phosphoric acid dropped to 0.58 wt.% Al2O3, 0.92 wt.% MgO, 0.21 wt.% Fe2O3, and 0.67 wt.% CaO, with a MER value of 0.095. (The metal ion content is mainly determined by the MER value, which is the ratio of the sum of the metal ion content to the phosphorus pentoxide content. Industrially, this value is below 0.1, and the metal ions have little impact on subsequent processes.) The P2O5 content was 25 wt.%, and the F content was 0.46 wt.%.

[0105] Silicon dioxide was added to wet-process phosphoric acid in an amount of 1.2 times the theoretical value. The mixture was reacted at 40° C. for 40 minutes and aged for 1 hour. The pretreated wet-process phosphoric acid was obtained after filtration.

[0106] Extraction is then carried out at 40°C. A 2:1 volume ratio of extractant and wet-process phosphoric acid (F concentration: 0.46 wt% and P2O5 concentration: 25 wt%) are added to the reactor. The extractant composition is trioctylamine (TOA): n-octanol: sulfonated kerosene (volume ratio: 0.15:0.20:0.65). After 60 minutes, the phases are separated to produce aqueous phase 1 and oil phase 1, which is the purified wet-process phosphoric acid. KOH aqueous solution is added to oil phase 1 for back extraction. The volume ratio of KOH aqueous solution to organic phase is 1:2. The reaction is allowed to proceed at 30°C for 40 minutes, followed by stratification and separation to produce aqueous phase 2 and oil phase 2. Aqueous phase 2 is heated and filtered at 70°C. The filtrate is a potassium dihydrogen phosphate solution. The filter cake is dried at 120°C for 6 hours to produce potassium fluorosilicate solid with a moisture content of less than 0.1% after drying. The filtrate is cooled, crystallized, and filtered to produce potassium dihydrogen phosphate crystals. The filtered mother liquor can be dissolved with KOH and returned to the back extraction stage. The oil phase is heated to 50°C for 40 minutes, allowed to stand for stratification, and then separated to obtain oil phase 3, trioctylamine, which is recovered and recycled for extraction. The aqueous phase 3 is water, which can dissolve KOH and return to the stripping stage. The dried filter cake, K2SiF6, is ground and calcined at 500°C under a nitrogen atmosphere for 2 hours. The SiF4 gas is collected using an air bag. After the reaction is complete, the solid KF after pyrolysis is collected and recovered as a byproduct. The SiF4 gas is purified by liquid nitrogen cryogenic distillation and further recovered. The final gas purity is >95%. In this example, the final fluorine extraction rate reached 82.55%, the phosphoric acid extraction rate was 22.38%, and the P2O5 content after extraction was 19.41wt.%, and the F content was 0.080wt.%.

[0107] Finally, lithium dihydrogen phosphate is prepared using purified wet-process phosphoric acid:

[0108] (1) Extraction: Methyl isobutyl ketone was used as the extractant to extract and purify wet-process phosphoric acid. The extraction temperature was 60°C, the volume ratio of the composite extractant to the purified wet-process phosphoric acid was 1:3, and the number of extraction stages was 3.

[0109] (2) Phase separation 1: The mixed solution of the composite extractant and wet-process phosphoric acid is allowed to stand for phase separation to obtain oil phase 1 and water phase 1. The organic oil phase of the coupled phosphoric acid is the extract, and the remaining water phase is water and a small amount of phosphoric acid. The water phase is used as the washing water for phosphogypsum. The washing water is returned to the phosphoric acid leaching tank. The phase separation temperature is 60°C and the phase separation time is 3 minutes.

[0110] (3) Stripping: The oil phase 1 obtained by phase separation 1 is stripped with lithium sulfate at a stripping temperature of 60° C. The volume ratio of the stripping agent to the organic phase is 1:1, and the molar ratio of lithium ions to phosphoric acid carried in the extractant is 1:1.05;

[0111] (4) Phase separation 2: The mixed solution of the extract and the stripping solution was allowed to stand for phase separation to obtain oil phase 2 and water phase 2. The main components of water phase 2 are: H2PO 4- 、Li + , H2O and a very small amount of SO4 2- The main component of oil phase 2 is the composite extractant carrying H2SO4, the separation temperature is 60℃, and the separation time is 3min;

[0112] (5) Stripping 2: Ammonia water was added to the oil phase 2 obtained from phase separation 2 for secondary stripping. The secondary stripping temperature was 50°C, the volume ratio of the stripping agent to the organic phase was 1:1, and the molar ratio of the stripping agent NH3 to the acid carried in the extractant was 1.1:1.

[0113] (6) Phase separation 3: The material obtained from stripping 2 is subjected to phase separation to obtain oil phase 3 and water phase 3. The oil phase of the composite extractant after stripping is returned to the extraction process, and the water phase is subjected to subsequent operations;

[0114] (7) Evaporation and concentration 1: The aqueous ammonium sulfate solution obtained in phase separation 3 was evaporated and concentrated to a specific gravity of 1.35 at a concentration temperature of 80°C and a pressure of 30 kPa;

[0115] (8) Cooling crystallization 2: The concentrated liquid is cooled and crystallized, and the cooling end temperature is 40°C;

[0116] (9) Filtration: The product obtained by cooling crystallization is filtered to obtain aqueous phase 4 and solid ammonium salt (NH4)2SO4. The mother liquor of aqueous phase 4 liquid crystallization is returned to (7) evaporation concentration stage 1;

[0117] (10) Evaporation and concentration 2: The aqueous phase 4 obtained in (9) was evaporated and concentrated to a specific gravity of 1.45 at a concentration temperature of 80°C and a pressure of 20 kPa;

[0118] (11) Cooling crystallization 3: The concentrated liquid is cooled and crystallized to obtain lithium dihydrogen phosphate with a purity of 98.5% at a crystallization temperature of 30°C.

[0119] Example 2

[0120] Wet-process phosphoric acid is produced using medium-grade phosphate rock (P2O5 content of 20%). Phosphate rock, sulfuric acid (98% concentration), and dilute phosphoric acid (20% concentration) are added to a decomposition tank in a specific ratio. The solid-to-liquid ratio is set at 5:1, the reaction temperature is controlled at 70°C, and the pH of the reaction mixture is adjusted to 1.2. The agitator is turned on and the reaction is stirred for 4 hours. The dilute acid concentration is maintained above 10%. After the reaction is completed, the phosphoric acid slurry is passed through a filtration device, where it is filtered to produce wet-process phosphoric acid with a P2O5 content of 30% and phosphogypsum.

[0121] Before the metal ions are removed by solvent extraction, the wet-process phosphoric acid is pretreated with decalcification. The amount of the decalcifying agent ammonium sulfate is 1.2 times the theoretical amount. The reaction temperature is 60° C., the reaction time is 40 minutes, the aging time is 1 hour, and the pretreated phosphoric acid is obtained after filtration.

[0122] Then extract the wet-process phosphoric acid to remove metal ions:

[0123] (1) Extraction: Wet-process phosphoric acid was subjected to three-stage three-stage countercurrent extraction. The extraction temperature of each stage was 60°C, and the O / A ratio was 4:1 (volume ratio).

[0124] (2) Washing: The metal ion-loaded extractant is first subjected to two-stage countercurrent washing in dilute phosphoric acid with a concentration of 2 wt.% P2O5 at a temperature of 55°C and an O / A ratio of 4:1 (volume ratio) to recover the phosphoric acid entrained in the extractant and improve the P2O5 yield in the phosphoric acid purification process.

[0125] (3) Stripping 1: Using 35 wt.% ammonium sulfate solution and 3 wt.% sulfuric acid solution as stripping agent 1, the metal ion-loaded extractant is subjected to two-stage three-stage countercurrent stripping. The stripping temperature of each stage is 30°C, and the O / A ratio is 3:1 (volume ratio), to obtain stripping solution 1 and organic phase 1. Stripping solution 1 is evaporated and concentrated, filtered to obtain calcium sulfate; the filtrate is continuously cooled to 75°C to obtain ammonium manganese sulfate; cooled to 45°C to obtain ammonium magnesium sulfate; cooled to 8°C to obtain mixed crystals of potassium magnesium sulfate and sodium magnesium sulfate; this step classifies and recovers calcium, manganese, magnesium, potassium, and sodium ions.

[0126] (4) Washing 1: The organic phase 1 was washed with a dilute ammonium sulfate solution having a concentration of 1 wt.% in two stages in countercurrent at a temperature of 35° C. and a phase O / A ratio of 4:1 (volume ratio) to recover the ammonium sulfate entrained in the organic phase 1.

[0127] 5) Stripping 2: Using a 35 wt.% sulfuric acid solution as stripping agent 2, stripping the metal ions in the organic phase 1 at a temperature of 25° C. and a phase O / A ratio of 2:1 (volume ratio) to obtain a stripping solution 2 and an organic phase 2; adding ammonium sulfate solid to the stripping solution 2, cooling the solution to 75° C. to obtain cesium aluminum sulfate; cooling the solution to 50° C. to obtain potassium aluminum sulfate; and cooling the solution to 15° C. to obtain ammonium aluminum sulfate; and recycling the crystallization mother liquor as stripping agent 2.

[0128] (6) Washing 2: The organic phase 2 is washed with dilute sulfuric acid having a concentration of 6 wt.% in two-stage countercurrent at a temperature of 25°C and an O / A ratio of 4:1 (volume ratio). The sulfuric acid entrained in the organic phase 2 is recovered and returned to the stripping agent 1 for recycling.

[0129] (7) Stripping 3: Using oxalic acid solution with a concentration of 12 wt.% as stripping agent 3, the organic phase 2 is subjected to two-stage countercurrent stripping at a temperature of 55°C and a phase O / A ratio of 2:1 (volume ratio) to obtain stripping solution 3 and organic phase 3. The stripping solution 3 is treated with ultraviolet light to obtain ferrous oxalate precipitate.

[0130] (8) Washing 3: The organic phase 3 is washed in two-stage countercurrent with dilute oxalic acid having a concentration of 0.1 wt.% at a temperature of 55°C and an O / A ratio of 4:1 (volume ratio). The oxalic acid entrained in the organic phase 3 is recovered to obtain a regenerated extractant that can be recycled.

[0131] After extraction, the metal ion content in phosphoric acid dropped to 0.73 wt.% Al2O3, 0.78 wt.% MgO, 0.22 wt.% Fe2O3, 0.62 wt.% CaO, with a MER value of 0.078. The P2O5 content was 30 wt.% and the F content was 0.65 wt.%.

[0132] Silica was added to wet-process phosphoric acid at an amount 1.2 times the theoretical value, reacted at 40°C for 40 minutes, aged for 1 hour, and filtered to obtain pretreated wet-process phosphoric acid. The purified phosphoric acid, from which metal ions had been removed, was subjected to extractive defluorination.

[0133] Extraction was performed at 40°C. A separatory funnel was charged with an extractant (volume ratio of 1.5:1) and wet-process phosphoric acid (F concentration of 0.65 wt% and P2O5 concentration of 30 wt%). The extractant composition (volume ratio of TOA: n-octanol: sulfonated kerosene = 0.15:0.20:0.65). After 60 minutes, the extractant was separated to produce aqueous phase 1 and oil phase 1, which is the purified wet-process phosphoric acid. NaOH solution was added to oil phase 1 for back extraction. The volume ratio of NaOH solution to organic phase was 1:2. The reaction was allowed to stand at 30°C for 40 minutes, followed by separation to produce aqueous phase 2 and oil phase 2. Aqueous phase 2 was heated and filtered at 70°C. The filtrate was a sodium dihydrogen phosphate solution. The filter cake was dried at 120°C for 6 hours to produce sodium fluorosilicate solid with a moisture content of less than 0.1% after drying. The filtrate was cooled, crystallized, and filtered to produce sodium dihydrogen phosphate crystals. The filtered mother liquor was then dissolved with NaOH and returned to the back extraction process. The oil phase is heated to 50°C for 2 minutes, allowed to stand for 40 minutes, and then separated. The resulting oil phase 3 is trioctylamine, which is recovered and recycled for extraction. The aqueous phase 3 is water, which can dissolve NaOH and then return to the stripping stage. The dried filter cake, fluorosilicate Na2SiF6, is ground and calcined at 500°C under a nitrogen atmosphere for 2 hours. SiF4 gas is collected using an air bag. After the reaction is complete, the solid NaF after pyrolysis is collected as a byproduct. The SiF4 gas is purified by liquid nitrogen cryogenic distillation and further recovered. The final gas purity is >95%. In this example, the final fluorine extraction rate reached 86.30%, and the phosphoric acid extraction rate was 25.73%.

[0134] Calcium oxide is added as a neutralizing agent to the purified wet-process phosphoric acid (Al2O3 content: 0.73 wt.%, MgO content: 0.78 wt.%, Fe2O3 content: 0.22 wt.%, CaO content: 0.62 wt.%, P2O5 content: 22.28 wt.%, and F content: 0.089 wt.%) after metal and fluoride ions have been removed. This is added to the reactor at a mass ratio of 1:1.3 (solution:mixture). The reaction temperature is maintained at 75°C and the pH is controlled at 3.5 to produce a slurry. A filter press is used to separate the slurry into a solid and a mother liquor, yielding a DCP-rich solid and a mother liquor. The solid is washed and dried to obtain the final product. The mother liquor is returned to the washing section of the phosphate rock decomposition process to produce wet-process phosphoric acid and used to wash filtered phosphogypsum, achieving resource recycling.

[0135] Example 3

[0136] Wet-process phosphoric acid is produced from low-grade phosphate rock (10% P2O5 content). Phosphate rock, sulfuric acid (98% concentration), and dilute phosphoric acid (20% concentration) are added to a decomposition tank in a specific ratio. The solid-to-liquid ratio is set at 5:1, the reaction temperature is controlled at 80°C, and the pH of the reaction mixture is adjusted to 1.0. The agitator is activated and the reaction is stirred for 6 hours. The dilute acid concentration is maintained above 10%. After the phosphate rock acid leachate is heated to 70°C, the adsorption assembly is controlled to descend until the phosphate rock acid leachate submerges the cation exchange membrane for 600 seconds. The adsorbed adsorption assembly is then immersed in a desorption tank containing a desorption solution of 30% sulfuric acid for 600 seconds. After the reaction, the phosphoric acid slurry is filtered to produce wet-process phosphoric acid with a 25% P2O5 content and phosphogypsum.

[0137] The adsorption process for purifying phosphate rock acid hydrolyzate by cation exchange membrane impregnation in this embodiment includes the following steps:

[0138] (1) Preparation of adsorption assembly: The cation exchange membrane loaded with sulfonic acid adsorption groups is installed and fixed in the adsorption assembly.

[0139] (2) Adsorption and desorption treatment of cationic impurities:

[0140] ① Adsorption: After heating the phosphate rock acid leachate to 70°C, control the adsorption component to descend until the phosphate rock acid leachate submerges the cation exchange membrane, with a residence time of 600 seconds;

[0141] ② Post-adsorption washing: lift the adsorption component and immerse it in washing tank 1; the residence time is 60 seconds; the washing tank 1 contains pure water;

[0142] ③ Desorption: The washed adsorption component is immersed in a desorption tank containing a desorption liquid for 600 seconds; the desorption liquid is 30% sulfuric acid.

[0143] ④ Post-desorption washing: After desorption, the adsorption separation membrane assembly is immersed in a washing tank 2 for 60 seconds and then taken out; the washing tank 2 contains 5% sulfuric acid;

[0144] (3) After the phosphate rock acid leachate crystallization was completed, solid-liquid separation was performed and the impurity contents in the liquid phase were measured. The results were: Al2O3 content of 0.48 wt.%, MgO content of 1.17 wt.%, Fe2O3 content of 0.21 wt.%, CaO content of 0.54 wt.%, and MER value of 0.096. P2O5 content of 25 wt.%, and F content of 0.46 wt.%.

[0145] Silicon dioxide was added to wet-process phosphoric acid in an amount of 1.2 times the theoretical value. The mixture was reacted at 40° C. for 40 minutes and aged for 1 hour. The pretreated wet-process phosphoric acid was obtained after filtration.

[0146] Extraction is then carried out at 40°C. A separatory funnel is charged with an extractant (volume ratio of 2:1) and wet-process phosphoric acid (F concentration of 0.46 wt% and P2O5 concentration of 25 wt%). The extractant composition is trioctylamine (TOA): n-octanol: sulfonated kerosene (volume ratio = 0.15:0.20:0.65). After 60 minutes, the phases separate to produce aqueous phase 1 and oil phase 1, which is the purified wet-process phosphoric acid. KOH aqueous solution is added to oil phase 1 for back extraction. The volume ratio of KOH aqueous solution to organic phase is 1:2. The reaction is allowed to stand at 30°C for 40 minutes, followed by separation to produce aqueous phase 2 and oil phase 2. Aqueous phase 2 is heated and filtered at 70°C. The filtrate is a potassium dihydrogen phosphate solution. The filter cake is dried at 120°C for 6 hours to produce potassium fluorosilicate solid with a moisture content of less than 0.1% after drying. The filtrate is cooled, crystallized, and filtered to produce potassium dihydrogen phosphate crystals. The filtered mother liquor can be dissolved with KOH and returned to the back extraction stage. The oil phase is heated to 50°C for 40 minutes, allowed to stand for stratification, and then separated. The resulting oil phase 3 is trioctylamine, which is recovered for re-extraction. The aqueous phase 3 is water, which can dissolve KOH and return to the stripping stage. The dried filter cake, fluorosilicate K2SiF6, is ground and calcined at 500°C under a nitrogen atmosphere for 2 hours. SiF4 gas is collected using an air bag. After the reaction is complete, the pyrolyzed solid KF is collected as a byproduct. The SiF4 gas is purified by cryogenic distillation with liquid nitrogen and further recovered. The final gas purity is >95%. In this case, the final fluorine extraction rate reached 82.72%, and the phosphoric acid extraction rate was 23.15%.

[0147] The metal ions and fluoride ions were removed to purify the wet-process phosphoric acid (MgO content was 1.17wt.%, Al2O3 content was 0.48wt.%, Fe2O3 content was 0.21wt.%, CaO content was 0.54wt.%, P2O5 content was 19.21wt.%, and F content was 0.079wt.%).

[0148] Finally, purified phosphoric acid is used to produce sodium dihydrogen phosphate:

[0149] (1) Extraction: Using di(2-ethylhexyl)hexyl phosphate as the extractant to extract and purify wet-process phosphoric acid, the extraction temperature is 60°C, the volume ratio of the composite extractant to the purified wet-process phosphoric acid is 1:3, and the number of extraction stages is 3;

[0150] (2) Phase separation 1: The mixed solution of the composite extractant and wet-process phosphoric acid is allowed to stand for phase separation to obtain oil phase 1 and water phase 1. The organic oil phase of the coupled phosphoric acid is the extract, and the remaining water phase is water and a small amount of phosphoric acid. The water phase is used as the washing water for phosphogypsum. The washing water is returned to the phosphoric acid leaching tank. The phase separation temperature is 60°C and the phase separation time is 3 minutes.

[0151] (3) Stripping 1: stripping the oil phase 1 obtained by phase separation 1 with sodium sulfate, the stripping temperature is 60°C, the volume ratio of the stripping agent to the organic phase is 1:1, and the molar ratio of lithium ions to phosphoric acid carried in the extractant is 1:1.05;

[0152] (4) Phase separation 2: The mixed solution of the extract and the stripping solution is allowed to stand for phase separation to obtain oil phase 2 and water phase 2. The main components of water phase 2 are: H2PO4 - 、Na + , H2O and a very small amount of SO4 2- The main component of oil phase 2 is the composite extractant carrying H2SO4, the separation temperature is 60℃, and the separation time is 3min;

[0153] (5) Stripping 2: Ammonia water was added to the oil phase 2 obtained from phase separation 2 for secondary stripping. The secondary stripping temperature was 50°C, the volume ratio of the stripping agent to the organic phase was 1:1, and the molar ratio of the stripping agent NH3 to the acid carried in the extractant was 1.1:1.

[0154] (6) Phase separation 3: The material obtained from stripping 2 is subjected to phase separation to obtain oil phase 3 and water phase 3. The composite extractant oil phase 3 after stripping is returned to the extraction process, and the water phase 3 is subjected to subsequent operations;

[0155] (7) Evaporation and concentration 1: The aqueous phase ammonium sulfate solution obtained by phase separation 3 is evaporated and concentrated to a specific gravity of 1.35 at a concentration temperature of 80°C and a pressure of 30 kPa;

[0156] (8) Cooling crystallization 2: The concentrated liquid is cooled and crystallized, and the cooling end temperature is 40°C;

[0157] (9) Filtration: The product obtained by cooling crystallization is filtered to obtain aqueous phase 4 and solid ammonium salt (NH4)2SO4. The liquid crystallization mother liquor is returned to the evaporation and concentration stage 1 in (7);

[0158] (10) Evaporation and concentration 2: The aqueous phase 4 obtained in (9) was evaporated and concentrated to a specific gravity of 1.45 at a concentration temperature of 80°C and a pressure of 20 kPa;

[0159] (11) Cooling crystallization 3: The concentrated liquid is subjected to cooling crystallization to obtain sodium dihydrogen phosphate with a purity of 99.2% at a crystallization temperature of 30°C.

[0160] Example 4

[0161] Wet-process phosphoric acid is produced using medium-grade phosphate rock (P2O5 content of 20%). Phosphate rock, sulfuric acid (98% concentration), and dilute phosphoric acid (20% concentration) are added to a decomposition tank in a specific ratio. The solid-to-liquid ratio is set at 5:1, the reaction temperature is controlled at 70°C, and the pH of the reaction mixture is adjusted to 1.2. The agitator is turned on and the reaction is stirred for 4 hours. The dilute acid concentration is maintained above 10%. After the reaction is completed, the phosphoric acid slurry is passed through a filtration device, where it is filtered to produce wet-process phosphoric acid with a P2O5 content of 30% and phosphogypsum.

[0162] In this embodiment, a cation exchange membrane fixed bed method is used to purify phosphate rock acid hydrolyzate, including the following steps:

[0163] (1) Prepare a cation exchange membrane loaded with sulfonic acid adsorption groups.

[0164] (2) Adsorption and desorption treatment of cationic impurities:

[0165] ① Adsorption: The phosphate rock acid leachate flows through the cation exchange membrane at a controlled flow rate of 10 cm / min;

[0166] ② Post-adsorption washing: After adsorption, washing liquid 1 is introduced at a flow rate of 20 cm / min; washing liquid 1 is pure water;

[0167] ③ Desorption: After washing, the desorption liquid is introduced at a flow rate of 10 cm / min; the desorption liquid is 30% sulfuric acid;

[0168] ④ Washing after desorption: introduce washing liquid 2 at a flow rate of 20 cm / min; washing liquid 2 is 5% sulfuric acid;

[0169] (3) After the phosphate rock acid leachate crystallization was completed, solid-liquid separation was performed and the impurity contents in the liquid phase were measured. The results were: Al2O3 content of 0.52 wt.%, MgO content of 1.36 wt.%, Fe2O3 content of 0.38 wt.%, CaO content of 0.46 wt.%, and MER value of 0.091. P2O5 content of 30 wt.%, and F content of 0.65 wt.%.

[0170] Silicon dioxide was added to wet-process phosphoric acid in an amount of 1.2 times the theoretical value. The mixture was reacted at 40° C. for 40 minutes and aged for 1 hour. The pretreated wet-process phosphoric acid was obtained after filtration.

[0171] The purified phosphoric acid, free of metal ions, undergoes extractive defluorination at 25°C. A reaction kettle is charged with an extractant (TOA:n-octanol:sulfonated kerosene) in a volume ratio of 2:1 and wet-process phosphoric acid (0.65 wt% F and 30 wt% P2O5) (total iodine, n-octanol, 0.15:0.2:0.65). After 60 minutes, the extractant is separated to produce an aqueous phase 1 and an oil phase 1, which is the purified wet-process phosphoric acid. NaOH solution is added to the oil phase 1 for back extraction, with a volume ratio of 1:2. The reaction is continued at 40°C for 40 minutes, followed by separation to produce an aqueous phase 2 and an oil phase 2. The aqueous phase 2 is heated and filtered at 70°C. The filtrate is a sodium dihydrogen phosphate solution. The filter cake is dried at 120°C for 6 hours to produce a sodium fluorosilicate solid with a moisture content of less than 0.1% after drying. The filtrate is cooled and crystallized to obtain sodium dihydrogen phosphate crystals. The filtered mother liquor can dissolve NaOH and return to the stripping stage. The oil phase 2 is heated to 50 degrees Celsius, allowed to stand for 40 minutes to separate the layers, and the oil phase 3 is trioctylamine, which is recovered and recycled for extraction. The aqueous phase 3 is water, which can dissolve NaOH and return to the stripping stage. The dried filter cake, fluorosilicate Na2SiF6, is ground and calcined at 500°C under a nitrogen atmosphere for 2 hours. The SiF4 gas is collected in an air bag. After the reaction is complete, the solid NaF after pyrolysis is collected as a by-product. The SiF4 gas is further recovered after purification by liquid nitrogen low-temperature distillation. The final gas purity is >95%. In this example, the final fluorine extraction rate reached 86.95%, and the phosphoric acid extraction rate was 26.16%. The metal ions and fluoride ions were removed to purify the wet-process phosphoric acid (Al2O3 content was 0.52wt.%, MgO content was 1.36wt.%, Fe2O3 content was 0.38wt.%, CaO content was 0.46wt.%, P2O5 content was 22.15wt.%, and F content was 0.085wt.%).

[0172] DCP is then produced by purifying and concentrating phosphoric acid. Calcium oxide is added as a neutralizing agent to the purified wet-process phosphoric acid, which has been free of metal ions and fluoride ions, and then fed into the reactor at a mass ratio of 1:1.3 (solution:mixture). The reaction temperature is maintained at 75°C and the pH is controlled at 3.5 to produce a slurry. A filter press is used to separate the slurry into a solid rich in DCP and a mother liquor. The solid is washed and dried to obtain the final product. The mother liquor is returned to the washing section of the phosphate rock decomposition process to produce wet-process phosphoric acid and used to wash the filtered phosphogypsum, thus achieving resource recycling.

[0173] Example 5

[0174] Wet-process phosphoric acid is produced using a mixture of medium-grade phosphate rock (P2O5 content of 20%) and white fertilizer (P2O5 content of 20%). High-grade phosphate rock, white fertilizer, sulfuric acid (98%), and dilute phosphoric acid (30%) are added to a decomposition tank in a specific ratio. The solid-to-liquid ratio is set at 10:1, the reaction temperature is controlled at 80°C, and the pH of the reaction mixture is adjusted to 0.8. The agitator is turned on and the reaction is stirred for 4 hours. The dilute acid concentration is maintained above 15%. After the reaction is completed, the phosphoric acid slurry is fed into a high-efficiency filtration device, where filtration produces wet-process phosphoric acid with a P2O5 content of 40% and solid phosphogypsum.

[0175] Then extract the wet-process phosphoric acid to remove metal ions:

[0176] (1) Extraction: Wet-process phosphoric acid was subjected to three-stage three-stage countercurrent extraction. The extraction temperature of each stage was 60°C, and the O / A ratio was 4:1 (volume ratio).

[0177] (2) Washing: a. The metal ion-loaded extractant is first washed in two stages of countercurrent with dilute phosphoric acid having a concentration of 3 wt.% P2O5 at a temperature of 60°C and an O / A ratio of 5:1 (volume ratio) to recover the phosphoric acid entrained in the extractant and improve the P2O5 yield in the phosphoric acid purification process.

[0178] (3) Stripping 1: Using 35 wt.% ammonium sulfate solution and 5 wt.% sulfuric acid solution as stripping agent 1, the metal ion-loaded extractant is subjected to two-stage three-stage countercurrent stripping. The stripping temperature of each stage is 35°C, and the O / A ratio is 4:1 (volume ratio), to obtain stripping solution 1 and organic phase 1. Stripping solution 1 is evaporated and concentrated, and filtered to obtain calcium sulfate; the filtrate is continuously cooled to 70°C to obtain ammonium manganese sulfate; the temperature is cooled to 40°C to obtain ammonium magnesium sulfate; the temperature is cooled to 6°C to obtain mixed crystals of potassium magnesium sulfate and sodium magnesium sulfate; this step classifies and recovers calcium, manganese, magnesium, potassium, and sodium ions.

[0179] (4) Washing 1: The organic phase 1 was washed with a 2 wt.% dilute ammonium sulfate solution in two stages in countercurrent at 40°C with an O / A ratio of 5:1 (volume ratio) to recover the ammonium sulfate entrained in the organic phase 1.

[0180] 5) Stripping 2: Using a 40 wt.% sulfuric acid solution as stripping agent 2, stripping the metal ions in the organic phase 1 at a temperature of 35° C. and a phase O / A ratio of 3:1 (volume ratio) to obtain a stripping solution 2 and an organic phase 2; adding ammonium sulfate solid to the stripping solution 2, cooling the solution to 70° C. to obtain cesium aluminum sulfate; cooling the solution to 45° C. to obtain potassium aluminum sulfate; and cooling the solution to 10° C. to obtain ammonium aluminum sulfate; and recycling the crystallization mother liquor as stripping agent 2.

[0181] (6) Washing 2: The organic phase 2 is washed with dilute sulfuric acid having a concentration of 7 wt.% in two-stage countercurrent at a temperature of 35°C and an O / A ratio of 5:1 (volume ratio). The sulfuric acid entrained in the organic phase 2 is recovered and returned to the stripping agent 1 for recycling.

[0182] (7) Stripping 3: Using 14 wt.% oxalic acid solution as stripping agent 3, the organic phase 2 is subjected to two-stage countercurrent stripping at a temperature of 60°C and an O / A ratio of 3:1 (volume ratio) to obtain stripping solution 3 and organic phase 3. The stripping solution 3 is treated with ultraviolet light to obtain ferrous oxalate precipitate.

[0183] (8) Washing 3: The organic phase 3 is washed in two stages in countercurrent with dilute oxalic acid having a concentration of 0.5 wt.% at a temperature of 60°C and an O / A ratio of 5:1 (volume ratio). The oxalic acid entrained in the organic phase 3 is recovered to obtain a regenerated extractant that can be recycled.

[0184] After extraction, the metal ion content in phosphoric acid dropped to 0.64 wt.% Al2O3, 1.15 wt.% MgO, 0.25 wt.% Fe2O3, 0.71 wt.% CaO, with a MER value of 0.069, 40 wt.% P2O5, and 0.84 wt.% F.

[0185] Silicon dioxide was added to wet-process phosphoric acid in an amount of 1.2 times the theoretical value. The mixture was reacted at 40° C. for 40 minutes and aged for 1 hour. The pretreated wet-process phosphoric acid was obtained after filtration.

[0186] The purified phosphoric acid, free of metal ions, was subjected to extractive defluorination. Extraction was performed at 40°C. A 2:1 ratio (volume ratio) of extractant and wet-process phosphoric acid with a 0.84 wt.% F concentration and a 40 wt.% P2O5 concentration was added to a reactor. The extractant composition was TOA: n-octanol: sulfonated kerosene (volume ratio = 0.15:0.20:0.65). After 60 minutes, the extractant was separated to produce aqueous phase 1 and oil phase 1, which is the purified wet-process phosphoric acid. NaCl solution was added to the oil phase 1 for back extraction. The volume ratio of NaCl solution to organic phase was 1:2. The reaction was allowed to stand at 40°C for 40 minutes, followed by separation to produce aqueous phase 2 and oil phase 2. Aqueous phase 2 was heated and filtered at 70°C. The filtrate was a sodium dihydrogen phosphate solution. The filter cake was dried at 120°C for 6 hours to produce sodium fluorosilicate solid. The moisture content after drying was less than 0.1%. The filtrate is filtered through cooling and crystallization to obtain sodium dihydrogen phosphate crystals. The mother liquor after filtration can dissolve NaCl and then return to the stripping stage. Ammonia water is added to the oil phase 2 for secondary stripping. After standing and stratification for 40 minutes, the liquid is separated to obtain oil phase 3 as trioctylamine. After recovery, the extraction cycle is repeated, and the aqueous phase 3 is an ammonium chloride salt solution. The dried filter cake fluorosilicate Na2SiF6 is ground and calcined at 500℃ under a nitrogen atmosphere for 2 hours. The SiF4 gas is collected by an air bag. After the reaction is complete, the solid NaF after pyrolysis is collected as a by-product for recovery. The SiF4 gas is further recovered after purification by liquid nitrogen low-temperature distillation. The final gas purity is >95%. In this example, the final fluorine extraction rate reaches 81.14%, the phosphoric acid extraction rate is 26.54%, the P2O5 content is 29.38wt.%, and the F content is 0.158wt.%.

[0187] Finally, magnesium dihydrogen phosphate is produced using purified wet-process phosphoric acid. The steps are as follows:

[0188] (1) Extraction: Cyclohexanone was used as the extractant to extract and purify wet-process phosphoric acid. The extraction temperature was 60°C, the volume ratio of the composite extractant to the purified wet-process phosphoric acid was 1:3, and the number of extraction stages was 3.

[0189] (2) Phase separation 1: The mixed solution of the composite extractant and wet-process phosphoric acid is allowed to stand for phase separation to obtain oil phase 1 and water phase 1. The organic oil phase of the coupled phosphoric acid is the extract, and the remaining water phase is water and a small amount of phosphoric acid. The water phase is used as the washing water for phosphogypsum. The washing water is returned to the phosphoric acid leaching tank. The phase separation temperature is 60°C and the phase separation time is 3 minutes.

[0190] (3) Stripping 1: Stripping the oil phase 1 obtained by separation 1 with magnesium sulfate, the stripping temperature is 60°C, the volume ratio of the stripping agent to the organic phase is 1:1, and the molar ratio of lithium ions to phosphoric acid carried in the extractant is 1:1.05;

[0191] (4) Phase separation 2: The mixed solution of the extract and the stripping solution is allowed to stand for phase separation to obtain oil phase 2 and water phase 2. The main components of water phase 2 are: H2PO4- Mg 2+ , H2O and a very small amount of SO4 2- The main component of oil phase 2 is the composite extractant carrying H2SO4, the separation temperature is 60℃, and the separation time is 3min;

[0192] (5) Stripping 2: Ammonia water was added to the oil phase 2 obtained from phase separation 2 for secondary stripping. The secondary stripping temperature was 50°C, the volume ratio of the stripping agent to the organic phase was 1:1, and the molar ratio of the stripping agent NH3 to the acid carried in the extractant was 1.1:1.

[0193] (6) Phase separation 3: The material obtained from stripping 2 is subjected to phase separation to obtain oil phase 3 and aqueous phase 3. After stripping, the composite extractant oil phase is returned to the extraction process, and the aqueous phase is subjected to subsequent operations;

[0194] (7) Evaporation and concentration 1: Evaporating and concentrating the aqueous phase 3 ammonium sulfate solution obtained from phase 3 to a specific gravity of 1.35 at a concentration temperature of 80°C and a pressure of 30 kPa;

[0195] (8) Cooling crystallization 2: The concentrated liquid is cooled and crystallized, and the cooling end temperature is 40°C;

[0196] (9) Filtration: The product obtained by cooling crystallization is filtered to obtain aqueous phase 4 and solid ammonium salt (NH4)2SO4. The liquid crystallization mother liquor is returned to the evaporation and concentration stage 1 in (7);

[0197] (10) Evaporation and concentration 2: The aqueous phase 4 obtained in (9) was evaporated and concentrated to a specific gravity of 1.45 at a concentration temperature of 80°C and a pressure of 20 kPa;

[0198] (11) Cooling crystallization 3: The concentrated liquid is subjected to cooling crystallization to obtain magnesium dihydrogen phosphate with a purity of 98.5% at a crystallization temperature of 30°C.

[0199] Example 6

[0200] Wet-process phosphoric acid is produced using a mixture of high-grade phosphate rock (30% P2O5 content) and white fertilizer (20% P2O5 content). High-grade phosphate rock, white fertilizer, sulfuric acid (98% concentration), and dilute phosphoric acid (30% concentration) are added to a decomposition tank in a specific proportion. The solid-to-liquid ratio is set at 10:1, the reaction temperature is controlled at 80°C, and the pH of the reaction mixture is adjusted to 0.8. The agitator is turned on and the reaction is stirred for 4 hours. The dilute acid concentration is maintained above 15%. After the reaction is completed, the phosphoric acid slurry is fed into a high-efficiency filtration device, where filtration produces wet-process phosphoric acid with a P2O5 content of 45% and solid phosphogypsum.

[0201] Then the metal ions are removed by solvent extraction.

[0202] The wet-process phosphoric acid was first subjected to decalcification pretreatment, the amount of the decalcifying agent sodium sulfate was 1.2 times the theoretical amount, the reaction temperature was 60° C., the reaction time was 40 minutes, the aging time was 1 hour, and the pretreated phosphoric acid was obtained after filtration.

[0203] Then extract the wet-process phosphoric acid to remove metal ions:

[0204] (1) Extraction: Wet-process phosphoric acid was subjected to three-stage three-stage countercurrent extraction. The extraction temperature of each stage was 60°C, and the O / A ratio was 4:1 (volume ratio).

[0205] (2) Washing: a. The metal ion-loaded extractant is first subjected to two-stage countercurrent washing with dilute phosphoric acid having a concentration of 4 wt.% P2O5 at a temperature of 65°C and an O / A ratio of 6:1 (volume ratio) to recover the phosphoric acid entrained in the extractant and improve the P2O5 yield in the phosphoric acid purification process.

[0206] (3) Stripping 1: Using 35 wt.% ammonium sulfate solution and 7 wt.% sulfuric acid solution as stripping agent 1, the metal ion-loaded extractant is subjected to two-stage three-stage countercurrent stripping. The stripping temperature of each stage is 45°C, and the O / A ratio is 5:1 (volume ratio), to obtain stripping solution 1 and organic phase 1. Stripping solution 1 is evaporated and concentrated, and filtered to obtain calcium sulfate; the filtrate is continuously cooled to 65°C to obtain ammonium manganese sulfate; the temperature is cooled to 35°C to obtain ammonium magnesium sulfate; the temperature is cooled to 4°C to obtain mixed crystals of potassium magnesium sulfate and sodium magnesium sulfate; this step classifies and recovers calcium, manganese, magnesium, potassium, and sodium ions.

[0207] (4) Washing 1: The organic phase 1 was washed with a 3 wt.% dilute ammonium sulfate solution in two stages in countercurrent at 45°C and a phase O / A ratio of 6:1 (volume ratio) to recover the ammonium sulfate entrained in the organic phase 1.

[0208] 5) Stripping 2: Using a 25 wt.% sulfuric acid solution as stripping agent 2, stripping the metal ions in the organic phase 1 at a temperature of 40° C. and a phase O / A ratio of 4:1 (volume ratio) to obtain a stripping solution 2 and an organic phase 2; adding ammonium sulfate solid to the stripping solution 2, cooling the solution to 65° C. to obtain cesium aluminum sulfate; cooling the solution to 40° C. to obtain potassium aluminum sulfate; and cooling the solution to 5° C. to obtain ammonium aluminum sulfate; and recycling the crystallization mother liquor as stripping agent 2.

[0209] (6) Washing 2: The organic phase 2 is washed with 8 wt.% dilute sulfuric acid in two stages in countercurrent at 40°C with an O / A ratio of 6:1 (volume ratio). The sulfuric acid entrained in the organic phase 2 is recovered and returned to the stripping agent 1 for recycling.

[0210] (7) Stripping 3: Using 16 wt.% oxalic acid solution as stripping agent 3, the organic phase 2 is subjected to two-stage countercurrent stripping at a temperature of 65°C and an O / A ratio of 4:1 (volume ratio) to obtain stripping solution 3 and organic phase 3. The stripping solution 3 is treated with ultraviolet light to obtain ferrous oxalate precipitate.

[0211] (8) Washing 3: The organic phase 3 is washed in two stages in countercurrent with dilute oxalic acid having a concentration of 1.5 wt.% at a temperature of 65°C and an O / A ratio of 6:1 (volume ratio). The oxalic acid entrained in the organic phase 3 is recovered to obtain a regenerated extractant that can be recycled.

[0212] After extraction, the metal ion content in phosphoric acid dropped to 0.69 wt.% Al2O3, 1.27 wt.% MgO, 0.20 wt.% Fe2O3, 0.53 wt.% CaO, with a MER value of 0.060. The P2O5 content was 45 wt.% and the F content was 1.46 wt.%.

[0213] Silicon dioxide was added to wet-process phosphoric acid in an amount of 1.2 times the theoretical value. The mixture was reacted at 40° C. for 40 minutes and aged for 1 hour. The pretreated wet-process phosphoric acid was obtained after filtration.

[0214] The purified phosphoric acid, free of metal ions, was subjected to extractive defluorination. Extraction was performed at 40°C. A 2:1 ratio (volume ratio) of extractant and wet-process phosphoric acid with a 1.46 wt% F concentration and a 45% P2O5 concentration was added to a reactor. The extractant composition was TOA: n-octanol: sulfonated kerosene (volume ratio = 0.15:0.15:0.70). After 60 minutes, the extractant was separated to produce aqueous phase 1 and oil phase 1, which is the purified wet-process phosphoric acid. KCl solution was added to the oil phase 1 for back extraction. The volume ratio of KCl solution to organic phase was 1:2. The reaction was continued at 40°C for 40 minutes, and the mixture was allowed to stand for stratification. The aqueous phase 2 was then separated to produce aqueous phase 2 and oil phase 2. The aqueous phase 2 was heated and filtered at 70°C. The filtrate was a potassium dihydrogen phosphate solution. The filter cake was dried at 120°C for 6 hours to produce potassium fluorosilicate solid. The moisture content after drying was less than 0.1%. The filtrate is cooled, crystallized and filtered to obtain potassium dihydrogen phosphate crystals. The mother liquor after filtration can dissolve KCl and return to the stripping stage. Ammonia water is added to the oil phase 2 for secondary stripping. After standing for 40 minutes and stratification, the liquid is separated to obtain oil phase 3 which is trioctylamine. After recovery, the extraction is circulated and the aqueous phase 3 is an ammonium chloride salt solution. The dried filter cake fluorosilicate K2SiF6 is ground and roasted at 500°C under a nitrogen atmosphere for 2 hours. The SiF4 gas is collected by an air bag. After the reaction is complete, the solid KF after pyrolysis is collected and recovered as a by-product. The SiF4 gas is further recovered after purification by liquid nitrogen low-temperature distillation. The final gas purity is >95%. In this example, the final fluorine extraction rate is 87.91%, the phosphoric acid extraction rate is 28.72%, the P2O5 content is 32.08wt.%, and the F content is 0.177wt.%. Finally, To produce DCP using purified phosphoric acid, calcium oxide (CaO) is added as a neutralizing agent to purified wet-process phosphoric acid (Al2O3 content: 0.69 wt.%, MgO content: 1.27 wt.%, Fe2O3 content: 0.20 wt.%, CaO content: 0.53 wt.%, MER value: 0.060, P2O5 content: 32.08 wt.%, and F content: 0.177 wt.%), which has been freed of metal ions and fluoride ions. The mixture is then added to a reactor at a mass ratio of 1:1.3 (solution:mixture). The reaction temperature is maintained at 75°C and the pH is controlled at 3.5 to produce a slurry. A filter press is used to separate the slurry into a solid rich in DCP and a mother liquor. The solid is washed and dried to obtain the final product. The mother liquor is returned to the washing section of the phosphate rock decomposition process to produce wet-process phosphoric acid and used to wash filtered phosphogypsum, achieving resource recycling.

[0215] Example 7

[0216] Decomposition of phosphate rock and white fertilizer → ion exchange to remove metal ions → extraction and defluorination → production of phosphate

[0217] Wet-process phosphoric acid is produced using a mixture of medium-grade phosphate rock (P2O5 content of 20%) and white fertilizer (P2O5 content of 20%). High-grade phosphate rock, white fertilizer, sulfuric acid (98%), and dilute phosphoric acid (30%) are added to a decomposition tank in a specific ratio. The solid-to-liquid ratio is set at 10:1, the reaction temperature is controlled at 80°C, and the pH of the reaction mixture is adjusted to 0.8. The agitator is turned on and the reaction is stirred for 4 hours. The dilute acid concentration is maintained above 15%. After the reaction is completed, the phosphoric acid slurry is fed into a high-efficiency filtration device, where filtration produces wet-process phosphoric acid with a P2O5 content of 40% and solid phosphogypsum.

[0218] This embodiment uses a cation exchange membrane impregnation method to purify wet-process phosphoric acid, including the following steps:

[0219] (1) Preparation of adsorption assembly: The cation exchange membrane loaded with sulfonic acid adsorption groups is installed and fixed in the adsorption assembly.

[0220] (2) Adsorption and desorption treatment of cationic impurities:

[0221] ① Adsorption: After heating the wet-process phosphoric acid to 50°C, the adsorption component is controlled to descend until the phosphate rock acid leaching liquid submerges the cation exchange membrane, with a residence time of 600 seconds;

[0222] ② Post-adsorption washing: lift the adsorption component and immerse it in washing tank 1; the residence time is 60 seconds; the washing tank 1 contains 3% nitric acid;

[0223] ③ Desorption: The washed adsorption component is immersed in a desorption tank containing a desorption liquid for 600 seconds; the desorption liquid is 30% sulfuric acid.

[0224] ④ Post-desorption washing: After desorption, the adsorption separation membrane assembly is immersed in a washing tank 2 for 60 seconds and then taken out; the washing tank 2 contains 5% sulfuric acid;

[0225] (3) The impurity content of the wet-process phosphoric acid after adsorption was tested: Al2O3 content was 0.43 wt.%, MgO content was 0.97 wt.%, Fe2O3 content was 0.27 wt.%, CaO content was 0.71 wt.%, and the MER value was 0.060. The P2O5 content was 40 wt.%, and the F content was 0.84 wt.%.

[0226] Silicon dioxide was added to wet-process phosphoric acid in an amount of 1.2 times the theoretical value. The mixture was reacted at 40° C. for 40 minutes and aged for 1 hour. The pretreated wet-process phosphoric acid was obtained after filtration.

[0227] The purified phosphoric acid, free of metal ions, was subjected to extractive defluorination. Extraction was performed at 40°C. A 2:1 ratio (volume ratio) of extractant and wet-process phosphoric acid with a 0.84 wt% F concentration and a 40% P2O5 concentration was added to a reactor. The extractant composition was TOA: n-octanol: sulfonated kerosene (volume ratio = 0.15:0.15:0.70). After 60 minutes, the extractant was separated to produce aqueous phase 1 and oil phase 1, which is the purified wet-process phosphoric acid. KNO3 solution was added to the oil phase 1 for back extraction. The volume ratio of KNO3 solution to organic phase was 1:2. The reaction was allowed to stand at 30°C for 40 minutes, followed by separation to produce aqueous phase 2 and oil phase 2. Aqueous phase 2 was heated and filtered at 70°C. The filtrate was a potassium dihydrogen phosphate solution. The filter cake was dried at 120°C for 6 hours to produce potassium fluorosilicate solid. The moisture content after drying was less than 0.1%. The filtrate is cooled, crystallized, and filtered to obtain potassium dihydrogen phosphate crystals. The filtered mother liquor can dissolve KNO3 and then return to the stripping stage. Ammonia water is added to the oil phase 2 for secondary stripping. After standing for 40 minutes and stratification, the liquid is separated to obtain oil phase 3, which is trioctylamine N235. After recovery, the extraction cycle is repeated, and the aqueous phase 3 is an ammonium nitrate salt solution. The dried filter cake, fluorosilicate K2SiF6, is ground and calcined at 500°C under a nitrogen atmosphere for 2 hours. SiF4 gas is collected using an air bag. After the reaction is complete, the solid KF after pyrolysis is collected as a by-product and recovered. The SiF4 gas is further recovered after purification by liquid nitrogen low-temperature distillation. The final gas purity is >95%. In this example, the final fluorine extraction rate reaches 88.44%, the phosphoric acid extraction rate is 26.54%, the P2O5 content is 28.87wt.%, and the F content is 0.097wt.%.

[0228] The metal ions and fluoride ions are removed to purify the wet-process phosphoric acid, and the Al2O3 content is 0.43wt.%, the MgO content is 0.97wt.%, the Fe2O3 content is 0.57wt.%, the CaO content is 0.71wt.%, the P2O5 content is 28.87wt.%, and the F content is 0.097wt.%.

[0229] Finally, potassium dihydrogen phosphate is prepared using purified wet-process phosphoric acid in the following steps:

[0230] (1) Extraction: Using tributyl phosphate as the extractant to extract and purify wet-process phosphoric acid, the extraction temperature is 60°C, the volume ratio of the composite extractant to the purified wet-process phosphoric acid is 1:3, and the number of extraction stages is 3;

[0231] (2) Phase separation 1: The mixed solution of the composite extractant and wet-process phosphoric acid is allowed to stand for phase separation to obtain an oil phase 1 and an aqueous phase 1. The organic oil phase of the coupled phosphoric acid is the extract, and the remaining aqueous phase is water and a small amount of phosphoric acid. The aqueous phase is used as the washing water for phosphogypsum. The washing water after washing is returned to the phosphoric acid leaching tank. The phase separation temperature is 60°C and the phase separation time is 3 minutes.

[0232] (3) Stripping 1: stripping the oil phase 1 obtained by phase separation 1 with KOH, the stripping temperature is 60°C, the volume ratio of the stripping agent to the organic phase is 1:1, and the molar ratio of lithium ions to phosphoric acid carried in the extractant is 1:1.05;

[0233] (4) Phase separation 2: The mixed solution of the extract and the stripping solution is allowed to stand for phase separation to obtain oil phase 2 and water phase 2. The main components of water phase 4 are: H2PO4 - , K + , H2O, the phase separation temperature is 60℃, and the phase separation time is 3min;

[0234] (5) Stripping 2: Ammonia water was added to the oil phase 2 obtained from phase separation 2 for secondary stripping. The secondary stripping temperature was 50°C, the volume ratio of the stripping agent to the organic phase was 1:1, and the molar ratio of the stripping agent NH3 to the acid carried in the extractant was 1.1:1.

[0235] (6) Phase separation 3: The material obtained from stripping 2 is subjected to phase separation to obtain oil phase 3 and water phase 3. The oil phase of the composite extractant after stripping is returned to the extraction process, and the water phase is subjected to subsequent operations;

[0236] (7) Evaporation and concentration 2: The aqueous phase 3 obtained in (6) was evaporated and concentrated to a specific gravity of 1.45 at a concentration temperature of 80°C and a pressure of 20 kPa;

[0237] (8) Cooling crystallization 3: The concentrated liquid is subjected to cooling crystallization to obtain potassium dihydrogen phosphate with a purity of 98.7% at a crystallization temperature of 30°C.

[0238] Example 8

[0239] Phosphate rock and white fertilizer decomposition → ion exchange demetallization → extraction defluorination → DCP

[0240] Wet-process phosphoric acid is produced using a mixture of high-grade phosphate rock (30% P2O5 content) and white fertilizer (20% P2O5 content). High-grade phosphate rock, white fertilizer, sulfuric acid (98% concentration), and dilute phosphoric acid (30% concentration) are added to a decomposition tank in a specific proportion. The solid-to-liquid ratio is set at 10:1, the reaction temperature is controlled at 80°C, and the pH of the reaction mixture is adjusted to 0.8. The agitator is turned on and the reaction is stirred for 4 hours. The dilute acid concentration is maintained above 15%. After the reaction is completed, the phosphoric acid slurry is fed into a high-efficiency filtration device, where filtration produces wet-process phosphoric acid with a P2O5 content of 45% and solid phosphogypsum.

[0241] This embodiment uses a cation exchange membrane fixed bed method to purify wet-process phosphoric acid, including the following steps:

[0242] (1) Prepare a cation exchange membrane loaded with sulfonic acid adsorption groups.

[0243] (2) Adsorption and desorption treatment of cationic impurities:

[0244] ① Adsorption: Wet-process phosphoric acid heated to 50°C flows through the cation exchange membrane at a controlled flow rate of 10 cm / min;

[0245] ② Post-adsorption washing: After adsorption, washing liquid 1 is introduced at a flow rate of 20 cm / min; washing liquid 1 is pure water;

[0246] ③ Desorption: After washing, the desorption liquid is introduced at a flow rate of 10 cm / min; the desorption liquid is 30% sulfuric acid;

[0247] ④ Washing after desorption: introduce washing liquid 2 at a flow rate of 20 cm / min; washing liquid 2 is 5% sulfuric acid;

[0248] (3) The impurity content of the wet-process phosphoric acid after adsorption was tested: Al2O3 content was 0.54 wt.%, MgO content was 1.15 wt.%, Fe2O3 content was 0.29 wt.%, CaO content was 0.65 wt.%, and the MER value was 0.058. P2O5 content was 45 wt.%, and F content was 1.46 wt.%.

[0249] Silicon dioxide was added to wet-process phosphoric acid in an amount of 1.2 times the theoretical value. The mixture was reacted at 40° C. for 40 minutes and aged for 1 hour. The pretreated wet-process phosphoric acid was obtained after filtration.

[0250] The purified phosphoric acid, free of metal ions, was subjected to extractive defluorination. Extraction was performed at 40°C. A 1.5:1 ratio (volume ratio) of extractant and wet-process phosphoric acid with a 1.46 wt% F concentration and a 45 wt% P2O5 concentration was added to a reactor. The extractant composition was TOA:n-octanol:sulfonated kerosene (volume ratio = 0.15:0.15:0.70). After 60 minutes, the extractant was separated to produce aqueous phase 1 and oil phase 1, which is the purified wet-process phosphoric acid. K2SO4 solution was added to the oil phase 1 for back extraction. The volume ratio of K2SO4 solution to organic phase was 1:2. The reaction was allowed to stand at 30°C for 40 minutes, followed by separation to produce aqueous phase 2 and oil phase 2. Aqueous phase 2 was heated and filtered at 70°C. The filtrate was a potassium dihydrogen phosphate solution. The filter cake was dried at 120°C for 6 hours to produce potassium fluorosilicate solid. The moisture content after drying was less than 0.1%. The filtrate is cooled, crystallized, and filtered to obtain potassium dihydrogen phosphate crystals. The filtered mother liquor can dissolve K2SO4 and then return to the stripping stage. Ammonia water is added to the oil phase 2 for secondary stripping. After standing for 40 minutes and stratification, the liquid is separated to obtain oil phase 3 as trioctylamine, which is recycled and extracted. The aqueous phase 3 is an ammonium sulfate salt solution. The dried filter cake fluorosilicate K2SiF6 is ground and calcined at 500°C under a nitrogen atmosphere for 2 hours. The SiF4 gas is collected using an air bag. After the reaction is complete, the solid KF after pyrolysis is collected as a by-product. The SiF4 gas is further recovered after purification by liquid nitrogen low-temperature distillation. The final gas purity is >95%. In this example, the final fluorine extraction rate reaches 90.02%, the phosphoric acid extraction rate is 28.43%, the P2O5 content is 32.21wt.%, and the F content is 0.146wt.%.

[0251] After removing metal ions and fluoride ions, the Al2O3 content in the purified wet-process phosphoric acid is 0.54wt.%, the MgO content is 1.15wt.%, the Fe2O3 content is 0.69wt.%, the CaO content is 0.65wt.%, and the MER value is 0.067.

[0252] The P2O5 content is 32.21wt.%, and the F content is 0.146wt.%.

[0253] Finally, purified phosphoric acid is used to produce DCP. Calcium oxide is added as a neutralizing agent to the purified wet-process phosphoric acid, which has been free of metal ions and fluoride ions. This is then added to the reactor at a mass ratio of 1:1.3 (solution:mixture). The reaction temperature is maintained at 75°C and the pH is controlled at 3.5 to produce a slurry. A filter press is used to separate the slurry into a solid rich in DCP and a mother liquor. The solid is washed and dried to produce the final product. The mother liquor is returned to the washing section of the phosphate rock decomposition process to produce wet-process phosphoric acid and used to wash the filtered phosphogypsum, thus achieving resource recycling.

[0254] Comparative Example 1

[0255] Wet-process phosphoric acid is produced from low-grade phosphate rock with a 20% P2O5 content. Phosphate rock, 98% sulfuric acid, and 20% dilute phosphoric acid are added to a decomposition tank in a specific ratio. The solid-liquid ratio is set at 5:1, the reaction temperature is controlled at 80°C, and the pH of the reaction mixture is adjusted to 1.0. The agitator is turned on and the reaction is stirred for 6 hours. The dilute acid concentration is maintained above 10%. After the reaction is completed, the phosphoric acid slurry is passed through a filtration device, where it is filtered to produce wet-process phosphoric acid with a 30% P2O5 content and phosphogypsum.

[0256] Before the metal ions are removed by solvent extraction, the wet-process phosphoric acid is pretreated with decalcification. The decalcification agent is ammonium sulfate, the amount of the decalcification agent is 1.2 times the theoretical amount, the reaction temperature is 60° C., the reaction time is 40 minutes, the aging time is 1 hour, and the pretreated phosphoric acid is obtained after filtration.

[0257] Then extract wet-process phosphoric acid to remove metal ions:

[0258] (1) Extraction: Wet-process phosphoric acid was subjected to three-stage three-stage countercurrent extraction. The extraction temperature of each stage was 60°C, and the O / A ratio was 4:1 (volume ratio).

[0259] (2) Washing: The metal ion-loaded extractant is first subjected to two-stage countercurrent washing in dilute phosphoric acid with a concentration of 1 wt.% P2O5 at a temperature of 50°C and an O / A ratio of 3:1 (volume ratio) to recover the phosphoric acid entrained in the extractant and improve the P2O5 yield in the phosphoric acid purification process.

[0260] (3) Stripping 1: Using 35 wt.% ammonium sulfate solution and 1 wt.% sulfuric acid solution as stripping agent 1, the metal ion-loaded extractant is subjected to two-stage three-stage countercurrent stripping. The stripping temperature of each stage is 30°C, and the O / A ratio is 2:1 (volume ratio), to obtain stripping solution 1 and organic phase 1. Stripping solution 1 is evaporated and concentrated, filtered to obtain calcium sulfate; the filtrate is continuously cooled to 80°C to obtain ammonium manganese sulfate; cooled to 50°C to obtain ammonium magnesium sulfate; cooled to 10°C to obtain mixed crystals of potassium magnesium sulfate and sodium magnesium sulfate; this step classifies and recovers calcium, manganese, magnesium, potassium, and sodium ions.

[0261] (4) Washing 1: The organic phase 1 was washed with a dilute ammonium sulfate solution having a concentration of 0.1 wt.% in two stages in countercurrent at a temperature of 30° C. and a phase O / A ratio of 3:1 (volume ratio) to recover the ammonium sulfate entrained in the organic phase 1.

[0262] (5) Stripping 2: Using a sulfuric acid solution with a concentration of 30 wt.% as stripping agent 2, stripping the metal ions in the organic phase 1 at a temperature of 30°C and a phase O / A ratio of 1:1 (volume ratio) to obtain stripping solution 2 and organic phase 2; adding ammonium sulfate solid to the stripping solution 2, cooling it to 80°C to obtain cesium aluminum sulfate; cooling it to 55°C to obtain potassium aluminum sulfate; cooling it to 20°C to obtain ammonium aluminum sulfate; the crystallization mother liquor is recycled as stripping agent 2.

[0263] (6) Washing 2: The organic phase 2 is washed with dilute sulfuric acid having a concentration of 5 wt.% in two-stage countercurrent at a temperature of 30°C and an O / A ratio of 3:1 (volume ratio). The sulfuric acid entrained in the organic phase 2 is recovered and returned to the stripping agent 1 for recycling.

[0264] (7) Stripping 3: Using 10 wt.% oxalic acid solution as stripping agent 3, the organic phase 2 is subjected to two-stage countercurrent stripping at a temperature of 50°C and an O / A ratio of 1:1 (volume ratio) to obtain stripping solution 3 and organic phase 3. The stripping solution 3 is treated with ultraviolet light to obtain ferrous oxalate precipitate.

[0265] (8) Washing 3: The organic phase 3 is washed in two stages in countercurrent with dilute oxalic acid having a concentration of 1 wt.% at a temperature of 50°C and an O / A ratio of 3:1 (volume ratio). The oxalic acid entrained in the organic phase 3 is recovered to obtain a regenerated extractant that can be recycled.

[0266] The extractant is returned to the wet-process phosphoric acid for recycling.

[0267] After extraction, the metal ion content in phosphoric acid dropped to 0.58 wt.% Al2O3, 1.14 wt.% MgO, 0.25 wt.% Fe2O3, 0.54 wt.% CaO, with a MER value of 0.084, 30 wt.% P2O5, and 0.68 wt.% F.

[0268] Silicon dioxide was added to wet-process phosphoric acid in an amount of 1.2 times the theoretical value. The mixture was reacted at 40° C. for 40 minutes and aged for 1 hour. The pretreated wet-process phosphoric acid was obtained after filtration.

[0269] Extraction is then carried out at 40°C. A separatory funnel is charged with an extractant (trioctylamine (TOA) / n-octanol / sulfonated kerosene) in a 1:1 volume ratio, along with wet-process phosphoric acid (0.68 wt% F and 30 wt% P2O5). The extractant composition is trioctylamine (TOA), n-octanol, and sulfonated kerosene (0.05:0.15:0.80). After 60 minutes, the phases separate to produce aqueous phase 1 and oil phase 1, which is the purified wet-process phosphoric acid. KOH aqueous solution is added to oil phase 1 for back extraction. The volume ratio of KOH aqueous solution to organic phase is 1:2. The reaction is allowed to proceed at 30°C for 40 minutes, followed by stratification and separation to produce aqueous phase 2 and oil phase 2. Aqueous phase 2 is heated and filtered at 70°C. The filtrate is a potassium dihydrogen phosphate solution. The filter cake is dried at 120°C for 6 hours to produce potassium fluorosilicate solid. The moisture content after drying is less than 0.1%. The filtrate is cooled, crystallized, and filtered to yield potassium dihydrogen phosphate crystals. The filtered mother liquor can be used to dissolve KOH and return to the stripping stage. The oil phase 2 is heated to 50°C, allowed to stand for 40 minutes, and then separated to yield trioctylamine (Olive Phase 3), which is recovered for recycling and extraction. The aqueous phase 3, water, can dissolve KOH and return to the stripping stage. The dried filter cake, K2SiF6 fluorosilicate, is ground and calcined at 500°C under a nitrogen atmosphere for 2 hours. SiF4 gas is collected using an air bag. After the reaction is complete, the pyrolyzed solid KF is collected as a byproduct and recovered. The SiF4 gas is purified by cryogenic distillation with liquid nitrogen and further recovered. The final gas purity is >95%. In this example, the final fluorine extraction rate reached 28.86%, the phosphoric acid extraction rate was 6.82%, the P2O5 content was 27.95 wt.%, and the F content was 0.484 wt.%.

[0270] Finally, purified phosphoric acid is used to produce DCP. Calcium oxide is added as a neutralizing agent to the purified wet-process phosphoric acid, which has been free of metal ions and fluoride ions. This is then added to the reactor at a mass ratio of 1:1.3 (solution:mixture). The reaction temperature is maintained at 75°C and the pH is controlled at 3.5 to produce a slurry. A filter press is used to separate the slurry into a solid rich in DCP and a mother liquor. The solid is washed and dried to produce the final product. The mother liquor is returned to the washing section of the phosphate rock decomposition process to produce wet-process phosphoric acid and used to wash the filtered phosphogypsum, thus achieving resource recycling.

Claims

1. A method for wet processing of phosphate rock, characterized in that: The method for wet processing of phosphate rock comprises: A. Acid leaching of phosphate rock: The phosphate rock raw material is acid-lyzed and leached, and the solid-liquid separation is carried out to obtain wet-process phosphoric acid; B. Removing metal ions from wet-process phosphoric acid; C. Wet-process phosphoric acid defluorination: using a wet-process phosphoric acid defluorination extractant to extract fluorosilicic acid in wet-process phosphoric acid at 10-70° C., and separating the phases to obtain defluorinated purified phosphoric acid and an organic phase loaded with fluorosilicate; the wet-process phosphoric acid defluorination extractant is a compound of an organic amine, a cosolvent, and a diluent in a volume ratio of 10-100:0-30:0-90, and the content of the cosolvent and the diluent is not 0; the organic amine includes a primary amine RNH2, a secondary amine R2NH, a tertiary amine R3N, and a quaternary ammonium base R4N + OH - At least one of the following, wherein R is a hydrocarbon group; the cosolvent includes at least one of n-butanol, n-hexanol, cyclohexanol, 2-ethyl-1-hexanol, and n-octanol; the diluent includes at least one of sulfonated kerosene, n-dodecane, isomeric dodecane, n-tetradecane, n-hexadecane, and 260# solvent oil.

2. The method for wet processing of phosphate rock according to claim 1, characterized in that: The phosphate rock raw material in step A comprises at least one of phosphate rock, white fertilizer, and slag acid, preferably, the phosphate rock raw material contains 5wt% to 35wt% of P2O5; Preferably, the acid used for the acidolysis is sulfuric acid or dilute phosphoric acid, and the acid concentration of the acidolysis is ≥5%; the solid-liquid ratio of the acidolysis is 0.1-15, the reaction temperature is 0-90° C., the pH value of the reaction material is controlled within 0.5-1.5, and the stirring reaction is carried out for 0.5-12 hours.

3. The method for wet processing of phosphate rock according to claim 1 or 2, characterized in that: The method for removing metal ions from wet-process phosphoric acid in step B is extraction or ion exchange; Preferably, a decalcification pretreatment is performed before the extraction and removal of metal ions, the amount of the decalcifying agent is 0.8 to 1.2 times the theoretical amount, the reaction temperature is 40 to 60° C., the reaction time is 20 to 40 minutes, the aging time is 0.5 to 2 hours, and the extraction and removal of metal ions are performed after solid-liquid separation; the decalcifying agent is preferably at least one of sulfuric acid, ammonium sulfate, sodium sulfate, and potassium sulfate; The method for extracting and removing metal ions preferably adopts the method disclosed in the invention patent publication number CN119265410A; Or the method for extracting and removing metal ions preferably adopts the method disclosed in the invention patent publication number CN119503888A, wherein the temperature of step (3) of CN119503888A is 25-85°C, and the volume ratio of O / A is 1:5-10:1; step (4) adopts method 3, and in method 3 of step (4), ammonia is passed through to adjust the pH to 7-12; the temperature of ammonia evaporation is 30-80°C; Alternatively, a mixed solution of oxalic acid and ammonium oxalate is used as a stripping agent instead of the stripping agent in step (3) of CN119503888A, wherein the ammonium oxalate concentration in step (3) is 1 wt.% to 20 wt.%, the oxalic acid concentration is 0 wt.% to 20 wt.%, the temperature is 25 to 85°C, and the volume ratio O / A is 1:5 to 10:1; step (4) adopts method 3, wherein ammonia is passed through method 3 to adjust the pH to 7 to 12; and the temperature of ammonia evaporation is 30 to 80°C; The method for removing metal ions by ion exchange preferably adopts the method disclosed in the invention patent publication number CN116786096A.

4. The method for wet processing of phosphate rock according to claim 1, characterized in that: The wet-process phosphoric acid defluorination extractant is prepared by mixing an organic amine, a cosolvent and a diluent in a volume ratio of 10-15:10-20:65-75.

5. The method for wet processing of phosphate rock according to claim 1 or 2, characterized in that: The step C further comprises: Stripping: reacting the fluorosilicate-loaded organic phase with a stripping agent, separating the phases to obtain a fluorosilicate solution and a regenerated organic phase, wherein the stripping agent is an inorganic alkali aqueous solution or an inorganic salt aqueous solution; Extractant regeneration: adding ammonia water to the regenerated organic phase to obtain an ammonium salt solution and a regenerated extractant, or heating the regenerated organic phase to obtain an aqueous solution and a regenerated extractant. The regenerated extractant is returned to step C for recycling, and the aqueous solution is returned to prepare a stripping agent.

6. The method for wet processing of phosphate rock according to claim 1, characterized in that: The method further comprises step C of supplementing silicon before extraction: adding a silicon supplement agent to the wet-process phosphoric acid raw material, wherein the amount of silicon supplemented is 1.1 to 1.2 times the theoretical value, reacting at 10 to 90° C. for 20 to 40 minutes, aging for 1 hour, and filtering; The silicon supplementing agent is preferably silicon dioxide, diatomaceous earth or sodium silicate.

7. The method for wet processing of phosphate rock according to claim 1, characterized in that: The wet-process phosphoric acid in step C has a mass concentration of 20% to 48% as calculated as P2O5; The mass concentration of fluorine in the wet-process phosphoric acid before defluorination in step C is 0.1 wt.% to 5.0 wt.%, preferably 0.1 wt.% to 1.5 wt.%; the extraction time in step C is preferably 10 to 90 minutes, and the volume ratio of the extractant to the wet-process phosphoric acid raw material is preferably 1 to 2; more preferably, the volume ratio of the extractant to the wet-process phosphoric acid raw material is 1.5 to 2.

8. The method for wet processing of phosphate rock according to claim 1, characterized in that: The fluorine extraction rate of the extraction in step C is above 60%, and the phosphoric acid extraction rate is below 30%; preferably, the fluorine extraction rate is above 82%.

9. The method for wet processing of phosphate rock according to claim 1, characterized in that: In step C, the volume ratio of the organic phase to the stripping agent is 2:1 to 1:5, the stripping temperature is 10 to 70° C., and the inorganic base is preferably KOH or NaOH; Step C also includes fluorine resource recovery: The fluorosilicate solution of step C is subjected to solid-liquid separation at 70-90° C. to obtain a fluorosilicate ointment and a filtrate, the fluorosilicate ointment is dried to obtain a solid product fluorosilicate, and the solid product fluorosilicate is pyrolyzed at 200-800° C. to generate SiF4 and a fluoride salt, or acidolyzed at 95-105% sulfuric acid to generate SiF4 and HF; Preferably, the pyrolysis is carried out under a nitrogen atmosphere, the pyrolysis temperature of fluorosilicate K2SiF6 is 400-600°C, and the pyrolysis temperature of Na2SiF6 is 300-800°C; Phosphoric acid recovery: the filtrate is cooled and crystallized, and solid-liquid separation is performed to obtain phosphate solid and mother liquor, and the mother liquor is recycled for the stripping step.

10. The method for wet processing of phosphate rock according to claim 1, characterized in that: The method further comprises the step D. preparing phosphate using the defluorinated purified phosphoric acid; Preferably, the step D comprises: neutralizing the defluorinated purified phosphoric acid with a neutralizing agent, separating the solid and liquid, washing and drying to obtain phosphate, wherein the neutralizing agent is at least one of calcium oxide, calcium hydroxide, and calcium carbonate, and controlling the pH range to be 2.0 to 6.0 to obtain calcium dihydrogen phosphate and / or calcium phosphate; Or step D adopts steps D to G of the invention patent with publication number CN 118419875A to prepare phosphate.

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