A purification method and purification system for wet-process phosphoric acid

Abstract

The present disclosure provides a purification method and system for wet-process phosphoric acid, belonging to the technical field of wet-process phosphoric acid purification. The present disclosure subjects the raffinate acid produced by extraction to secondary extraction, returns it to be mixed with the wet-process phosphoric acid, and returns the washed raffinate acid to the desulfurization treatment step for recycling to form an acid closed loop. The entire wet-process phosphoric acid purification process does not produce acid byproducts, is environmentally friendly, and the yield of the purified phosphoric acid product is greater than 95%, eliminating the need to worry about the treatment of acid byproducts.

Description

Technical Field

[0001] The present disclosure relates to the technical field of refined phosphoric acid, and in particular to a purification method and purification system for wet-process phosphoric acid. Background Art

[0002] Phosphoric acid is an important chemical raw material and intermediate product, widely used in fertilizer, electronics, medicine, food and other industries. At present, phosphoric acid production is mainly divided into two categories: thermal process and wet process. Due to the high energy consumption, severe environmental pollution and high requirements for ore in thermal process production, the world's energy shortage has caused the price of thermal process phosphoric acid to continue to rise. Therefore, wet process phosphoric acid has begun to gradually replace thermal process phosphoric acid. The impurity content in wet process phosphoric acid is relatively high, mainly including anion SO4 2- 、F - As well as cations such as Fe, Al, Mg, Ca, and Si. Industrial phosphoric acid production of end products generally requires phosphoric acid of relatively good quality. Therefore, wet-process phosphoric acid must be purified to meet production quality standards.

[0003] Currently, the main methods for purifying wet-process phosphoric acid include crystallization, ion exchange, chemical precipitation, and solvent extraction. Crystallization, due to the need for high concentration, presents challenges in selecting corrosion-resistant materials. Furthermore, due to the presence of entrained crystals, it is difficult to obtain high-purity acid, and the yield of the crystalline product is low. While the ion exchange process equipment is relatively simple, the ion exchange resin is expensive, the phosphoric acid exiting the resin has a low concentration, and the evaporation rate is high. This high energy consumption makes it unsuitable for large-scale production, and the initial investment is high. Chemical precipitation can remove some cations or anions, but the purification depth is limited and it introduces another ion. Solvent extraction, due to its excellent purification results, mild conditions, high yield, and stable operation, has become the most commonly used method.

[0004] The core technology behind solvent extraction purification of wet-process phosphoric acid involves extracting and separating impurity-laden wet-process phosphoric acid through an organic solvent to produce industrial-grade phosphoric acid. However, the P2O5 extraction rate for solvent extraction purification of wet-process phosphoric acid is generally 50% to 70%, resulting in a low purification yield. Furthermore, the byproduct, raffinate acid, is phosphoric acid containing P2O5 and high levels of iron, aluminum, and magnesium impurities. The impurity ion content in raffinate acid is exponentially higher than that of the original phosphoric acid, significantly limiting its application. Currently, domestic raffinate acid utilization is primarily focused on the production of monoammonium phosphate fertilizer. In recent years, due to overcapacity in high-concentration compound fertilizer production and significant price fluctuations in the phosphate fertilizer market, the rational utilization of the byproduct raffinate acid has become a pressing technical challenge for manufacturers. Recovering the phosphorus from the byproduct raffinate acid, or eliminating its production, would significantly improve economic efficiency. Summary of the Invention

[0005] The purpose of the present invention is to overcome the deficiencies of the prior art and provide a method and system for purifying wet-process phosphoric acid, thereby solving the problem of recycling raffinate and improving the purification effect of wet-process phosphoric acid.

[0006] To achieve the above objectives, the technical solution adopted by the present invention is as follows: On the one hand, the present disclosure provides a method for purifying wet-process phosphoric acid, comprising the following steps:

[0007] The wet-process phosphoric acid is sequentially subjected to pretreatment, extraction, desulfurization, washing, stripping, and final treatment to obtain a phosphoric acid product;

[0008] The extraction steps include: performing a first extraction on the crude phosphoric acid obtained from the pretreatment with a first extractant to obtain a first extraction phase and a first raffinate acid; performing a second extraction on the first raffinate acid with a second extractant to obtain a second raffinate acid and a second extraction phase, and returning the second raffinate acid to the pretreatment step to be mixed with wet-process phosphoric acid; and removing impurities from the second extraction phase and returning it to the second extraction step as a second extraction agent.

[0009] The first extractant is a mixture of tributyl phosphate, sulfonated kerosene and tridecanol.

[0010] In some embodiments, the volume ratio of tributyl phosphate, sulfonated kerosene and tridecanol is tributyl phosphate: sulfonated kerosene: tridecanol = 60-75: 20-30: 5-10;

[0011] And / or, the mass fraction of P2O5 in the crude phosphoric acid obtained by the pretreatment is 48-50%.

[0012] In some embodiments, the steps include:

[0013] Pretreatment: Pre-treating wet-process phosphoric acid to obtain crude phosphoric acid;

[0014] Extraction: The crude phosphoric acid is subjected to a first extraction with a first extractant to obtain a first extraction phase and a first raffinate acid; the first raffinate acid is subjected to a second extraction to obtain a second raffinate acid and a second extraction phase, and the second raffinate acid is returned to the pretreatment step to be mixed with wet-process phosphoric acid; the second extraction phase is subjected to impurity removal and then returned to the second extraction step to serve as a second extraction agent; wherein the first extraction agent is a mixture of tributyl phosphate, sulfonated kerosene, and tridecanol;

[0015] Desulfurization treatment: desulfurizing the first extraction phase obtained by extraction with a desulfurizing agent to obtain a desulfurized organic phase and desulfurized sulfuric acid;

[0016] Washing: washing the desulfurized organic phase with washing acid to obtain a washed organic phase and a washing residual acid, and returning the washing residual acid to the desulfurization treatment step to prepare a desulfurizer;

[0017] Stripping: stripping the washed organic phase with a stripping agent to obtain stripping acid and a stripping organic phase, regenerating the stripping organic phase to obtain a regenerated organic phase, and returning the regenerated organic phase to the extraction step as the first extractant;

[0018] Final treatment: The stripping acid obtained by stripping is returned to the washing step as washing acid, and the remaining part is concentrated and then subjected to final treatment to obtain the phosphoric acid product.

[0019] In some embodiments, the pretreatment comprises the following steps: sequentially desulfurizing, defluorinating, dearsenicizing, dechlorinating, and decolorizing the wet-process phosphoric acid, and then filtering and collecting the filtrate;

[0020] Preferably, the steps of desulfurization, defluorination, dearsenicization, dechlorination and decolorization are:

[0021] Mixing phosphate rock powder, sodium carbonate, silicon dioxide and the desulfurized water to obtain a desulfurization and defluorination agent, and mixing a sulfur-containing compound with the desulfurized water to obtain a dearsenicizing agent; wherein the sulfur-containing compound is sodium sulfide or phosphorus pentasulfide;

[0022] The wet-process phosphoric acid is sequentially added with a desulfurization and defluorination agent for desulfurization and defluorination, a dearsenicizing agent for arsenic removal, ozone for dechlorination, and activated carbon for decolorization; wherein the stoichiometric ratio of phosphate rock to sulfate in the wet-process phosphoric acid is 1.05-1.2, the stoichiometric ratio of the total amount of sodium carbonate and silicon dioxide to the fluoride ion in the wet-process phosphoric acid is 1.2-1.4, the stoichiometric ratio of sulfur atoms in the dearsenicizing agent to arsenic ions in the wet-process phosphoric acid is 10-20, and the volume ratio of ozone to the volume of the wet-process phosphoric acid is 500-1000.

[0023] More preferably, the temperature of the desulfurization, defluorination, dearsenicization, dechlorination and decolorization is 50-60° C., and / or the negative pressure of the dearsenicization, dechlorination and decolorization is 0.3-1.5 kPa.

[0024] In some embodiments, during the extraction step, at least one of the following is satisfied:

[0025] (1) The first extraction is a multi-stage countercurrent extraction with 4-6 extraction stages;

[0026] (2) the volume ratio of the first extractant to crude phosphoric acid is 4-6;

[0027] (3) the temperature of the first extraction is 40-50° C.;

[0028] (4) The second extraction step comprises: adding the first raffinate acid to hydrogen peroxide and reacting at a temperature of 60-70° C. to obtain an oxidized first raffinate acid; then adding a second extractant and performing multi-stage cross-current extraction to obtain a second raffinate acid and a second extract phase;

[0029] Preferably, the mass of the hydrogen peroxide is 0.1-0.5% of the mass of the first raffinate acid;

[0030] Preferably, the second extractant is prepared by mixing the extractant and sulfonated kerosene, adding sodium hydroxide solution to carry out saponification reaction, and separating the phases to obtain an organic phase, which is the second extractant; wherein the extractant is at least one of 2-ethylhexyl phosphate, di(2-ethylhexyl) phosphate, dinonylnaphthalenesulfonic acid, cyclohexane acid, trialkylphosphine oxide, and acetylacetone, and the saponification rate of the second extractant is 10%-30%;

[0031] More preferably, the volume ratio of the extractant to sulfonated kerosene is 1-2.5, the number of stages of the multi-stage cross-current extraction is 3-5, the volume ratio of the second extractant to the first raffinate after oxidation is 3-5, and the temperature of the multi-stage cross-current extraction is 40-60°C;

[0032] (5) The impurity removal step of the second extraction phase is as follows: the second extraction phase is washed with 1%-10% by mass phosphoric acid, and then stripped with 30%-40% by mass sulfuric acid to obtain a metal stripping solution and a metal organic phase.

[0033] In some embodiments, the desulfurization treatment step is:

[0034] Solution preparation: Calcium salt and residual washing acid are prepared into calcium salt solution, and barium salt and residual washing acid are prepared into barium salt solution;

[0035] Desulfurization: The first extraction phase obtained from the first extraction and the second desulfurization phase are mixed to undergo a first reaction, followed by phase separation to obtain a first organic phase and a first desulfurization phase, and the first desulfurization phase is returned to the pretreatment step;

[0036] mixing the first organic phase and the calcium salt solution and performing a second reaction to obtain a second organic phase and a second desulfurization;

[0037] The second organic phase and the barium salt solution are mixed and then subjected to a third reaction to obtain a desulfurized organic phase and a third desulfurization step; wherein the third desulfurization step is returned to the pretreatment step;

[0038] Preferably, the volume ratio of the first extraction phase to the second desulfurization phase is 10-15, the volume ratio of the first organic phase to the calcium salt solution is 10-15, the volume ratio of the second organic phase to the barium salt solution is 10-15, and the temperature of the first reaction, the second reaction and the third reaction is independently 40-50°C;

[0039] More preferably, the calcium salt is calcium carbonate or calcium hydroxide, and the barium salt is barium carbonate.

[0040] In some embodiments, in the washing step, the washing is countercurrent washing, and the number of washing stages is 2-3;

[0041] and / or, the volume ratio of the desulfurized organic phase to the washing acid is 12-18;

[0042] And / or, the washing temperature is 50-60°C.

[0043] In some embodiments, in the stripping step, the stripping agent is desalted water;

[0044] and / or, the volume ratio of the washing organic phase to the stripping agent is 4-6;

[0045] And / or, the number of stages of the stripping is 2-4;

[0046] And / or, the stripping temperature is 60-70°C;

[0047] And / or, the regeneration step of the stripping organic phase comprises: washing the stripping organic phase with a sodium hydroxide aqueous solution, adding activated carbon to adsorb impurities, and then washing with desalted water to obtain a regenerated organic phase;

[0048] Preferably, the mass fraction of the sodium hydroxide aqueous solution is 4%-6%, and the mass of the activated carbon is 0.5%-1% of the mass of the stripping organic phase.

[0049] On the other hand, a purification system for the wet-process phosphoric acid purification method is provided, comprising a pretreatment system, a first extraction system, a desulfurization treatment system, a washing and stripping system, a second extraction system and a regeneration system.

[0050] In some embodiments, the first extraction system includes an extraction tank, an emulsion breaking and adjustment tank, and a phase separation tank;

[0051] And / or, the desulfurization treatment system includes a solution preparation system, a phase separation system, a first desulfurization system, and a second desulfurization system;

[0052] And / or, the washing and stripping system includes a washing tower and a stripping tower;

[0053] And / or, the second extraction system includes an extraction tank, a phase separation tank and a stripping tank;

[0054] And / or, the regeneration system includes a regeneration tank, a phase separation tank and a water washing tower.

[0055] Compared with the prior art, the present invention has the following beneficial effects:

[0056] 1. The present invention performs secondary extraction on the first raffinate produced by extraction, returns the raffinate to be mixed with wet-process phosphoric acid, and returns the washed raffinate to the desulfurization treatment step for recycling to form an acid closed loop. The entire wet-process phosphoric acid purification process does not produce acid byproducts, is environmentally friendly, and the yield of the purified phosphoric acid product is greater than 95%, eliminating the need to worry about the disposal of acid byproducts.

[0057] 2. The present invention uses a mixture of tributyl phosphate, sulfonated kerosene, and tridecanol as the first extractant. Tributyl phosphate has a good removal effect on phosphate cations, is safe, non-toxic, and has a low water solubility, making it a good phosphoric acid extractant. The combination of tributyl phosphate and sulfonated kerosene can improve the fluidity of the first extractant, increase the extraction mass transfer rate, and improve the extraction efficiency. However, when the mixture of tributyl phosphate and sulfonated kerosene is at low temperatures or when the phosphoric acid extraction rate is high, the kerosene easily precipitates from the extraction phase to form a third phase, thereby affecting the extraction effect of the first extractant. The present invention adds tridecanol as a modifier to change the polarity between the first extractants, increase the compatibility between organic matter, prevent the formation of the third phase, and improve the extraction rate and extraction efficiency of the first extractant.

[0058] 3. The present invention adopts three precipitation desulfurization processes. The first desulfurization process is the desulfurization process in the pretreatment process. The pretreatment process uses phosphate rock powder for desulfurization. The phosphate rock powder is widely available and has a low desulfurization cost. The desulfurization rate in the pretreatment process is 60-80%. The second desulfurization process uses a calcium salt solution for desulfurization, and the desulfurization rate can reach more than 90%. The third desulfurization process uses a barium salt solution for desulfurization, and the desulfurization rate can reach more than 99%. The three precipitation desulfurization processes all use an acid phase (washing residual acid or desulfurizing) to prepare a solution and then react with the phosphoric acid phase, that is, they are all liquid-liquid contact reactions, which improves the desulfurization efficiency and desulfurization rate. In addition, the introduced impurities are carried away by the acid phase, thereby improving the purification effect of the phosphoric acid. In addition, the extract phase is desulfurized before washing, achieving a pre-washing effect and improving the purity of the phosphoric acid product.

[0059] 4. The present invention first mixes the first extraction phase and the second desulfurization phase in the desulfurization process and performs phase separation, which has the following advantages: first, the calcium ion concentration in the second desulfurization phase is high, which can achieve the pre-desulfurization effect of the first extraction phase, improve the utilization rate of calcium salt, and reduce the cost of calcium salt; second, the calcium ions in the first desulfurization phase are reduced, avoiding the formation of precipitation during the pretreatment process of the first desulfurization phase returning to step S1, resulting in blockage of the pretreatment system structure; third, if the second desulfurization phase is directly returned to the pretreatment process, the calcium content of the finished phosphoric acid will be too high, making the quality of the finished phosphoric acid unqualified. BRIEF DESCRIPTION OF THE DRAWINGS

[0060] Figure 1 This is a process flow chart of the wet-process phosphoric acid purification method disclosed herein;

[0061] Figure 2 This is a process flow diagram of the first extraction system in the purification system used in the wet-process phosphoric acid purification method disclosed herein;

[0062] Figure 3 This is a process flow chart of a desulfurization system in a purification system used in the wet-process phosphoric acid purification method disclosed herein;

[0063] Figure 4 This is a process flow chart of a washing and stripping system in a purification system used in the wet-process phosphoric acid purification method disclosed herein;

[0064] Figure 5 A process flow chart of a regeneration system in a purification system for the wet-process phosphoric acid purification method disclosed herein;

[0065] Figure 6 This is a process flow diagram of the second extraction system in the purification system used in the wet-process phosphoric acid purification method disclosed herein. DETAILED DESCRIPTION

[0066] To facilitate understanding of the present disclosure, a more comprehensive description will be given below. However, the present disclosure can be implemented in many different forms and is not limited to the embodiments described herein. Rather, the purpose of providing these embodiments is to provide a more thorough and comprehensive understanding of the disclosure of the present disclosure.

[0067] As used herein:

[0068] "Prepared from" is synonymous with "comprising." As used herein, the terms "comprising," "including," "having," "containing," or any other variations thereof, are intended to cover a non-exclusive inclusion. For example, a composition, process, method, article, or apparatus that comprises the listed elements is not necessarily limited to only those elements but may include other elements not expressly listed or inherent to such composition, process, method, article, or apparatus.

[0069] The conjunction "consisting of" excludes any unspecified element, step, or component. If used in a claim, this phrase renders the claim closed, excluding materials other than those described, except for conventional impurities associated therewith. When the phrase "consisting of" appears in a clause of the body of a claim, rather than immediately following the subject matter, it limits only the elements described in that clause; other elements are not excluded from the claim as a whole.

[0070] When amount, concentration or other value or parameter are represented with range, preferred range or the range that a series of upper preferred value and lower preferred value limit are expressed, this should be understood as specifically disclosing all ranges formed by any pairing of any range upper limit or preferred value and any range lower limit or preferred value, and no matter whether this scope is disclosed separately.For example, when disclosing scope " 1-5 ", described scope should be interpreted as including scope " 1-4 ", " 1-3 ", " 1-2 ", " 1-2 and 4-5 ", " 1-3 and 5 " etc.When numerical range is described in this article, unless otherwise stated, otherwise this scope is intended to include its end value and all integers and fractions within this range.

[0071] In these examples, parts and percentages are by mass unless otherwise indicated.

[0072] "Parts by mass" refers to the basic unit of measurement used to express the mass ratio of multiple components. One part can represent any unit of mass, such as 1g or 2.689g. If we say that the mass of component A is a parts and the mass of component B is b parts, this means the ratio of the mass of component A to the mass of component B is a:b. Alternatively, we could say that the mass of component A is aK and the mass of component B is bK (K is an arbitrary number representing a multiplication factor). It's important to note that, unlike parts by mass, the sum of the mass of all components is not limited to 100 parts.

[0073] "And / or" is used to indicate that one or both of the stated situations may occur, for example, A and / or B includes (A and B) and (A or B).

[0074] In order to improve the utilization rate of the raffinate acid generated by the purification of wet-process phosphoric acid and the extraction rate of wet-process phosphoric acid, Figure 1 As shown, the present disclosure provides a method for purifying wet-process phosphoric acid, comprising the following steps:

[0075] S1: pretreatment;

[0076] Specifically, the pre-processing steps are:

[0077] S11: Wet-process phosphoric acid and the second raffinate acid are mixed to obtain first crude phosphoric acid.

[0078] Specifically, the mass fraction of P2O5 in the wet-process phosphoric acid is 45-50%, for example, but not limited to 45%, 46%, 47%, 48%, 49%, 50%;

[0079] Specifically, the mass ratio of the wet-process phosphoric acid to the second raffinate acid is 1-2, for example, but not limited to 1, 1.2, 1.4, 1.6, 1.8, or 2.

[0080] S12: adding a desulfurization and defluorination agent to the first crude phosphoric acid to carry out a desulfurization and defluorination reaction to obtain a first slurry.

[0081] Specifically, the desulfurization and defluorination agent includes phosphate rock powder, sodium carbonate, silicon dioxide and desulfurization, and the desulfurization is the first desulfurization and / or the third desulfurization;

[0082] Specifically, the stoichiometric ratio of the phosphate rock in the desulfurization and defluorination agent to the sulfate in the first crude phosphoric acid is 1.05-1.2; for example, it can be but not limited to 1.05, 1.1, 1.15, 1.2;

[0083] Specifically, the stoichiometric ratio of the total stoichiometric amount of sodium carbonate and silicon dioxide in the desulfurization and defluorination agent to the fluoride ion in the first crude phosphoric acid is 1.2-1.4, for example, but not limited to 1.2, 1.25, 1.3, 1.35, 1.4;

[0084] Specifically, the mass ratio of sodium carbonate to silicon dioxide is 2-4, for example, but not limited to 2, 2.5, 3, 3.5, 4.

[0085] S13: adding a dearsenicizing agent to the first slurry to perform a dearsenicizing reaction, and filtering the obtained product using a filter press to obtain a second slurry.

[0086] Specifically, the arsenic removal agent comprises a mixture of a sulfur-containing compound and desulfurization to obtain the arsenic removal agent, the sulfur-containing compound is sodium sulfide or phosphorus pentasulfide, and the desulfurization is a first desulfurization and / or a third desulfurization;

[0087] Specifically, the stoichiometric ratio of sulfur atoms in the arsenic removal agent to arsenic ions in the first crude phosphoric acid is 10-20, for example, but not limited to 10, 12, 14, 16, 18, 20.

[0088] S14: adding the second slurry to a dechlorination tower, introducing ozone to dechlorinate, and obtaining a third slurry.

[0089] Specifically, the volume ratio of the ozone to the second slurry is 500-1000, for example, but not limited to 500, 600, 700, 800, 900, 1000;

[0090] Specifically, the negative pressures of the dearsenicization reaction and the dechlorination reaction are each independently 0.3 kPa-1.5 kPa, for example, but not limited to 0.3 kPa, 0.5 kPa, 0.7 kPa, 0.9 kPa, 1.1 kPa, 1.3 kPa, 1.5 kPa;

[0091] The dearsenicization reaction will produce hydrogen sulfide gas, and the dechlorination reaction will produce ozone gas, both of which are toxic gases. The present invention adopts a negative pressure machine to carry out dearsenicization and dechlorination under the above pressure, thereby preventing the toxic gas from overflowing and transferring the toxic gas to the tail gas receiving device and the washing device, thereby improving the safety of the pretreatment process.

[0092] S15: adding activated carbon to the third slurry for decolorization, filtering the obtained product to obtain a fourth slurry, and concentrating the fourth slurry to obtain crude phosphoric acid with a P2O5 mass fraction of 48%-50%.

[0093] Specifically, the mass of the activated carbon is 0.5%-1.0% of the mass of the third slurry, for example, it can be but not limited to 0.5%, 0.6%, 0.7%, 0.8%, 0.9%, 1.0%.

[0094] Specifically, the temperature of the desulfurization and defluorination reaction, the dearsenicization reaction, the dechlorination, and the decolorization reaction is independently 50-60°C, for example, but not limited to 50°C, 52°C, 54°C, 56°C, 58°C, and 60°C;

[0095] Specifically, the time for the desulfurization and defluorination reaction, the dearsenicization reaction and the decolorization is independently 1-2 hours.

[0096] The present invention removes some impurities in wet-process phosphoric acid through the above-mentioned pretreatment step, and can remove impurities introduced during the pretreatment process through the subsequent extraction step without affecting the quality of the finished phosphoric acid.

[0097] The present invention uses the by-product of desulfurization to prepare a chemical impurity removal agent used in a pretreatment process, which not only ensures continuous and stable feeding, but also allows the chemical impurity removal agent and wet-process phosphoric acid to undergo a liquid-liquid contact reaction, thereby improving the reaction efficiency of the impurity removal reaction. At the same time, the by-product of desulfurization can be recycled, reducing the content of by-product acid, which is beneficial to environmental protection.

[0098] S2: extraction;

[0099] Specifically, the extraction steps are:

[0100] S21: performing multi-stage countercurrent extraction on the crude phosphoric acid obtained in step S15 and the first extractant to obtain a first extract phase and a first raffinate acid;

[0101] Specifically, the volume ratio of the first extractant to the crude phosphoric acid is 4-6, for example, but not limited to 4, 4.2, 4.5, 4.7, 5, 5.3, 5.5, 5.8, 6;

[0102] Specifically, the number of extraction stages of the multi-stage countercurrent extraction is 4-6 stages; for example, it can be 4 stages, 5 stages, or 6 stages; the contact time of each stage of extraction is 10-15 minutes, for example, it can be but not limited to 10 minutes, 13 minutes, or 15 minutes;

[0103] Specifically, the temperature of the multi-stage countercurrent extraction is 40-50°C, for example, but not limited to 40°C, 42°C, 44°C, 46°C, 48°C, 50°C.

[0104] Specifically, the first extractant is a mixture of tributyl phosphate, sulfonated kerosene and tridecanol, wherein the volume ratio of tributyl phosphate: sulfonated kerosene: tridecanol is 60-75:20-30:5-10; for example, it can be but not limited to 60:30:10, 65:25:10, 70:23:7, 75:20:5;

[0105] The present invention uses a mixture of tributyl phosphate, sulfonated kerosene, and tridecanol as a first extractant. Tributyl phosphate has a good removal effect on phosphate cations, is safe, non-toxic, and has a low water solubility, making it a good phosphoric acid extractant. The combination of tributyl phosphate and sulfonated kerosene can improve the fluidity of the first extractant, increase the extraction mass transfer rate, and improve the extraction efficiency. However, when the mixture of tributyl phosphate and sulfonated kerosene is at low temperatures or when the phosphoric acid extraction rate is high, the kerosene easily precipitates from the extraction phase to form a third phase, thereby affecting the extraction effect of the first extractant. The present invention adds tridecanol as a modifier to change the polarity between the first extractants, increase the compatibility between organic substances, prevent the formation of the third phase, and improve the extraction rate and extraction efficiency of the first extractant.

[0106] S22: adding hydrogen peroxide to the first raffinate acid for oxidation reaction to obtain an oxidized first raffinate acid; wherein the amount of hydrogen peroxide added is 0.1%-0.5% of the mass of the first raffinate acid, for example, but not limited to 0.1%, 0.2%, 0.3%, 0.4%, or 0.5%;

[0107] Specifically, the temperature of the oxidation reaction is 60-70°C, for example, but not limited to, 60°C, 62°C, 64°C, 66°C, 68°C, or 70°C.

[0108] Specifically, the oxidation reaction time is 30-60 min, for example, but not limited to 30 min, 35 min, 40 min, 45 min, 50 min, 55 min, 60 min;

[0109] S23: subjecting the oxidized first raffinate acid and the second extractant to multi-stage cross-current extraction to obtain a second extract phase and a second raffinate acid, and returning the second raffinate acid to step S11 to be mixed with wet-process phosphoric acid;

[0110] Specifically, the volume ratio of the second extractant to the oxidized first raffinate acid is 5-3:1, for example, but not limited to, 5:1, 4.5:1, 4:1, 3.5:1, 3:1;

[0111] Specifically, the number of stages of the multi-stage cross-flow extraction is 3-5, for example, 3, 4, or 5, and the extraction time of each stage is 20-40 min, for example, but not limited to 20 min, 25 min, 30 min, 35 min, or 40 min;

[0112] Specifically, the temperature of the multi-stage cross-current extraction is 40-60°C, for example, but not limited to 40°C, 45°C, 50°C, 55°C, and 60°C.

[0113] S24: After washing the second extract phase with dilute phosphoric acid, stripping with 30%-40% by mass sulfuric acid to obtain a metal stripping solution and a metal-organic phase, and returning the metal-organic phase to step S23 as the second extractant;

[0114] Specifically, the preparation method of the second extractant is as follows: after mixing the extractant and sulfonated kerosene, adding sodium hydroxide solution to carry out saponification reaction, phase separation, and the obtained organic phase is the second extractant;

[0115] The extractant is at least one of 2-ethylhexyl phosphate (P507), di(2-ethylhexyl) phosphate (P204), dinonylnaphthalenesulfonic acid (DNNSA), cyclohexane acid, trialkylphosphine oxide, and acetylacetone.

[0116] The saponification rate of the second extractant is 10%-30%, for example, but not limited to 10%, 12%, 15%, 17%, 20%, 23%, 25%, 28%, 30%;

[0117] The volume ratio of the extractant to sulfonated kerosene is 1-2.5, for example, but not limited to 1, 1.2, 1.5, 1.7, 2, 2.3, 2.5;

[0118] The mass fraction of the sodium hydroxide solution is less than 15%. The mass fraction of the sodium hydroxide solution within the above range is conducive to the saponification reaction between the second extractant and the sodium hydroxide solution.

[0119] Specifically, the saponification rate of the second extractant is calculated as follows:

[0120] Saponification rate = n1 / n2×100%,

[0121] n1=ρ1V1w / M1,

[0122] n2=ρ2V2 / M2,

[0123] n1-amount of sodium hydroxide substance mol; n2-amount of extractant substance mol; ρ1-density of sodium hydroxide solution g / l; V1-volume of sodium hydroxide solution l; w-mass fraction of sodium hydroxide solution; M1-molar mass of sodium hydroxide g / mol; ρ2-density of extractant g / l; V2-volume of extractant l; M2-molar mass of extractant g / mol.

[0124] Specifically, the mass fraction of the dilute phosphoric acid is 1%-10%, for example, but not limited to 1%, 3%, 5%, 7%, and 10%.

[0125] Specifically, the stripping temperature is 50-60°C, for example, but not limited to, 50°C, 52°C, 55°C, 57°C, and 60°C.

[0126] Specifically, the stripping time is 40-60 min, for example, but not limited to 40 min, 45 min, 50 min, 55 min, 60 min;

[0127] Specifically, the number of stripping stages is 2-4, for example, but not limited to 2, 3, and 4.

[0128] In the present disclosure, after the first raffinate acid is subjected to secondary extraction, the secondary extraction can not only remove metal impurities in the first raffinate acid, but also reduce phosphorus loss in the secondary extraction. The second raffinate acid obtained by the secondary extraction is returned to step S11 and mixed with wet-process phosphoric acid, and the acid is recycled, which can improve the yield of phosphoric acid. In addition, after the second extraction phase produced by the secondary extraction is reprocessed, the metal impurities in the second extraction phase can be separated, which is beneficial to the subsequent treatment of the metal impurities and the reuse of the second extraction phase.

[0129] S3: Desulfurization treatment; specifically, the steps of desulfurization treatment are:

[0130] S31: The first extraction phase and the second desulfurization phase of step S21 are mixed to perform a first reaction, phase separation is performed to obtain a first organic phase and a first desulfurization phase, and the first desulfurization phase is returned to the pretreatment step to prepare a desulfurization and defluorination agent and / or a dearsenicizing agent;

[0131] S32: mixing the first organic phase and the calcium salt solution and performing a second reaction, separating the phases to obtain a second organic phase and a second desulfurization;

[0132] S33: mixing the second organic phase and the barium salt solution and performing a third reaction, separating the phases to obtain a desulfurized organic phase and a third desulfurization phase; and returning the third desulfurization phase to the pretreatment step to prepare a desulfurization, defluorination agent and / or a dearsenicizing agent;

[0133] Specifically, the calcium salt solution includes calcium salt and washing residual acid, and the barium salt solution includes barium salt and washing residual acid, wherein the calcium salt is carbonate and / or calcium hydroxide, and the barium salt is barium carbonate;

[0134] Specifically, the temperature of the first reaction, the second reaction and the third reaction is independently 40-50°C, for example, but not limited to 40°C, 42°C, 44°C, 46°C, 48°C, 50°C;

[0135] Specifically, the time for the first reaction, the second reaction and the third reaction is independently 30-60 min, for example, but not limited to 30 min, 35 min, 40 min, 45 min, 50 min, 55 min, 60 min;

[0136] Specifically, the volume ratio of the first extraction phase to the second desulfurization phase is 10-15, the volume ratio of the first organic phase to the calcium salt solution is 10-15, and the volume ratio of the second organic phase to the barium salt solution is 10-15; for example, it can be but not limited to 10, 11, 12, 13, 14, 15.

[0137] In step S31, the second desulfurization phase is mixed with the first extraction phase. On the one hand, the second desulfurization phase contains a relatively high concentration of calcium ions, which can pre-desulfurize the first extraction phase, fully utilize the calcium ions, improve the utilization rate of the calcium salt, and reduce the cost of the calcium salt. On the other hand, it can reduce the calcium ions in the first desulfurization phase, avoiding the formation of precipitation during the first desulfurization phase returning to the pretreatment process of step S1, which may lead to blockage of the pretreatment system structure. On the other hand, if the second desulfurization phase is directly returned to the pretreatment process of step S1, the calcium content of the finished phosphoric acid will be too high, resulting in substandard quality of the finished phosphoric acid.

[0138] Steps 12, S32, and S33 in the present disclosure constitute three precipitation desulfurization steps in the wet-process phosphoric acid purification process. The desulfurization rate of step 12 is 60-80%, the desulfurization rate of step S32 can reach over 90%, and the desulfurization rate of step S33 can reach over 99%. Steps S32 and S33 use calcium salt solution and barium salt solution respectively for desulfurization. First, the calcium salt solution and barium salt solution react with the organic phase in liquid-liquid contact, resulting in a fast desulfurization reaction and high desulfurization efficiency. In addition, the introduced impurities are all carried away by the acid phase, thereby improving the purity of the finished phosphoric acid. Second, the calcium salt solution and barium salt solution perform desulfurization, which can achieve a pre-washing effect on the extract phase, further improving the purity of the finished phosphoric acid. In addition, the present disclosure uses calcium carbonate and / or calcium hydroxide as calcium salt, which not only has a good desulfurization effect but also has low cost. Combining with barium carbonate for desulfurization can improve the desulfurization effect.

[0139] S4: washing: countercurrently washing the desulfurized organic phase with washing acid to obtain a washed organic phase and a washing residual acid, and returning the washing residual acid to the desulfurization step to prepare a desulfurizer;

[0140] Specifically, the volume ratio of the desulfurized organic phase to the washing acid is 12-18, for example, but not limited to 12, 13, 14, 15, 16, 17, 18;

[0141] Specifically, the temperature of the countercurrent washing is 50-60°C, for example, but not limited to, 50°C, 52°C, 54°C, 56°C, 58°C, and 60°C;

[0142] Specifically, the number of countercurrent washing stages is 2-3;

[0143] S5: Stripping: Specifically, the stripping steps are as follows:

[0144] S51: stripping the washed organic phase with a stripping agent to obtain stripping acid and a stripping organic phase, and returning part of the stripping acid to step S4 as a washing acid;

[0145] Specifically, the stripping agent is desalted water;

[0146] Specifically, the number of stages of stripping is 2-4;

[0147] Specifically, the volume ratio of the washed organic phase to the stripping agent is 4-6; for example, it can be but not limited to 4, 4.5, 5, 5.5, 6;

[0148] Specifically, the stripping temperature is 60-70°C; for example, it can be but not limited to 60°C, 62°C, 64°C, 66°C, 68°C, or 70°C.

[0149] In the present disclosure, the stripping acid is phosphoric acid having a certain concentration and a low impurity content. The use of the stripping acid for washing can not only achieve the same washing effect as water, but also introduces much less water into the system than water washing, thereby ensuring the concentration of the washed organic phase and reducing the difficulty and energy consumption of subsequent concentration. Specifically, the source of the stripping acid is unconcentrated phosphoric acid.

[0150] S52: The stripped organic phase is washed with a sodium hydroxide aqueous solution, then activated carbon is added for adsorption, and then phase separation is performed. The resulting organic phase is washed with desalted water to obtain a regenerated organic phase. After adding an appropriate amount of the first extractant to the regenerated organic phase, the process returns to step S21 to perform the first extraction using the first extractant;

[0151] Specifically, the volume ratio of the stripping organic phase to the sodium hydroxide aqueous solution is 1-3, for example, but not limited to 1, 1.5, 2, 2.5, 3;

[0152] Specifically, the mass fraction of the sodium hydroxide aqueous solution is 4%-6%, for example, it can be but not limited to 4%, 4.5%, 5%, 5.5%, 6%;

[0153] Specifically, the amount of activated carbon added is 0.5%-1% of the mass of the stripping organic phase, for example, it can be but not limited to 0.5%, 0.6%, 0.7%, 0.8%, 0.9%, 1.0%;

[0154] After long-term use, the first extractant will produce decomposition products and interphase contaminants. These two substances will increase the difficulty of phase separation during extraction, affect the quality of subsequent products and increase the loss of the extractant. By regenerating and removing them, a better extraction effect can be maintained, and the stripping organic phase can be recycled and reused to obtain better economic value.

[0155] The present invention recycles the desulfurization, residual acid washing and stripping acid in the by-products to form an acid closed loop. No by-product acid is produced during the entire wet-process phosphoric acid purification process, eliminating the need to worry about by-product acid treatment. This is not only environmentally friendly, but also has a high yield of phosphoric acid product (>95%), which is at least 40% higher than the yield of phosphoric acid product (54%) obtained by the traditional wet-process phosphoric acid purification method.

[0156] S6: Concentration and final treatment; specifically, the steps of concentration and final treatment are:

[0157] S61: The remaining stripping acid obtained in step S51 is concentrated for the first time to obtain phosphoric acid with a mass fraction of 60%, then concentrated for the second time to obtain phosphoric acid with a mass fraction of 75%, and then concentrated for the third time to obtain phosphoric acid with a mass fraction of 85%;

[0158] S62: Add activated carbon and hydrogen peroxide to 85% by mass phosphoric acid for decolorization, and filter to obtain a phosphoric acid product.

[0159] The present disclosure also provides a purification system for the wet-process phosphoric acid purification method, comprising a pretreatment system, a first extraction system, a desulfurization treatment system, a washing and stripping system, a second extraction system, and a regeneration system.

[0160] Specifically, the first extraction system includes an extraction tank, an emulsion breaking and regulating tank, and a phase separation tank;

[0161] like Figure 2 As shown, taking the number of extraction stages as 4 as an example, the first extraction system includes a primary extraction tank, a primary demulsification adjustment tank, a primary phase separation tank, a secondary extraction tank, a secondary demulsification adjustment tank, a secondary phase separation tank, a tertiary extraction tank, a tertiary demulsification adjustment tank, a tertiary phase separation tank, a quaternary extraction tank, a quaternary demulsification adjustment tank, a quaternary phase separation tank, a raffinate acid mixing tank, a raffinate acid settling tank, and an extractant collecting tank;

[0162] The crude phosphoric acid is added to the primary extraction tank to start the primary extraction. After the primary extraction is completed, the mixed slurry enters the primary demulsification adjustment tank for primary demulsification. After the primary demulsification is completed, the mixed slurry enters the primary phase separation tank for primary phase separation. The first extraction phase in the primary phase separation tank enters the desulfurization treatment system.

[0163] The acid phase in the primary phase separation tank is partially refluxed to the primary demulsification adjustment tank, and the remaining portion enters the secondary extraction tank for secondary extraction; after the secondary extraction, the resulting mixed slurry enters the secondary demulsification adjustment tank for secondary demulsification. After the secondary demulsification is completed, the resulting mixed slurry enters the secondary phase separation tank for secondary phase separation, and the organic phase in the secondary phase separation tank enters the primary extraction tank for primary extraction. The acid phase in the secondary phase separation tank is partially refluxed to the secondary demulsification adjustment tank, and the remaining portion enters the tertiary extraction tank for tertiary extraction;

[0164] After the third stage extraction, the mixed slurry enters the third stage demulsification adjustment tank for third stage demulsification. After the third stage demulsification, the mixed slurry enters the third stage phase separation tank for third stage phase separation. The organic phase in the third stage phase separation tank enters the second stage extraction tank for second stage extraction. The acid phase in the third stage phase separation tank is partially refluxed to the third stage demulsification adjustment tank, and the remaining part enters the fourth stage extraction tank for fourth stage extraction.

[0165] After the fourth-stage extraction, the mixed slurry enters the fourth-stage demulsification adjustment tank for fourth-stage demulsification. After the fourth-stage demulsification, the mixed slurry enters the fourth-stage phase separation tank for fourth-stage phase separation. The organic phase in the fourth-stage phase separation tank enters the third-stage extraction tank for third-stage extraction. The acid phase in the fourth-stage phase separation tank is partially refluxed to the fourth-stage demulsification adjustment tank, and the remaining part enters the raffinate acid mixing tank for mixing. The mixed liquid obtained in the raffinate acid mixing tank enters the raffinate acid settling tank, and the organic phase in the raffinate acid settling tank enters the extractant collecting tank. The extractant in the extractant collecting tank is circulated back to the raffinate acid mixing tank. The acid phase in the raffinate acid settling tank is the first raffinate acid, and a secondary extraction process is performed.

[0166] Specifically, the viscosity of the extraction phase is high, the mass transfer is slow, and the phase separation is slow. The extraction tank and demulsification adjustment tank disclosed in the present invention use a double-layer screen agitator for stirring, which fully disperses the system into droplets. The oil-water phase contact area is large, the mass transfer efficiency is high, and different stirring speeds are used to achieve the effects of emulsification and demulsification, thereby increasing the extraction efficiency and phase separation efficiency.

[0167] Furthermore, the stirring speed of the extraction tank is 80-120 r / min, and the stirring speed of the demulsification adjustment tank is 20-60 r / min;

[0168] During each extraction process, part of the acid phase in the phase separation tank is returned to the demulsification adjustment tank, which can control the volume ratio of the organic phase to the aqueous phase in the demulsification adjustment tank to 2-3, which is beneficial to demulsification and phase separation. For example, when the phase ratio in the demulsification adjustment tank is too large and phase separation is difficult, the returned part of the acid phase increases the contact area of ​​the acid phase, improves phase separation, reduces the acid entrained in the extraction phase, and is beneficial to improving the purity of the phosphoric acid product.

[0169] Specifically, the desulfurization treatment system includes a solution preparation system, a phase separation system, a first desulfurization system, and a second desulfurization system;

[0170] like Figure 3 As shown, the first extraction phase is added to the decalcification reaction tank for decalcification reaction, and the slurry obtained by the decalcification reaction enters the decalcification sedimentation tank for phase separation. The acid clear liquid in the decalcification sedimentation tank is returned to the decalcification reaction tank; the acid residue slurry in the decalcification sedimentation tank enters the decalcification thickening tank for thickening, and then enters the decalcification filter press for filtration. The obtained decalcification filtrate is the first desulfurization, part of the first desulfurization is returned to the pretreatment section for extraction, and the remaining part of the first desulfurization enters the calcium salt preparation tank to prepare the calcium salt solution;

[0171] The organic phase in the decalcification sedimentation tank enters the crude desulfurization reaction tank and is mixed with the calcium salt solution to perform the first desulfurization reaction. The slurry obtained from the first desulfurization reaction enters the crude desulfurization sedimentation tank for phase separation. The acid clear liquid in the crude desulfurization sedimentation tank is returned to the crude desulfurization reaction tank for the first desulfurization reaction. The acid slag slurry in the crude desulfurization sedimentation tank enters the crude desulfurization thickening tank for thickening. The second desulfurization in the crude desulfurization thickening tank enters the decalcification reaction tank for decalcification reaction.

[0172] The organic phase and barium salt solution in the coarse desulfurization sedimentation tank enter the fine desulfurization reaction tank for the second desulfurization reaction. The slurry obtained from the second desulfurization reaction enters the fine desulfurization sedimentation tank for phase separation. The acid clear liquid in the fine desulfurization sedimentation tank is returned to the fine desulfurization reaction tank for the second desulfurization reaction; the organic phase in the fine desulfurization sedimentation tank enters the washing and stripping section; the acid slag slurry in the fine desulfurization sedimentation tank enters the fine desulfurization thickening tank for thickening, and the third desulfurization in the fine desulfurization thickening tank enters the fine desulfurization membrane filtration and concentration equipment for filtration. The obtained filtrate enters the barium salt preparation tank and the calcium salt preparation tank respectively, and the obtained precipitate enters the decalcification thickening tank after concentration.

[0173] During the desulfurization process, part of the acid phase in the coarse desulfurization sedimentation tank of the desulfurization section is refluxed to the coarse desulfurization reaction tank, and part of the acid phase in the fine desulfurization sedimentation tank is refluxed to the fine desulfurization reaction tank, which has the following advantages: first, the phase ratio in the desulfurization reaction tank is reduced, ensuring that the volume ratio of the organic phase and the aqueous phase in the desulfurization reaction tank is 3-5, which is beneficial to the subsequent phase separation; second, the reflux is increased to control the supersaturation of the crystals, which is beneficial to the growth of calcium sulfate and barium sulfate crystals, without scaling and easy to filter; third, the acid in the preparation of calcium salt solution and barium salt solution is reduced, the acid circulation preparation of calcium salt solution and barium salt solution is reduced, and the operation steps are reduced.

[0174] Specifically, the washing and stripping system includes a washing tower and a stripping tower;

[0175] like Figure 4 As shown, taking the countercurrent washing stage as 2 and the countercurrent stripping stage as 2 as an example, the washing and stripping system includes a primary washing tower, a secondary washing tower, a washing ester phase buffer tank, a primary stripping tower, a secondary stripping tower, and a stripping acid buffer tank;

[0176] The fine desulfurized organic phase is added to the primary washing tower, and the residual acid obtained from the secondary washing is used for primary washing. The residual acid obtained from the primary washing is respectively fed into the calcium salt preparation tank and the barium salt preparation tank of the desulfurization process;

[0177] The ester phase obtained from the primary washing enters the secondary washing tower and is washed with stripping acid for secondary washing. The acid phase obtained from the secondary washing returns to the primary washing tower for primary washing.

[0178] The ester phase obtained from the secondary washing enters the primary stripping tower after passing through the ester washing buffer tank, and undergoes primary stripping with the acid obtained from the secondary stripping. The acid phase obtained from the primary stripping is collected in the stripping acid buffer tank, and part of the stripping acid enters the washing tower as washing acid, and the remaining stripping acid enters the concentration step;

[0179] The ester phase obtained from the primary stripping enters the secondary stripping tower for secondary stripping with desalted water. The stripped organic phase obtained from the secondary stripping enters the first extraction section. The acid phase obtained from the secondary stripping enters the primary stripping tower for primary stripping.

[0180] Specifically, the washing tower and the stripping tower are vibrating sieve plate towers. During the washing process, the organic phase is the continuous phase and the acid phase is the dispersed phase; during the stripping process, the organic phase is the dispersed phase and the acid phase is the continuous phase. In order to improve the phase separation effect, the vibration frequency of the washing tower is 15-25Hz, and the vibration frequency of the stripping tower is 20-30Hz. The viscosity of the fine desulfurization organic phase is low, and the mass transfer is mainly affected by the concentration driving force. A vibrating sieve plate tower is used for washing and stripping. Increasing the number of tower plates and countercurrent contact can increase the washing and stripping effects.

[0181] Specifically, the regeneration extraction system includes an extraction regeneration tank, a regeneration phase separation tank and a water washing tower;

[0182] like Figure 5 As shown, the stripping organic phase and alkali solution obtained by stripping are added to the extraction regeneration tank A for reaction, the reaction product is added to the extraction regeneration tank B to adsorb the first extractant decomposition product with activated carbon, and then enters the filter press for filtration, the filtrate obtained by filtration enters the regeneration phase separation tank for phase separation, and the organic phase obtained by phase separation enters the water washing tower for washing with desalted water to separate and obtain the regenerated extractant and the washing alkali solution; the regenerated extractant is returned to the first extraction tank for the first extraction.

[0183] Specifically, the second extraction system includes an extraction tank, a phase separation tank and a stripping tank;

[0184] like Figure 6 As shown, taking the extraction stages as 3 and the stripping stages as 2 as an example, the second extraction system includes a primary extraction tank, a primary phase separation tank, a secondary extraction tank, a secondary phase separation tank, a tertiary extraction tank, a tertiary phase separation tank, a primary stripping tank, a primary stripping phase separation tank, a secondary stripping tank, a secondary stripping phase separation tank, a washing tank and a washing phase separation tank;

[0185] The first raffinate is added to the primary extraction tower and subjected to primary extraction with the second extractant. The slurry obtained from the primary extraction enters the primary phase separation tank for primary phase separation. The organic phase in the primary phase separation tank is collected in a metal ion-loaded extraction phase buffer tank.

[0186] The acid phase in the primary phase separation tank enters the secondary extraction tank for secondary extraction, and the slurry obtained from the secondary extraction enters the secondary phase separation tank for secondary phase separation. The organic phase in the secondary phase separation tank is collected in the metal ion loaded extraction phase buffer tank.

[0187] The acid phase in the secondary phase separation tank enters the tertiary extraction tank for tertiary extraction. The slurry obtained by the tertiary extraction enters the tertiary phase separation tank for tertiary phase separation. The organic phase in the tertiary phase separation tank is collected in a metal ion-loaded extraction phase buffer tank, and the acid phase in the tertiary phase separation tank is collected in a raffinate storage tank.

[0188] The organic phase collected in the metal ion-loaded extraction phase buffer tank is added to the washing tank and washed with dilute phosphoric acid. The mixture in the washing tank enters the washing phase separation tank for phase separation. The organic phase in the washing phase separation tank enters the primary stripping tank for primary stripping with sulfuric acid, and then enters the primary stripping phase separation tank. The organic phase in the primary stripping phase separation tank enters the secondary stripping tank for secondary stripping with sulfuric acid, and then enters the secondary stripping phase separation tank. The organic phase in the secondary stripping phase separation tank is returned to the extraction step as the second extractant.

[0189] In the following examples and comparative examples, the preparation methods of the used medicaments are as follows:

[0190] Desulfurization and defluorination agent: prepared by mixing phosphate rock powder, sodium carbonate, silicon dioxide and desulfurization, wherein the desulfurization is the first desulfurization and / or the third desulfurization;

[0191] Arsenic removal agent: prepared by mixing sodium sulfide and desulfurization, wherein the desulfurization is the first desulfurization and / or the third desulfurization;

[0192] Calcium salt solution: Mix calcium carbonate with residual washing acid;

[0193] Barium salt solution: Mix barium carbonate with residual washing acid;

[0194] The first extractant is prepared by uniformly mixing tributyl phosphate, sulfonated kerosene and tridecanol;

[0195] Second extractant: Mix the extractant and sulfonated kerosene evenly, then saponify with sodium hydroxide solution, separate the phases, and the resulting organic phase is the second extractant;

[0196] The components of wet-process phosphoric acid are shown in Table 1 below:

[0197] Table 1

[0198] Element <![CDATA[P2O5]]> F Cl <![CDATA[SO4]]> Fe Mg Al Ca As Pb content% 47.8 0.75 0.025 4.20 0.52 0.95 0.69 0.1 0.003 0.002

[0199] Example 1

[0200] This embodiment provides a method for purifying wet-process phosphoric acid and a purification system for the wet-process phosphoric acid purification method. The details are as follows:

[0201] S1: pretreatment;

[0202] S11: mixing wet-process phosphoric acid and the second raffinate acid at a mass ratio of 1.5:1 to obtain first crude phosphoric acid;

[0203] S12: adding a desulfurization and defluorination agent to the first crude phosphoric acid to carry out a desulfurization and defluorination reaction at a temperature of 50° C. for 1 hour to obtain a first slurry; wherein the stoichiometric ratio of the phosphate rock in the desulfurization and defluorination agent to the sulfate in the first crude phosphoric acid is 1.05, the total stoichiometric ratio of sodium carbonate and silicon dioxide in the desulfurization and defluorination agent to the stoichiometric ratio of fluoride ions in the first crude phosphoric acid is 1.2, and the mass ratio of sodium carbonate to silicon dioxide is 2;

[0204] S13: adding an arsenic removal agent to the first slurry and carrying out an arsenic removal reaction for 1 hour at a temperature of 50° C. and a negative pressure blower pressure of 0.3 kPa. The resulting product is filtered using a filter press, and the resulting filtrate is used as the second slurry. The gas generated by the arsenic removal reaction is collected using an exhaust gas collection device. The stoichiometric ratio of sulfur atoms in the arsenic removal agent to arsenic ions in the first crude phosphoric acid is 10.

[0205] S14: adding the second slurry to a dechlorination tower, introducing ozone, and performing a dechlorination reaction at a temperature of 50° C. and a negative pressure blower pressure of 0.3 kPa to obtain a third slurry. The gas obtained by the dechlorination reaction is collected by an exhaust gas collection device; wherein the volume ratio of ozone to the second slurry is 500;

[0206] S15: adding activated carbon to the third slurry for decolorization, filtering the resulting product to obtain a fourth slurry, and concentrating the fourth slurry to obtain crude phosphoric acid with a P2O5 mass fraction of 50%; wherein the amount of activated carbon added is 0.5% of the mass of the third slurry;

[0207] S2: extraction;

[0208] S21: If Figure 2As shown, the crude phosphoric acid obtained in step S15 is subjected to multi-stage countercurrent extraction with the first extractant at a temperature of 40° C. to obtain a first extract phase and a first raffinate acid; wherein the volume ratio of tributyl phosphate, sulfonated kerosene and tridecanol in the first extractant is 60:30:10; the number of stages of the multi-stage countercurrent extraction is 4, and the contact time of each stage of extraction is 10 minutes; the volume ratio of the crude phosphoric acid obtained in step S15 to the first extractant is 4; each stage of extraction equipment includes an extraction tank, an emulsion breaking adjustment tank and a phase separation tank, and the extraction Both the extraction tank and the demulsification adjustment tank are equipped with double-layer mesh agitators. The stirring speeds of the first to fourth extraction tanks are 120r / min, 110r / min, 100r / min, and 95r / min, respectively. The stirring speeds of the first to fourth demulsification phase separation tanks are 20r / min, 25r / min, 30r / min, and 35r / min, respectively. Part of the acid phase in each phase separation tank is refluxed to the demulsification adjustment tank, and the volume ratio of the organic phase to the aqueous phase in the demulsification adjustment tank is controlled to be 2:1.

[0209] S22: adding hydrogen peroxide to the first raffinate obtained in step S21, and performing an oxidation reaction at 60° C. for 30 minutes to obtain an oxidized first raffinate; wherein the amount of hydrogen peroxide added is 0.1% of the mass of the first raffinate;

[0210] S23: If Figure 5 As shown, the oxidized first raffinate obtained in step S22 and a second extractant with a saponification rate of 10% are subjected to multi-stage cross-current extraction at a temperature of 60° C. to obtain a second extract phase and a second raffinate acid, and the second raffinate acid is returned to step S11 to be mixed with wet-process phosphoric acid; wherein the volume ratio of the oxidized first raffinate acid to the second extractant is 1:3, the volume ratio of P204, DNNSA, and sulfonated kerosene in the second extractant is 3:3:4, the number of stages of the multi-stage cross-current extraction is 3, and the extraction time of each stage is 30 minutes;

[0211] S24: Figure 4 As shown, the second extract phase is washed with 5% dilute phosphoric acid and then stripped with 40% by mass sulfuric acid to obtain a metal stripping solution and a metal organic phase, and the metal organic phase is returned to step S23 as the second extractant;

[0212] S3: If Figure 3 As shown, the steps of desulfurization treatment are as follows:

[0213] S31: The first extraction phase and the second desulfurization phase of step S21 are added to the decalcification reaction tank for mixing and performing a first reaction. The slurry in the decalcification reaction tank enters the decalcification sedimentation tank for phase separation. The acid clear liquid in the decalcification sedimentation tank is returned to the decalcification reaction tank; the acid residue slurry in the decalcification sedimentation tank enters the decalcification thickening tank for thickening, and then enters the decalcification filter press for filtration. The obtained decalcification filtrate is the first desulfurization phase. Part of the first desulfurization phase is returned to the pretreatment of step S1 to prepare the desulfurization and defluorination agent and / or the dearsenicating agent, and the remaining part of the first desulfurization phase enters the calcium salt preparation tank to prepare the calcium salt solution;

[0214] S32: The organic phase in the decalcification sedimentation tank enters the coarse desulfurization reaction tank and is mixed with the calcium salt solution to perform the first desulfurization reaction. The slurry obtained from the first desulfurization reaction enters the coarse desulfurization sedimentation tank for phase separation. The acid clear liquid in the coarse desulfurization sedimentation tank is returned to the coarse desulfurization reaction tank for the first desulfurization reaction. The acid slag slurry in the coarse desulfurization sedimentation tank enters the coarse desulfurization thickening tank for thickening. The second desulfurization in the coarse desulfurization thickening tank enters the decalcification reaction tank for decalcification. The precipitate in the coarse desulfurization thickening tank enters the decalcification filter press.

[0215] S33: The organic phase and barium salt solution in the coarse desulfurization sedimentation tank enter the fine desulfurization reaction tank for the second desulfurization reaction. The slurry obtained from the second desulfurization reaction enters the fine desulfurization sedimentation tank for phase separation. The acid supernatant in the fine desulfurization sedimentation tank is returned to the fine desulfurization reaction tank for the second desulfurization reaction. The organic phase in the fine desulfurization sedimentation tank enters the washing and stripping section. The acid slag slurry in the fine desulfurization sedimentation tank enters the fine desulfurization thickening tank for thickening. The third desulfurization in the fine desulfurization thickening tank enters the fine desulfurization membrane filtration and concentration equipment for filtration. The resulting filtrate enters the barium salt preparation tank and the calcium salt preparation tank respectively. The resulting precipitate enters the decalcification thickening tank after concentration.

[0216] The temperature during the desulfurization process was 40°C, the reaction time in each tank was 30 minutes, the volume ratio of the organic phase to the aqueous phase in the reaction feed was 10:1, and the volume ratio of the organic phase to the aqueous phase in the decalcification tank, the coarse desulfurization tank, and the fine desulfurization tank was 3:1.

[0217] S4: As Figure 4 As shown, washing: the desulfurized organic phase obtained by desulfurization is added to a washing tower, and multi-stage countercurrent washing is performed with washing acid at a temperature of 50°C to obtain a washed organic phase and a washing residual acid, and the washing residual acid is returned to the desulfurization step to prepare a desulfurizer; wherein, the number of countercurrent washing stages is 2, the washing tower is a vibrating screen plate tower, the vibration frequencies of the first washing tower and the second washing tower are 15Hz and 18Hz respectively, and the volume ratio of the desulfurized organic phase to the washing acid is 15:1;

[0218] S5: Stripping:

[0219] S51: If Figure 4As shown, the washed organic phase obtained by washing is added to a stripping tower, and stripped with desalted water at a temperature of 60° C. to obtain stripping acid and a stripping organic phase, and part of the stripping acid is returned to step S4 as washing acid; wherein the number of stripping stages is 2, the stripping tower is a vibrating sieve plate tower, the vibration frequencies of the first stripping tower and the second stripping tower are 20 Hz and 25 Hz respectively, and the volume ratio of the washed organic phase to the desalted water is 5:1;

[0220] S52: Figure 5 As shown, the stripping organic phase with a specific gravity greater than 0.92 is added to the extraction regeneration tank A, washed with a sodium hydroxide aqueous solution with a mass fraction of 5%, then activated carbon is added for adsorption, and then phase separation is performed. The resulting organic phase is washed with desalted water to obtain a regenerated organic phase, and after adding an appropriate amount of the first extractant to the regenerated organic phase, the process returns to step S21 to perform the first extraction as the first extractant; wherein, the volume ratio of the stripping organic phase to the sodium hydroxide aqueous solution is 2:1, and the amount of activated carbon added is 0.5% of the total mass of the stripping organic phase and the sodium hydroxide aqueous solution;

[0221] S6: Concentration and final treatment:

[0222] S61: The remaining stripping acid obtained in step S51 is concentrated for the first time to obtain phosphoric acid with a mass fraction of 60%, then concentrated for the second time to obtain phosphoric acid with a mass fraction of 75%, and then concentrated for the third time to obtain phosphoric acid with a mass fraction of 85%;

[0223] S62: Adding activated carbon and hydrogen peroxide to 85% by mass phosphoric acid for decolorization, and filtering to obtain finished phosphoric acid.

[0224] The yield of the finished phosphoric acid obtained in this example was 98.5%. The composition of the finished phosphoric acid is shown in Table 2 below.

[0225] Table 2

[0226] Element <![CDATA[P2O5]]> F Cl SO4 Fe Mg Al Ca As Pb content% 61.81 0.0085 0.0005 0.0010 0.0009 0.0010 0.0008 0.0011 0.0002 0.0001

[0227] Example 2

[0228] This embodiment provides a method for purifying wet-process phosphoric acid and a purification system for the method.

[0229] The difference between this embodiment and embodiment 1 is only the pre-processing step, and the remaining steps are the same as those in embodiment 1; the pre-processing steps of this embodiment are as follows:

[0230] S1: pretreatment;

[0231] S11: mixing wet-process phosphoric acid and the second raffinate acid at a mass ratio of 1.5:1 to obtain first crude phosphoric acid;

[0232] S12: adding a desulfurization and defluorination agent to the first crude phosphoric acid to carry out a desulfurization and defluorination reaction at a temperature of 50° C. for 1 hour to obtain a first slurry; wherein the stoichiometric ratio of the phosphate rock in the desulfurization and defluorination agent to the sulfate in the first crude phosphoric acid is 1.2, the total stoichiometric ratio of the sodium carbonate and silicon dioxide in the desulfurization and defluorination agent to the stoichiometric ratio of the fluoride ion in the first crude phosphoric acid is 1.4, and the mass ratio of the sodium carbonate to the silicon dioxide is 3;

[0233] S13: adding an arsenic removal agent to the first slurry and performing a arsenic removal reaction for 1 hour at a temperature of 50° C. and a negative pressure blower pressure of 0.5 kPa. The resulting product is filtered using a filter press, and the resulting filtrate is used as the second slurry. The gas generated by the arsenic removal reaction is collected using an exhaust gas collection device. The stoichiometric ratio of sulfur atoms in the arsenic removal agent to arsenic ions in the first crude phosphoric acid is 15.

[0234] S14: adding the second slurry to a dechlorination tower, introducing ozone, and performing a dechlorination reaction at a temperature of 50° C. and a negative pressure blower pressure of 0.3 kPa to obtain a third slurry. The gas obtained by the dechlorination reaction is collected by an exhaust gas collection device; wherein the volume ratio of ozone to the second slurry is 800;

[0235] S15: adding activated carbon to the third slurry for decolorization, filtering the resulting product to obtain a fourth slurry, and concentrating the fourth slurry to obtain crude phosphoric acid with a P2O5 mass fraction of 48%; wherein the amount of activated carbon added is 0.8% of the mass of the third slurry;

[0236] The yield of the finished phosphoric acid obtained in this example was 97.5%. The composition of the finished phosphoric acid is shown in Table 3 below.

[0237] Table 3

[0238] Element <![CDATA[P2O5]]> F Cl SO4 Fe Mg Al Ca As Pb content% 61.68 0.0065 0.0002 0.0009 0.0008 0.0009 0.0008 0.0009 0.0001 0.0001

[0239] Example 3

[0240] This embodiment provides a method for purifying wet-process phosphoric acid and a purification system for the method.

[0241] The difference between this embodiment and embodiment 2 is only step S21, and the remaining steps are the same as those in embodiment 2; step S21 of this embodiment is as follows:

[0242] S21: If Figure 2As shown, the crude phosphoric acid obtained in step S15 is subjected to multi-stage countercurrent extraction with a first extractant at a temperature of 50° C. to obtain a first extract phase and a first raffinate acid; wherein the volume ratio of tributyl phosphate, sulfonated kerosene, and tridecanol in the first extractant is 70:25:5; the number of stages of the multi-stage countercurrent extraction is 4, and the contact time of each stage of extraction is 10 minutes; the volume ratio of the crude phosphoric acid obtained in step S15 to the first extractant is 1:5; each stage of extraction equipment includes an extraction tank, an emulsion breaking adjustment tank, and a phase separation tank. The extraction tank and demulsification adjustment tank are both equipped with double-layer mesh agitators. The stirring speeds of the first to fourth-stage extraction tanks are 110r / min, 105r / min, 100r / min, and 90r / min, respectively. The stirring speeds of the first to fourth-stage demulsification phase separation tanks are 15r / min, 20r / min, 25r / min, and 30r / min, respectively. Part of the acid phase in each phase separation tank is refluxed to the demulsification adjustment tank, and the volume ratio of the organic phase to the aqueous phase in the demulsification adjustment tank is controlled to be 3:1.

[0243] The yield of the finished phosphoric acid obtained in this example was 98.1%. The composition of the finished phosphoric acid is shown in Table 4 below.

[0244] Table 4

[0245] Element <![CDATA[P2O5]]> F Cl SO4 Fe Mg Al Ca As Pb content% 61.72 0.0075 0.0004 0.0012 0.0010 0.0009 0.0010 0.0010 0.0001 0.0001

[0246] Example 4

[0247] This embodiment provides a method for purifying wet-process phosphoric acid and a purification system for the method.

[0248] The difference between this embodiment and embodiment 2 is only the desulfurization treatment step. The other steps are the same as those in embodiment 2. The desulfurization treatment steps of this embodiment are as follows:

[0249] S3: If Figure 3 As shown, the steps of desulfurization treatment are as follows:

[0250] S31: The first extraction phase and the second desulfurization phase of step S21 are added to the decalcification reaction tank for mixing and performing a first reaction. The slurry in the decalcification reaction tank enters the decalcification sedimentation tank for phase separation. The acid clear liquid in the decalcification sedimentation tank is returned to the decalcification reaction tank; the acid residue slurry in the decalcification sedimentation tank enters the decalcification thickening tank for thickening, and then enters the decalcification filter press for filtration. The obtained decalcification filtrate is the first desulfurization phase. Part of the first desulfurization phase is returned to the pretreatment of step S1 to prepare the desulfurization and defluorination agent and / or the dearsenicating agent, and the remaining part of the first desulfurization phase enters the calcium salt preparation tank to prepare the calcium salt solution;

[0251] S32: The organic phase in the decalcification sedimentation tank enters the coarse desulfurization reaction tank and is mixed with the calcium salt solution to perform the first desulfurization reaction. The slurry obtained from the first desulfurization reaction enters the coarse desulfurization sedimentation tank for phase separation. The acid clear liquid in the coarse desulfurization sedimentation tank is returned to the coarse desulfurization reaction tank for the first desulfurization reaction. The acid slag slurry in the coarse desulfurization sedimentation tank enters the coarse desulfurization thickening tank for thickening. The second desulfurization in the coarse desulfurization thickening tank enters the decalcification reaction tank for decalcification. The precipitate in the coarse desulfurization thickening tank enters the decalcification filter press.

[0252] S33: The organic phase and barium salt solution in the coarse desulfurization sedimentation tank enter the fine desulfurization reaction tank for the second desulfurization reaction. The slurry obtained from the second desulfurization reaction enters the fine desulfurization sedimentation tank for phase separation. The acid supernatant in the fine desulfurization sedimentation tank is returned to the fine desulfurization reaction tank for the second desulfurization reaction. The organic phase in the fine desulfurization sedimentation tank enters the washing and stripping section. The acid slag slurry in the fine desulfurization sedimentation tank enters the fine desulfurization thickening tank for thickening. The third desulfurization in the fine desulfurization thickening tank enters the fine desulfurization membrane filtration and concentration equipment for filtration. The resulting filtrate enters the barium salt preparation tank and the calcium salt preparation tank respectively. The resulting precipitate enters the decalcification thickening tank after concentration.

[0253] The temperature during the desulfurization process was 50°C, the reaction time in each tank was 30 minutes, the reaction feed ratio was 12:1, and the volume ratio of the organic phase to the aqueous phase in the decalcification tank, the rough desulfurization tank, and the fine desulfurization tank was 4:1.

[0254] The yield of the finished phosphoric acid obtained in this example was 97.6%. The composition of the finished phosphoric acid is shown in Table 5 below.

[0255] Table 5

[0256] Element <![CDATA[P2O5]]> F Cl SO4 Fe Mg Al Ca As Pb content% 61.66 0.0072 0.0004 0.0006 0.0010 0.0009 0.0010 0.0010 0.0001 0.0001

[0257] Example 5

[0258] This embodiment provides a method for purifying wet-process phosphoric acid and a purification system for the method.

[0259] The difference between this embodiment and embodiment 4 is only that step S4 and step S51 are different, and the remaining steps are the same as those in embodiment 4; step S4 and step S51 of this embodiment are as follows:

[0260] S4: As Figure 4 As shown, washing: the desulfurized organic phase obtained by desulfurization is added to a washing tower, and multi-stage countercurrent washing is performed with washing acid at a temperature of 60°C to obtain a washed organic phase and a washing residual acid, and the washing residual acid is returned to the desulfurization step to prepare a desulfurizer; wherein, the number of countercurrent washing stages is 2, the washing tower is a vibrating screen plate tower, the vibration frequencies of the first washing tower and the second washing tower are 20Hz and 25Hz respectively, and the volume ratio of the desulfurized organic phase to the washing acid is 12:1;

[0261] S5: Stripping:

[0262] S51: If Figure 4 As shown, the washed organic phase obtained by washing is added to a stripping tower, and stripped with desalted water at a temperature of 60° C. to obtain stripping acid and a stripping organic phase, and part of the stripping acid is returned to step S4 as washing acid; wherein the number of stripping stages is 2, the stripping tower is a vibrating sieve plate tower, the vibration frequencies of the first stripping tower and the second stripping tower are 25 Hz and 30 Hz respectively, and the volume ratio of the washed organic phase to the desalted water is 4:1;

[0263] The yield of the finished phosphoric acid obtained in this example was 96.1%. The composition of the finished phosphoric acid is shown in Table 6 below.

[0264] Table 6

[0265]

[0266]

[0267] Example 6

[0268] This embodiment provides a method for purifying wet-process phosphoric acid and a purification system for the method.

[0269] The difference between this embodiment and embodiment 4 is only that step S22 and step S23 are different, and the remaining steps are the same as those in embodiment 4; step S22 and step S23 of this embodiment are as follows:

[0270] S22: adding hydrogen peroxide to the first raffinate obtained in step S21, and performing an oxidation reaction at 60° C. for 30 minutes to obtain an oxidized first raffinate; wherein the amount of hydrogen peroxide added is 0.3% of the mass of the first raffinate;

[0271] S23: If Figure 5 As shown, the oxidized raffinate obtained in step S22 and a second extractant with a saponification rate of 20% are subjected to multi-stage cross-current extraction at a temperature of 60° C. to obtain a second extract phase and a second raffinate acid, and the second raffinate acid is returned to step S11 to be mixed with wet-process phosphoric acid; wherein the volume ratio of the oxidized first raffinate acid to the second extractant is 1:3, the volume ratio of P204, trialkylphosphine oxide, and sulfonated kerosene in the second extractant is 2:5:3, the number of stages of the multi-stage cross-current extraction is 3, and the extraction time of each stage is 50 minutes;

[0272] The yield of the finished phosphoric acid obtained in this example was 96.2%. The composition of the finished phosphoric acid is shown in Table 7 below.

[0273] Table 7

[0274] Element <![CDATA[P2O5]]> F Cl SO4 Fe Mg Al Ca As Pb content% 61.72 0.0061 0.0002 0.0006 0.0002 0.0002 0.0003 0.0004 0.0001 0.0001

[0275] Finally, it should be noted that the above embodiments are used to illustrate the technical solutions of the present invention rather than to limit the scope of protection of the present invention. Although the present invention has been described in detail with reference to the preferred embodiments, those skilled in the art should understand that the technical solutions of the present invention may be modified or replaced by equivalents without departing from the essence and scope of the technical solutions of the present invention.

Claims

1. A method for purifying wet-process phosphoric acid, characterized in that: The following steps are involved: The wet-process phosphoric acid is sequentially subjected to pretreatment, extraction, desulfurization, washing, stripping, and final treatment to obtain a phosphoric acid product; The extraction step comprises: performing a first extraction on the crude phosphoric acid obtained by pretreatment with a first extractant to obtain a first extraction phase and a first raffinate acid; The first raffinate is subjected to a second extraction with a second extractant to obtain a second raffinate and a second extract phase, and the second raffinate is returned to the pretreatment step to be mixed with wet-process phosphoric acid; the second extract phase is subjected to impurity removal and then returned to the second extraction step as a second extractant; The first extractant is a mixture of tributyl phosphate, sulfonated kerosene and tridecanol.

2. The method for purifying wet-process phosphoric acid according to claim 1, wherein The volume ratio of tributyl phosphate, sulfonated kerosene and tridecanol is tributyl phosphate: sulfonated kerosene: tridecanol = 60-75: 20-30: 5-10; And / or, the mass fraction of P2O5 in the crude phosphoric acid obtained by the pretreatment is 48%-50%.

3. The method for purifying wet-process phosphoric acid according to claim 1, wherein The following steps are involved: Pretreatment: Pretreating wet-process phosphoric acid to obtain crude phosphoric acid; Extraction: performing a first extraction on the crude phosphoric acid with a first extractant to obtain a first extract phase and a first raffinate; The first raffinate is subjected to a second extraction to obtain a second raffinate and a second extract phase, and the second raffinate is returned to the pretreatment step to be mixed with wet-process phosphoric acid; the second extract phase is subjected to impurity removal and then returned to the second extraction step to serve as a second extractant; wherein the first extractant is a mixture of tributyl phosphate, sulfonated kerosene, and tridecanol; Desulfurization treatment: desulfurizing the first extraction phase obtained by extraction with a desulfurizing agent to obtain a desulfurized organic phase and desulfurized sulfuric acid; Washing: washing the desulfurized organic phase with washing acid to obtain a washed organic phase and a washing residual acid, and returning the washing residual acid to the desulfurization treatment step to prepare a desulfurizer; Stripping: stripping the washed organic phase with a stripping agent to obtain stripping acid and a stripping organic phase, regenerating the stripping organic phase to obtain a regenerated organic phase, and returning the regenerated organic phase to the extraction step as the first extractant; Final treatment: The stripping acid obtained by stripping is returned to the washing step as washing acid, and the remaining part is concentrated and then subjected to final treatment to obtain the phosphoric acid product.

4. The method for purifying wet-process phosphoric acid according to claim 3, wherein: The pretreatment comprises the following steps: sequentially performing desulfurization, defluorination, dearsenicization, dechlorination and decolorization on the wet-process phosphoric acid, and then filtering and collecting the filtrate.

5. The method for purifying wet-process phosphoric acid according to claim 4, wherein: The steps of desulfurization, defluorination, dearsenicization, dechlorination and decolorization are: mixing phosphate rock powder, sodium carbonate, silicon dioxide and the desulfurized water to obtain a desulfurization and defluorination agent, and mixing a sulfur-containing compound with the desulfurized water to obtain a dearsenicizing agent, wherein the sulfur-containing compound is sodium sulfide or phosphorus pentasulfide; A desulfurization and defluorination agent is sequentially added to wet-process phosphoric acid for desulfurization and defluorination, a dearsenicizing agent is added for dearsenicization, ozone is added for dechlorination, and activated carbon is added for decolorization; wherein, the stoichiometric ratio of phosphate rock to sulfate in the wet-process phosphoric acid is 1.05-1.2, the stoichiometric ratio of the total stoichiometric amount of sodium carbonate and silicon dioxide to the fluoride ion in the wet-process phosphoric acid is 1.2-1.4, the stoichiometric ratio of sulfur atoms in the dearsenicizing agent to arsenic ions in the wet-process phosphoric acid is 10-20, and the volume ratio of ozone to the volume of wet-process phosphoric acid is 500-1000.

6. The method for purifying wet-process phosphoric acid according to claim 4, wherein: The temperature of the desulfurization, defluorination, dearsenicization, dechlorination and decolorization is 50-60° C., and / or the negative pressure of the dearsenicization, dechlorination and decolorization is 0.3-1.5 kPa.

7. The method for purifying wet-process phosphoric acid according to claim 3, wherein: In the extraction step, at least one of the following is satisfied: (1) The first extraction is a multi-stage countercurrent extraction with 4-6 extraction stages; (2) the volume ratio of the first extractant to crude phosphoric acid is 4-6; (3) The temperature of the first extraction is 40-50°C; (4) The second extraction step comprises: adding the first raffinate acid to hydrogen peroxide and reacting at a temperature of 60-70° C. to obtain an oxidized first raffinate acid; then adding a second extractant to perform multi-stage cross-current extraction to obtain a second raffinate acid and a second extract phase; (5) The impurity removal step of the second extraction phase is as follows: the second extraction phase is washed with 1%-10% by mass phosphoric acid, and then stripped with 30%-40% by mass sulfuric acid to obtain a metal stripping solution and a metal organic phase.

8. The method for purifying wet-process phosphoric acid according to claim 7, wherein: The mass of the hydrogen peroxide is 0.1%-0.5% of the mass of the first raffinate acid.

9. The method for purifying wet-process phosphoric acid according to claim 7, wherein: The preparation method of the second extractant comprises: mixing the extractant and sulfonated kerosene, adding sodium hydroxide solution to carry out saponification reaction, and separating the phases to obtain an organic phase, which is the second extractant; wherein the extractant is at least one of 2-ethylhexyl phosphate, di(2-ethylhexyl) phosphate, dinonylnaphthalenesulfonic acid, cyclohexane acid, trialkylphosphine oxide, and acetylacetone, and the saponification rate of the second extractant is 10%-30%.

10. The method for purifying wet-process phosphoric acid according to claim 9, wherein: The volume ratio of the extractant to sulfonated kerosene is 1-2.

5.

11. The method for purifying wet-process phosphoric acid according to claim 7, wherein: The number of stages of the multi-stage cross-current extraction is 3-5, the volume ratio of the second extractant to the oxidized first raffinate acid is 3-5, and the temperature of the multi-stage cross-current extraction is 40-60°C.

12. The method for purifying wet-process phosphoric acid according to claim 3, wherein: The steps of the desulfurization treatment are: Solution preparation: Calcium salt and residual washing acid are prepared into calcium salt solution, and barium salt and residual washing acid are prepared into barium salt solution; Desulfurization: The first extraction phase and the second desulfurization phase are mixed to perform a first reaction, and phase separation is performed to obtain a first organic phase and a first desulfurization phase, and the first desulfurization phase is returned to the pretreatment step; mixing the first organic phase and the calcium salt solution and performing a second reaction to obtain a second organic phase and a second desulfurization; The second organic phase and the barium salt solution are mixed and then subjected to a third reaction to obtain a desulfurized organic phase and a third desulfurized phase; The third desulfurization is returned to the pretreatment step.

13. The method for purifying wet-process phosphoric acid according to claim 12, wherein: The volume ratio of the first extraction phase to the second desulfurization phase is 10-15, the volume ratio of the first organic phase to the calcium salt solution is 10-15, the volume ratio of the second organic phase to the barium salt solution is 10-15, and the temperatures of the first reaction, the second reaction and the third reaction are each independently 40-50°C.

14. The method for purifying wet-process phosphoric acid according to claim 12, wherein: The calcium salt is calcium carbonate or calcium hydroxide, and the barium salt is barium carbonate.

15. The method for purifying wet-process phosphoric acid according to claim 3, wherein: In the washing step, the washing is countercurrent washing, and the number of washing stages is 2-3; and / or, the volume ratio of the desulfurized organic phase to the washing acid is 12-18; And / or, the washing temperature is 50-60°C.

16. The method for purifying wet-process phosphoric acid according to claim 3, wherein: In the stripping step, the stripping agent is desalted water; and / or, the volume ratio of the washing organic phase to the stripping agent is 4-6; And / or, the number of stages of the stripping is 2-4; And / or, the stripping temperature is 60-70°C; And / or, the regeneration step of the stripping organic phase is: washing the stripping organic phase with a sodium hydroxide aqueous solution, adding activated carbon to adsorb impurities, and then washing with desalted water to obtain a regenerated organic phase.

17. The method for purifying wet-process phosphoric acid according to claim 16, wherein: The mass fraction of the sodium hydroxide aqueous solution is 4%-6%, and the mass of the activated carbon is 0.5%-1% of the mass of the stripping organic phase.

18. A purification system for the wet-process phosphoric acid purification method according to any one of claims 1 to 17, characterized in that: It includes a pre-treatment system, a first extraction system, a desulfurization treatment system, a washing and stripping system, a second extraction system and a regeneration system.

19. The system of claim 18, wherein: The first extraction system includes an extraction tank, an emulsion breaking and regulating tank, and a phase separation tank; And / or, the desulfurization treatment system includes a solution preparation system, a phase separation system, a first desulfurization system, and a second desulfurization system; And / or, the washing and stripping system includes a washing tower and a stripping tower; And / or, the second extraction system includes an extraction tank, a phase separation tank and a stripping tank; And / or, the regeneration system includes a regeneration tank, a phase separation tank and a water washing tower.

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