Purification method and purification system for wet-process phosphoric acid

Abstract

The invention provides a wet-process phosphoric acid purification method and purification system, and belongs to the technical field of wet-process phosphoric acid purification. Raffinate acid generated through extraction is subjected to secondary extraction and then returns to be mixed with wet-process phosphoric acid, washing residual acid is returned to a desulfurization treatment step and is recycled to form an acid closed loop, and the wet-process phosphoric acid is obtained. The whole wet-process phosphoric acid purification process does not generate acid by-products and is environment-friendly, the yield of the purified phosphoric acid product is greater than 95%, and the treatment trouble of the acid by-products does not need to be worried about.

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 of thermal phosphoric acid production, severe environmental pollution, and high requirements for ore, the world's energy is increasingly tight, causing the price of thermal phosphoric acid to continue to rise. Therefore, wet phosphoric acid has begun to gradually replace thermal phosphoric acid. The impurity content in wet phosphoric acid is relatively high, mainly including anions SO4 2- 、F - As well as cations such as Fe, A1, Mg, Ca, Si, etc. The end product of industrial phosphoric acid production generally requires phosphoric acid of relatively good quality. Therefore, the wet-process phosphoric acid must be purified to meet production quality standards.

[0003] At present, the purification methods of wet phosphoric acid mainly include crystallization, ion exchange, chemical precipitation, solvent extraction, etc. The crystallization method needs to be concentrated to a higher concentration, which has problems in the selection of anti-corrosion materials. In addition, due to crystal entrainment, it is difficult to obtain high-purity acid, and the yield of crystallized products is low. Although the process equipment of the ion exchange method is relatively simple, the ion exchange resin is expensive, the phosphoric acid concentration from the resin is low, the evaporation volume is too large, the energy consumption is too high, and it is not suitable for large-scale production, and the one-time investment is large. The chemical precipitation method can remove some cations or anions, but the purification depth is not high, and another ion is introduced. The solvent extraction method has become the most commonly used method due to its good purification effect, mild conditions, high output and stable operation.

[0004] The core technology of solvent extraction purification of wet-process phosphoric acid is to extract and separate wet-process phosphoric acid with high impurity content through organic solvent to obtain industrial-grade phosphoric acid. The P2O5 extraction rate of wet-process phosphoric acid solvent extraction purification is generally 50% to 70%, and the purification yield is low. In addition, the by-product raffinate acid is phosphoric acid containing P2O5 and high iron, aluminum, and magnesium impurities. The impurity ion content in the raffinate acid is doubled compared with the original phosphoric acid, which greatly limits the application scope of the raffinate acid. At present, the utilization of domestic raffinate acid is mainly concentrated in the production of monoammonium phosphate fertilizer. In recent years, due to the overcapacity of high-concentration phosphate fertilizers and large fluctuations in the market price of phosphate fertilizers, the rational utilization of by-product raffinate acid has become a technical problem that enterprises need to solve urgently. If the phosphorus in the by-product raffinate acid can be recovered or no raffinate acid is produced, the economic benefits will be greatly improved. Summary of the invention

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

[0006] To achieve the above object, 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 treatment, washing, stripping, and final treatment to obtain a phosphoric acid product;

[0008] The extraction steps are as follows: subjecting the crude phosphoric acid obtained by pretreatment to a first extraction with a first extractant to obtain a first extraction phase and a first raffinate acid; subjecting the first raffinate acid to a second extraction 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 subjecting the second extraction phase to impurities removal and then returning it to the second extraction step as a second extractant;

[0009] The first extracting agent 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: subjecting the crude phosphoric acid to a first extraction with a first extractant to obtain a first extraction phase and a first raffinate acid; subjecting the first raffinate acid to a second extraction 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; after impurities are removed from the second extraction phase, returning the second extraction phase to the second extraction step as a second extractant; wherein the first extractant 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 a 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] The phosphate rock powder, sodium carbonate, silicon dioxide and the desulfurized water are mixed to obtain a desulfurization and defluorination agent, and the sulfur-containing compound and the desulfurized water are mixed 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 dearsenicization, ozone for dechlorination, and activated carbon for decolorization; wherein the stoichiometric ratio of phosphate rock powder to sulfate in the wet-process phosphoric acid is 1.05-1.2, and the stoichiometric ratio of the total stoichiometric ratio of sodium carbonate and silicon dioxide to fluoride ions 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;

[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, in the extraction step, at least one of the following is satisfied:

[0025] (1) The first extraction is a multi-stage countercurrent extraction, and the number of extraction stages is 4-6;

[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, 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;

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

[0030] Preferably, the preparation method of the second extractant is: 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; wherein the extractant is at least one of 2-ethylhexyl phosphate 2-ethylhexyl ester, 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 acid 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: washing the second extraction phase with 1%-10% by mass phosphoric acid, and then stripping with 30%-40% by mass sulfuric acid to obtain a metal stripping solution and a metal organic phase.

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

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

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

[0036] The first organic phase and the calcium salt solution are mixed and then subjected to 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 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;

[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 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 comprises a washing tower and a stripping tower;

[0053] And / or, the second extraction system comprises 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 acid produced by extraction, returns it to mix with 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 by-products, which is environmentally friendly. The yield of the purified phosphoric acid product is greater than 95%, and there is no need to worry about the treatment of acid by-products.

[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 and non-toxic, and has a low water solubility. It is a good phosphoric acid extractant. The compounding 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 compound of tributyl phosphate and sulfonated kerosene is at low temperature or when the phosphoric acid extraction rate is high, the kerosene is easily precipitated 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, which can change the polarity between the first extractants, increase the compatibility between organic substances, prevent the generation 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 of the desulfurization process 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 removing sulfuric acid) 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 taken away by the acid phase, which improves the purification effect of phosphoric acid. In addition, the extract phase is desulfurized before washing, which achieves the effect of pre-washing and improves 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 separates the phases, which has the following advantages. First, the calcium ion concentration in the second desulfurization phase is high, which can achieve the effect of pre-desulfurization 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 to avoid the formation of precipitation during the pretreatment process of the first desulfurization phase returning to step S1, which leads to 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 The process flow chart of the purification method of wet-process phosphoric acid disclosed in the present invention is as follows;

[0061] Figure 2 A process flow chart of a first extraction system in a purification system for the purification method of wet-process phosphoric acid disclosed in the present invention;

[0062] Figure 3 A process flow chart of a desulfurization system in a purification system used in the purification method of wet-process phosphoric acid disclosed in the present invention;

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

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

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

[0066] In order to facilitate understanding of the present disclosure, the present disclosure will be described more fully below. However, the present disclosure can be implemented in many different forms and is not limited to the embodiments described herein. On the contrary, the purpose of providing these embodiments is to make the understanding of the disclosure of the present disclosure more thorough and comprehensive.

[0067] As used herein:

[0068] "Prepared from" is synonymous with "comprising." As used herein, the terms "comprising," "including," "having," "containing," or any other variation 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-ended so that it does not include 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 is represented by range, preferred range or a series of upper preferred value and lower preferred value, 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, regardless of whether the range is disclosed separately. For example, when disclosing range "1-5", described range should be interpreted as including range "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, the range is intended to include its end value and all integers and fractions within the 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 for expressing the mass ratio of multiple components. 1 part can represent any unit 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, it means the ratio of the mass of component A to the mass of component B is a:b. Alternatively, it means that the mass of component A is aK and the mass of component B is bK (K is an arbitrary number, indicating a multiple factor). It should not be misunderstood that, unlike the mass parts, 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 phosphoric acid and the extraction rate of wet 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: The 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, it can be 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, 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 powder 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 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 1.2-1.4, for example, it can be 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 dearsenicating agent to the first slurry to carry out a dearsenicating reaction, and filtering the obtained product with a filter press to obtain a second slurry.

[0086] Specifically, the arsenic removal agent includes a sulfur-containing compound and a desulfurization agent mixed 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 into 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 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 dearsenicification 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 to avoid the toxic gas from overflowing, and transfer 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 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 the chemical impurity removal reagent used in the pretreatment process, which can not only ensure continuous and stable feeding, but also the chemical impurity removal reagent and wet-process phosphoric acid undergo 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 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, and 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, and 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, and the volume ratio of tributyl phosphate, sulfonated kerosene and tridecanol is tributyl phosphate: sulfonated kerosene: tridecanol = 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 the first extractant. Tributyl phosphate has a good removal effect on phosphate cations, is safe and non-toxic, and has a low water solubility. It is a good phosphoric acid extractant. The compound use 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 compound of tributyl phosphate and sulfonated kerosene is at a low temperature or when the phosphoric acid extraction rate is high, the kerosene is easily precipitated 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, which can change the polarity between the first extractants, increase the compatibility between organic substances, prevent the generation 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%, 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, 70°C.

[0108] Specifically, the oxidation reaction time is 30-60 min, for example, it can be 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 extraction 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, it can be but not limited to 5:1, 4.5:1, 4:1, 3.5:1, 3:1;

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

[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, 60°C.

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

[0114] Specifically, the preparation method of the second extractant is: after mixing the extractant and sulfonated kerosene, adding sodium hydroxide solution to carry out saponification reaction, separating the phases, 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, it can be 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 calculation formula of the second extractant is 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, it can be but not limited to 1%, 3%, 5%, 7%, 10%.

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

[0126] Specifically, the stripping time is 40-60 min, for example, it can be 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, it can be but not limited to 2, 3, 4.

[0128] In the present invention, 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 the loss of phosphorus 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 to improve the yield of phosphoric acid. In addition, after the second extraction phase generated 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, defluorination agent and / or a dearsenicizing agent;

[0131] S32: mixing the first organic phase and the calcium salt solution and performing a second reaction, and performing phase separation 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 residual washing acid, and the barium salt solution includes barium salt and residual washing acid, wherein the calcium salt is carbonate and / or calcium hydroxide, and the barium salt is barium carbonate;

[0134] Specifically, the temperatures of the first reaction, the second reaction and the third reaction are each 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 of 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 is mixed with the first extraction phase. On the one hand, the second desulfurization contains a relatively high concentration of calcium ions, which can pre-desulfurize the first extraction phase, make full use of calcium ions, improve the utilization rate of calcium salts, and reduce the cost of calcium salts; on the other hand, it can reduce the calcium ions in the first desulfurization, and avoid the formation of precipitation during the pretreatment process of the first desulfurization returning to step S1, resulting in blockage of the pretreatment system structure; on the other hand, if the second desulfurization is directly returned to the pretreatment process of step S1, the calcium content of the finished phosphoric acid will be too high, making the quality of the finished phosphoric acid unqualified;

[0138] Step 12, step S32 and step S33 in the present disclosure constitute three precipitation desulfurization in the wet phosphoric acid purification process. The desulfurization rate of step 12 is 60-80%, the desulfurization rate of step S32 can reach more than 90%, and the desulfurization rate of step S33 can reach more than 99%; Steps S32 and S33 use calcium salt solution and barium salt solution for desulfurization in turn. On the one hand, the calcium salt solution and the barium salt solution are both liquid-liquid contact reactions with the organic phase, the desulfurization reaction is fast, the desulfurization efficiency is high, and the introduced impurities are all taken away by the acid phase, which improves the purity of the finished phosphoric acid. On the second hand, the calcium salt solution and the barium salt solution are used for desulfurization, which can achieve the effect of pre-washing the extraction 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: washing the desulfurized organic phase with a washing acid in countercurrent 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, it can be 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, it can be but not limited to 50°C, 52°C, 54°C, 56°C, 58°C, 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 obtained by washing 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 the stripping is 2-4;

[0147] Specifically, the volume ratio of the washing 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, 70°C.

[0149] In the present disclosure, the stripping acid is phosphoric acid with a certain concentration and low impurity content. The stripping acid washing can not only achieve the effect of water washing, but also bring 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: washing the stripping organic phase with a sodium hydroxide aqueous solution, then adding activated carbon for adsorption, and then performing phase separation, washing the obtained organic phase with desalted water to obtain a regenerated organic phase, and adding an appropriate amount of the first extractant to the regenerated organic phase, and returning to step S21 as the first extractant for the first extraction;

[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] The first extractant will produce decomposition products and interphase contaminants after long-term use. 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 to obtain better economic value.

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

[0161] like Figure 2 As shown, taking the extraction stage 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 into the primary extraction tank to start the primary extraction. After the primary extraction, the mixed slurry enters the primary demulsification adjustment tank for primary demulsification. After the primary demulsification, 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 part enters the secondary extraction tank for secondary extraction; after the secondary extraction, the mixed slurry enters the secondary demulsification adjustment tank for secondary demulsification, after the secondary demulsification, the mixed slurry enters the secondary phase separation tank for secondary phase separation, 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 part 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 flows back 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 part in the fourth-stage phase separation tank flows back 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 large, the mass transfer is slow, and the phase separation is slow. The extraction tank and the 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] In each extraction process, part of the acid phase in the phase separation tank flows back to the demulsification adjustment tank, and the volume ratio of the organic phase and the aqueous phase in the demulsification adjustment tank can be controlled to be 2-3, which is beneficial to demulsification and phase separation. For example, when the phase ratio in the demulsification adjustment tank is too large, phase separation is difficult, and part of the acid phase that is returned increases the contact area of ​​the acid phase, which allows for better phase separation and reduces the acid entrained in the extraction phase, which 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, the slurry obtained by the decalcification reaction enters the decalcification sedimentation tank for phase separation, and the acid clear liquid in the decalcification sedimentation tank is returned to the decalcification reaction tank; the acid slag slurry in the decalcification sedimentation tank enters the decalcification thickening tank for thickening, and then enters the decalcification filter press for filtration, and 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, and 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, the third desulfurization in the fine desulfurization thickening tank enters the fine desulfurization membrane filtration and concentration equipment for filtration, and 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 to ensure 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 is reduced during the preparation of calcium salt solution and barium salt solution, 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 enters the calcium salt preparation tank and the barium salt preparation tank of the desulfurization treatment process respectively.

[0177] The ester phase obtained from the primary washing enters the secondary washing tower and is subjected to secondary washing with stripping acid, and the acid phase obtained from the secondary washing is returned 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. After the acid phase obtained from the primary stripping is collected in the stripping acid buffer tank, 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, and 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 product obtained by the reaction 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, the organic phase obtained by phase separation enters the washing tower for washing with desalted water, and the regenerated extractant and the washing alkali solution are separated; 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 stage as 3 and the stripping stage 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 acid is added to a primary extraction tower and subjected to primary extraction with a second extractant. The slurry obtained by the primary extraction enters a primary phase separation tank for primary phase separation. The organic phase in the primary phase separation tank is collected by 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, the slurry obtained from the secondary extraction enters the secondary phase separation tank for secondary phase separation, and the organic phase in the secondary phase separation tank is collected by the metal ion loading 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 by the metal ion loading extraction phase buffer tank, and the acid phase in the tertiary phase separation tank is collected by the 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 use sodium hydroxide solution to carry out saponification reaction, separate the phases, and the obtained organic phase is the second extractant;

[0196] The composition of wet-process phosphoric acid is 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 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 phosphate rock powder in the desulfurization and defluorination agent to 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 a dearsenicating agent to the first slurry to carry out a dearsenicating reaction for 1 hour at a temperature of 50° C. and a pressure of 0.3 kPa from a negative pressure blower, filtering the obtained product with a filter press, and obtaining a filtrate as the second slurry; and collecting the gas obtained from the dearsenicating reaction with a tail gas collection device; wherein the stoichiometric ratio of sulfur atoms in the dearsenicating 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, and collecting the gas obtained by the dechlorination reaction using a tail 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 obtained 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: Figure 2As shown, the crude phosphoric acid obtained in step S15 and the first extractant are subjected to multi-stage countercurrent extraction at a temperature of 40° C. to obtain a first extraction 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 min; the volume ratio of the crude phosphoric acid obtained in step S15 to the first extractant is 4; each stage of the extraction equipment includes an extraction tank, a demulsification adjustment tank and a phase separation tank, and the extraction Both the extraction tank and the demulsification adjustment tank are 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 and the aqueous phase in the demulsification adjustment tank is controlled to be 2:1.

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

[0210] S23: Figure 5 As shown, the first oxidized raffinate obtained in step S22 and the 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 extraction 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 first oxidized 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 extraction 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 extraction agent;

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

[0213] S31: adding the first extraction phase and the second desulfurization of step S21 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, and the acid clear liquid in the decalcification sedimentation tank is returned to the decalcification reaction tank; the acid slag slurry in the decalcification sedimentation tank enters the decalcification thickening tank for thickening, and then enters the decalcification filter press for filtration, and the obtained decalcification filtrate is the first desulfurization, part of the first desulfurization 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 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 reaction. 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, and 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, and 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;

[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 and the aqueous phase in the reaction feed was 10:1, and the volume ratio of the organic phase and the aqueous phase in the decalcification reaction tank, the rough desulfurization reaction tank, and the fine desulfurization reaction tank was 3:1.

[0217] S4: Figure 4 As shown, washing: adding the desulfurized organic phase obtained by desulfurization into a washing tower, performing multi-stage countercurrent washing with washing acid at a temperature of 50°C to obtain a washing organic phase and washing residual acid, and returning the washing residual acid 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: 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 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%, and then activated carbon is added for adsorption, and then phase separation is performed, and the obtained 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, it returns to step S21 as the first extractant for the first extraction; 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 residual 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 phosphoric acid with a mass fraction of 85% to decolorize the phosphoric acid, and filtering to obtain finished phosphoric acid.

[0224] The yield of the finished phosphoric acid obtained in this example is 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 for purifying wet-process phosphoric acid.

[0229] The difference between this embodiment and embodiment 1 is only the pretreatment step, and the other steps are the same as those of embodiment 1; the pretreatment 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 phosphate rock powder in the desulfurization and defluorination agent to sulfate in the first crude phosphoric acid is 1.2, 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 1.4, and the mass ratio of sodium carbonate to silicon dioxide is 3;

[0233] S13: adding a dearsenicating agent to the first slurry to carry out a dearsenicating reaction for 1 hour at a temperature of 50° C. and a pressure of 0.5 kPa from a negative pressure blower, filtering the obtained product with a filter press, and obtaining a filtrate as the second slurry; and collecting the gas obtained from the dearsenicating reaction with a tail gas collection device; wherein the stoichiometric ratio of sulfur atoms in the dearsenicating agent to arsenic ions in the first crude phosphoric acid is 15;

[0234] S14: adding the second slurry into 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, and collecting the gas obtained by the dechlorination reaction using a tail 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 obtained 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 is 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 for purifying wet-process phosphoric acid.

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

[0242] S21: Figure 2As shown, the crude phosphoric acid obtained in step S15 and the first extractant are subjected to multi-stage countercurrent extraction at a temperature of 50° C. to obtain a first extraction 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 min; the volume ratio of the crude phosphoric acid obtained in step S15 to the first extractant is 1:5; each stage of the extraction equipment includes an extraction tank, a demulsification adjustment tank and a phase separation tank, The extraction tank and demulsification adjustment tank are both double-layer mesh agitators. The stirring speeds of the first to fourth extraction tanks are 110r / min, 105r / min, 100r / min, and 90r / min, respectively. The stirring speeds of the first to fourth 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 and 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 is 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 for purifying wet-process phosphoric acid.

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

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

[0250] S31: adding the first extraction phase and the second desulfurization of step S21 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, and the acid clear liquid in the decalcification sedimentation tank is returned to the decalcification reaction tank; the acid slag slurry in the decalcification sedimentation tank enters the decalcification thickening tank for thickening, and then enters the decalcification filter press for filtration, and the obtained decalcification filtrate is the first desulfurization, part of the first desulfurization 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 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 reaction. 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, and 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, and 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;

[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 reaction tank, the rough desulfurization reaction tank, and the fine desulfurization reaction tank was 4:1.

[0254] The yield of the finished phosphoric acid obtained in this example is 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 for purifying wet-process phosphoric acid.

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

[0260] S4: Figure 4 As shown, washing: adding the desulfurized organic phase obtained by desulfurization into a washing tower, performing multi-stage countercurrent washing with washing acid at a temperature of 60°C to obtain a washing organic phase and washing residual acid, and returning the washing residual acid 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: 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 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 is 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 for purifying wet-process phosphoric acid.

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

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

[0271] S23: Figure 5 As shown, the oxidized raffinate obtained in step S22 and the 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 extraction 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 min;

[0272] The yield of the finished phosphoric acid obtained in this example is 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 solution 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 solution of the present invention can be modified or replaced by equivalents without departing from the essence and scope of the technical solution 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 treatment, washing, stripping, and final treatment to obtain a phosphoric acid product; The extraction step comprises: subjecting the crude phosphoric acid obtained by pretreatment 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 with a second extractant to obtain a second raffinate acid and a second extract phase, and the second raffinate acid 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 extracting agent is a mixture of tributyl phosphate, sulfonated kerosene and tridecanol.

2. The method for purifying wet-process phosphoric acid according to claim 1, characterized in that: 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, characterized in that: The following steps are involved: Pretreatment: pretreating wet-process phosphoric acid to obtain crude phosphoric acid; Extraction: subjecting the crude phosphoric acid 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 as a second extraction agent; wherein the first extraction agent 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 a phosphoric acid product.

4. The method for purifying wet-process phosphoric acid according to claim 3, characterized in that: 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; Preferably, the steps of desulfurization, defluorination, dearsenicization, dechlorination and decolorization are: The phosphate rock powder, sodium carbonate, silicon dioxide and the desulfurized water are mixed to obtain a desulfurization and defluorination agent, and the sulfur-containing compound and the desulfurized water are mixed to obtain a dearsenicizing agent, wherein the sulfur-containing compound is sodium sulfide or phosphorus pentasulfide; The wet-process phosphoric acid is sequentially added with a desulfurization and defluorination agent for desulfurization and defluorination, a dearsenicizing agent for dearsenicization, ozone for dechlorination, and activated carbon for decolorization; wherein the stoichiometric ratio of phosphate rock powder to sulfate in the wet-process phosphoric acid is 1.05-1.2, the stoichiometric ratio of the total stoichiometric ratio of sodium carbonate and silicon dioxide to fluoride ions 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; 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.

5. The method for purifying wet-process phosphoric acid according to claim 3, characterized in that: In the extraction step, at least one of the following is satisfied: (1) The first extraction is a multi-stage countercurrent extraction, and the number of extraction stages is 4-6; (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, 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; Preferably, the mass of the hydrogen peroxide is 0.1%-0.5% of the mass of the first raffinate acid; Preferably, the preparation method of the second extractant is: 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; wherein the extractant is at least one of 2-ethylhexyl phosphate 2-ethylhexyl ester, di(2-ethylhexyl) phosphate, dinonylnaphthalenesulfonic acid, cyclohexane acid, trialkylphosphine oxide, and acetylacetone, and the saponification rate of the second extractant is 10%-30%; 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 acid after oxidation is 3-5, and the temperature of the multi-stage cross-current extraction is 40-60°C; (5) The impurity removal step of the second extraction phase is: washing the second extraction phase with 1%-10% by mass phosphoric acid, and then stripping with 30%-40% by mass sulfuric acid to obtain a metal stripping solution and a metal organic phase.

6. The method for purifying wet-process phosphoric acid according to claim 3, characterized in that: The steps of the desulfurization treatment are: Solution preparation: calcium salt and washing residual acid are prepared into calcium salt solution, and barium salt and washing residual 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; The first organic phase and the calcium salt solution are mixed and then subjected to 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 desulfurization step; wherein the third desulfurization step is returned to the pretreatment step; 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; More preferably, the calcium salt is calcium carbonate or calcium hydroxide, and the barium salt is barium carbonate.

7. The method for purifying wet-process phosphoric acid according to claim 3, characterized in that: 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.

8. The method for purifying wet-process phosphoric acid according to claim 3, characterized in that: 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; 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.

9. A purification system for the wet-process phosphoric acid purification method according to any one of claims 1 to 8, 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.

10. The system according to claim 9, characterized in that The first extraction system comprises 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 comprises a washing tower and a stripping tower; And / or, the second extraction system comprises 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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