Method for co-production of phosphoric acid and iron phosphate by membrane separation wet process

Abstract

According to the method for co-production of the wet-process phosphoric acid and the iron phosphate through membrane separation, an acid-resistant nanofiltration membrane is adopted to perform first-stage membrane separation on the wet-process phosphoric acid, and a first-stage clear phase is obtained; performing second-stage membrane separation on the first-stage clear phase through an acid-resistant nanofiltration membrane to obtain a second-stage concentrated phase and purified acid; mixing the purified acid, the filter cake and water to perform crystal transformation reaction, filtering and drying slurry, and calcining a solid phase to obtain the anhydrous iron phosphate. According to the method, the wet-process phosphoric acid is subjected to membrane separation by using the two-stage acid-resistant nanofiltration membrane to obtain the purified acid and the concentrated acid, and the purified acid and the concentrated acid are both used as raw materials for producing the battery-grade iron phosphate, so that raw material waste is avoided, and atomic economy is realized. The membrane separation is simple in operation, low in cost and low in energy consumption, phosphorus combined with hydrogen ions is separated into purified acid through membrane separation, phosphorus combined with impurity ions is separated into concentrated acid, and phosphorus combined with the hydrogen ions contained in the concentrated acid is precipitated into a filter cake through auxiliary precipitation and ammonia neutralization to be used as a fertilizer.

Description

Technical Field

[0001] The present invention belongs to the technical field of phosphoric acid purification and co-production of iron phosphate, and relates to a method for co-producing iron phosphate by membrane separation of wet-process phosphoric acid. Background Art

[0002] Under the current background, the reconstruction of China's energy system is imperative. With the support of the state for the new energy industry, the new energy vehicle and photovoltaic industries have entered a stage of rapid development, and the demand has been continuously increasing, thus driving the rapid growth of the demand for various sub-products in the upstream links of the industrial chain.

[0003] According to the growth trend of the cathode materials for new energy batteries, the future market demand for lithium iron phosphate materials is likely to maintain a high-speed growth. As one of the main precursors for the production of lithium iron phosphate, the demand for iron phosphate will also maintain a rapid growth momentum. There are many domestic iron phosphate production technologies, and among them, the ammonium method using industrial ammonium phosphate as the raw material is one of the current mainstream processes, accounting for about 60%; followed by the sodium method, accounting for about 25%, and then the iron method, accounting for about 10%. In the raw and auxiliary material costs of iron phosphate, the proportion of phosphorus source raw materials (ammonium phosphate or 85% phosphoric acid) is the largest. In the ammonium method process, 0.78t of ammonium phosphate and 0.13t of 85% phosphoric acid are consumed to produce one ton of iron phosphate, and the costs of ammonium phosphate and 85% phosphoric acid account for 81.1% of the raw and auxiliary material costs. Therefore, traditional iron phosphate production processes generally use high-purity monoammonium phosphate and industrial-grade 85% phosphoric acid as raw materials, and the raw material costs are too high. Further exploring low-cost preparation processes for battery-grade iron phosphate has important research and application significance.

[0004] Wet-process phosphoric acid is phosphoric acid containing various impurities obtained by decomposing phosphate rock with strong acids such as sulfuric acid or hydrochloric acid and then performing liquid-solid separation. Wet-process phosphoric acid has low cost and wide sources, but unpurified wet-process phosphoric acid contains impurities such as Ca, Mg, Al, Fe, and F, and cannot be directly used in the production of battery-grade iron phosphate. For example, the existing patent CN 118929598 discloses a method for preparing iron phosphate and iron phosphate using phosphorus-containing wastewater precipitation slag. It uses an acid solution to leach the phosphorus-containing wastewater precipitation slag as the phosphorus source, uses by-product ferrous sulfate heptahydrate from titanium dioxide as the iron source, drops the acidic leaching solution into the iron source solution, and obtains a solid-liquid mixture after reaction; after filtering the solid-liquid mixture, H2O2 is added to the filtrate to obtain crude iron phosphate; 85% phosphoric acid is added to the crude iron phosphate in the step to obtain iron phosphate dihydrate. Although this technical solution uses the leaching solution of phosphorus-containing wastewater precipitation slag as the phosphorus source to save a certain amount of cost, expensive 85% phosphoric acid still has to be used in the subsequent steps.

[0005] Therefore, how to find a more suitable method to solve the above problems existing in the current battery-grade iron phosphate process has become one of the problems that many front-line researchers in the industry urgently need to solve. Summary of the Invention

[0006] In view of this, the technical problem to be solved by the present invention is to provide a method for co-producing iron phosphate from wet-process phosphoric acid by membrane separation. The present invention separates wet-process phosphoric acid into membrane-purified acid and concentrated acid through secondary membrane separation of wet-process phosphoric acid. The concentrated acid is subjected to auxiliary two-stage ammonia neutralization for impurity removal to obtain a phosphonium ammonium solution, and then ferrous sulfate and hydrogen peroxide are added for a synthesis reaction, followed by pressure filtration. The filter cake is washed and then made into a slurry with pure water, and the membrane-purified acid is added for a crystal transformation reaction. After pressure filtration, a crystal transformation filter cake is obtained. The filter cake is washed, dried, and calcined to obtain anhydrous battery-grade iron phosphate.

[0007] The present invention provides a method for co-producing iron phosphate from wet-process phosphoric acid by membrane separation, comprising the following steps:

[0008] 1) Using an acid-resistant nanofiltration membrane, after the first-stage membrane separation of wet-process phosphoric acid, concentrated acid and a first-stage clear phase are obtained;

[0009] 2) After the first-stage clear phase obtained in the above step is subjected to second-stage membrane separation through an acid-resistant nanofiltration membrane, a second-stage concentrated phase and purified acid are obtained;

[0010] 3) Mixing the purified acid, filter cake, and water obtained in the above steps for a crystal transformation reaction, and after the obtained slurry is filtered and dried, the obtained solid phase is calcined to obtain anhydrous iron phosphate;

[0011] The preparation process of the filter cake comprises the following steps:

[0012] a) After the concentrated acid, ammonia gas, and a co-precipitant are subjected to an ammonia neutralization reaction and filtration, a filtrate is obtained;

[0013] b) After the filtrate, hydrogen peroxide, and ferrous sulfate solution obtained in the above step are subjected to a synthesis reaction, filtration is carried out again to obtain a filter cake.

[0014] Preferably, by mass content, it includes P2O5: 22% - 23%, Fe: 650 - 700 ppm, Al: 500 - 520 ppm, Ca: 1430 - 1500 ppm, Mg: 3900 - 4100 ppm, Na: 170 - 200 ppm, K: 200 - 300 ppm, and SO4 2- : 35 - 45 ppm;

[0015] The flux of the first-stage membrane separation of the wet-process phosphoric acid is 30 - 35 LMH;

[0016] The operating pressure of the first-stage membrane separation is 5 - 6 MPa;

[0017] The temperature of the first-stage membrane separation is 40 - 45 °C;

[0018] The flux of the clear phase after the first-stage membrane separation is 30 - 35 LMH.

[0019] Preferably, the secondary dense phase returns to the primary membrane separation process in step 1).

[0020] The operating pressure of the secondary membrane separation is 4 - 5 MPa.

[0021] The temperature of the secondary membrane separation is 25 - 30 °C.

[0022] The flux of the clear phase after the secondary membrane separation is 50 - 60 LMH.

[0023] Preferably, in the concentrated acid after the primary membrane separation, by mass content, it includes P2O5: 23.5% - 24.5%, Fe: 2000 - 2100 ppm, Al: 1900 - 2000 ppm, Ca: 3300 - 3400 ppm, Mg: 8000 - 8300 ppm, Na: 1500 - 1600 ppm, K: 500 - 550 ppm, and SO4 2- : 50 - 60 ppm;

[0024] In the purified acid after the secondary membrane separation, by mass content, it includes P2O5: 21.5% - 22.5%, Fe: 0 ppm, Al: 0 ppm, Ca: 0 ppm, Mg: 0 - 1 ppm, Na: 10 - 25 ppm, K: 10 - 45 ppm, and SO4 2- : 0.5 - 1 ppm;

[0025] The mass ratio of the purified acid to the concentrated acid is (5 - 7):(5 - 3).

[0026] Preferably, the ammonia neutralization reaction includes a two-stage ammonia neutralization reaction process.

[0027] Step a) specifically includes the following steps:

[0028] a1) Mix the concentrated acid obtained in step 1) with a co-precipitant, then introduce ammonia gas to adjust the pH value, carry out the first-stage ammonia neutralization reaction, and then filter to obtain the first-stage ammonia neutralization filtrate.

[0029] a2) Mix the first-stage ammonia neutralization filtrate obtained in the above step with an alkaline pH regulator again, carry out the second-stage ammonia neutralization reaction, and then filter to obtain the filtrate.

[0030] The alkaline pH regulator includes ammonia gas and / or ammonia water.

[0031] The co-precipitant includes fluoride and silica white.

[0032] The fluoride includes one or more of ammonium fluoride, sodium fluoride, and potassium fluoride.

[0033] Preferably, the addition amount of the fluoride is 0.003% - 0.005% of the concentrated acid;

[0034] The addition amount of the precipitated silica is 0.001% - 0.003% of the concentrated acid;

[0035] In the step a1), the pH value is adjusted to 4.0 - 4.5;

[0036] The temperature of the first-stage ammonia neutralization reaction is 70 - 75 °C;

[0037] The time of the first-stage ammonia neutralization reaction is 20 - 30 min.

[0038] Preferably, the pH value of the second-stage ammonia neutralization reaction is controlled to be 6.0 - 6.5;

[0039] The temperature of the second-stage ammonia neutralization reaction is 75 - 80 °C;

[0040] The time of the second-stage ammonia neutralization reaction is 20 - 30 min;

[0041] In the step a), the filter cake obtained after filtration is ammonium phosphate compound fertilizer;

[0042] The ferrous sulfate is specifically the by-product ferrous sulfate heptahydrate produced in titanium dioxide production.

[0043] Preferably, the added mass of the hydrogen peroxide is 1 - 1.5 times the theoretical amount required for oxidizing ferrous sulfate;

[0044] The Fe element concentration of the ferrous sulfate solution is 6 wt% - 9 wt%;

[0045] In the step b), the ratio of nFe / nP in the ferrous sulfate to the filtrate is (1.03 - 1.18):1;

[0046] In the step b), the temperature of the synthesis reaction is 60 - 70 °C;

[0047] In the step b), the time of the synthesis reaction is 0.5 - 2 h;

[0048] In the step b), after the second filtration, a washing step is further included.

[0049] Preferably, the mass ratio of the filter cake to water is (0.8 - 1.2):2;

[0050] The ratio of the purified acid to the filter cake, calculated by nFe / nP, is (0.95 - 0.99):1;

[0051] In the mixing process of the step 3), a pH regulator is also added;

[0052] The adjusted pH value is 0.6 to 1;

[0053] The temperature of the crystal conversion reaction is 90 to 95 °C;

[0054] The time of the crystal conversion reaction is 2 to 4 h.

[0055] Preferably, in step 3), after filtration, there is also a washing step;

[0056] The drying temperature is 120 to 150 °C;

[0057] The drying time is 1.5 to 3 h;

[0058] The calcination process is: calcination at 250 to 300 °C for 1 to 1.5 h, calcination at 350 to 400 °C for 1 to 1.5 h, and calcination at 500 to 580 °C for 1 to 1.5 h;

[0059] The purity of the anhydrous iron phosphate is greater than or equal to 99.9%;

[0060] The anhydrous iron phosphate includes battery-grade iron phosphate.

[0061] The present invention provides a method for co-producing iron phosphate by membrane separation of wet-process phosphoric acid, which includes the following steps. First, a acid-resistant nanofiltration membrane is used to perform primary membrane separation on wet-process phosphoric acid to obtain concentrated acid and a primary clear phase; then the primary clear phase obtained in the above step is subjected to secondary membrane separation through an acid-resistant nanofiltration membrane to obtain a secondary concentrated phase and purified acid; finally, the purified acid, filter cake, and water obtained in the above steps are mixed for a crystal conversion reaction, and the obtained slurry is filtered and dried, and the obtained solid phase is then calcined to obtain anhydrous iron phosphate; the preparation process of the filter cake includes the following steps: concentrated acid, ammonia gas, and a co-precipitant are subjected to ammonia neutralization reaction and filtration to obtain a filtrate; the filtrate, hydrogen peroxide, and ferrous sulfate solution obtained in the above step are subjected to a synthesis reaction and then filtered again to obtain a filter cake. Compared with the prior art, the present invention creatively designs a process route for preparing iron phosphate based on membrane separation of wet-process phosphoric acid with a specific route and steps, which has rarely been reported in the previous iron phosphate production technologies. The present invention separates wet-process phosphoric acid into membrane-purified acid and concentrated acid through secondary membrane separation of wet-process phosphoric acid. The concentrated acid is subjected to auxiliary two-stage ammonia neutralization for impurity removal to obtain a phosphonium ammonium solution, which is added to a ferrous sulfate solution together with hydrogen peroxide for a synthesis reaction and then pressure-filtered. The filter cake is washed and made into a slurry with pure water, and membrane-purified acid is added for a crystal conversion reaction. After pressure filtration, a crystal conversion filter cake is obtained. Finally, the filter cake is dried and calcined to obtain anhydrous battery-grade iron phosphate.

[0062] The present invention uses a secondary acid-resistant nanofiltration membrane to perform membrane separation on wet-process phosphoric acid into purified acid and concentrated acid. Both the purified acid and the concentrated acid are used as raw materials for producing battery-grade iron phosphate, without raw material waste, achieving atom economy. Moreover, the membrane separation operation is simple, with low cost and low energy consumption. The acid-resistant nanofiltration membrane has a high rejection rate for impurity ions. After separation by two-stage acid-resistant nanofiltration membrane, the purified acid basically contains no impurity ions, and most of the impurity ions are retained in the concentrated acid. At the same time, through the method of membrane separation, the phosphorus combined with hydrogen ions is separated into the purified acid, and the phosphorus combined with impurity ions is separated into the concentrated acid. However, the concentrated acid still contains part of the phosphorus combined with hydrogen ions. Through the technology of auxiliary precipitation + ammonia neutralization, the phosphorus combined with impurity ions is precipitated into the filter cake and used as fertilizer. Through ammonia neutralization, the part of the phosphorus combined with hydrogen ions is converted into ammonium phosphate salt as the phosphorus source for the iron phosphate synthesis reaction.

[0063] The present invention realizes the maximum utilization of the effective components of phosphoric acid in wet-process phosphoric acid with high added value. The part of the phosphorus combined with impurity ions is used as fertilizer to exert its fertilizer effect. Moreover, the first-stage acid-resistant NF uses high temperature and high pressure to increase the membrane flux, thereby increasing the yield of the purified acid. The second-stage acid-resistant NF uses low temperature and high pressure to ensure the quality of the purified acid. In particular, the concentrated acid passes through the formulation combination of a co-precipitant, and then two-stage ammonia neutralization. Most of the impurity ions intercepted by the membrane in the concentrated acid are precipitated in the ammonium phosphate compound fertilizer after two-stage ammonia neutralization, making the impurity index of the filtrate lower, so that it can be used as a high-quality production raw material for the iron phosphate synthesis reaction.

[0064] The trial production results show that the unique membrane-purified acid raw material and auxiliary ammonia-neutralized raw material of the present invention make the prepared iron phosphate have a higher tapped density and a lower BET, which is more conducive to improving the electrochemical performance of lithium iron phosphate. At the same time, the latter-stage iron phosphate production process of the present invention is similar to the conventional process of iron phosphate, so the existing iron phosphate production equipment can be directly used for production, which can reduce the investment cost and rapidly expand the scale. BRIEF DESCRIPTION OF THE DRAWINGS

[0065] Figure 1 It is a schematic flow chart of the membrane separation of wet-process phosphoric acid and co-production of iron phosphate provided by the present invention. DETAILED DESCRIPTION OF THE INVENTION

[0066] In order to further understand the present invention, the preferred embodiments of the present invention will be described below in conjunction with examples. However, it should be understood that these descriptions are only for further explaining the features and advantages of the present invention, rather than limiting the claims of the invention.

[0067] For all raw materials of the present invention, there is no particular limitation on their sources, and they can be purchased on the market or prepared by conventional methods well-known to those skilled in the art.

[0068] For all raw materials of the present invention, there is no particular limitation on their purity. Preferably, industrial pure or conventional purity used in the wet-process phosphoric acid process is adopted in the present invention.

[0069] The present invention provides a method for co-producing phosphoric acid iron by membrane separation of wet-process phosphoric acid, comprising the following steps:

[0070] 1) Using an acid-resistant nanofiltration membrane, after performing primary membrane separation on wet-process phosphoric acid, concentrated acid and a primary clear phase are obtained;

[0071] 2) After subjecting the primary clear phase obtained in the above step to secondary membrane separation through an acid-resistant nanofiltration membrane, a secondary concentrated phase and purified acid are obtained;

[0072] 3) Mixing the purified acid, filter cake and water obtained in the above step for a conversion reaction, after the obtained slurry is filtered and dried, the obtained solid phase is calcined to obtain anhydrous iron phosphate;

[0073] The preparation process of the filter cake includes the following steps:

[0074] a) After subjecting concentrated acid, ammonia gas and a co-precipitant to an ammonia neutralization reaction and filtration, a filtrate is obtained;

[0075] b) After subjecting the filtrate, hydrogen peroxide and ferrous sulfate solution obtained in the above step to a synthesis reaction, filtration is performed again to obtain a filter cake.

[0076] The present invention first uses an acid-resistant nanofiltration membrane to perform primary membrane separation on wet-process phosphoric acid to obtain concentrated acid and a primary clear phase.

[0077] In the present invention, in the wet-process phosphoric acid, by mass content, preferably, it includes P2O5: 22% - 23%, Fe: 650 - 700 ppm, Al: 500 - 520 ppm, Ca: 1430 - 1500 ppm, Mg: 3900 - 4100 ppm, Na: 170 - 200 ppm, K: 200 - 300 ppm and SO4 2- : 35 - 45 ppm, more preferably P2O5: 22.2% - 22.8%, Fe: 660 - 690 ppm, Al: 504 - 516 ppm, Ca: 1440 - 1490 ppm, Mg: 3940 - 4060 ppm, Na: 175 - 195 ppm, K: 220 - 280 ppm and SO4 2- : 37 - 43 ppm, more preferably P2O5: 22.4% - 22.6%, Fe: 670 - 680 ppm, Al: 508 - 512 ppm, Ca: 1450 - 1480 ppm, Mg: 3980 - 4020 ppm, Na: 180 - 190 ppm, K: 240 - 260 ppm and SO4 2- : 39 - 41 ppm.

[0078] In the present invention, the flux of the first-stage membrane separation of the wet-process phosphoric acid is preferably 30 - 35 LMH, more preferably 31 - 34 LMH, and even more preferably 32 - 33 LMH.

[0079] In the present invention, the operating pressure of the first-stage membrane separation is preferably 5 - 6 MPa, more preferably 5.2 - 5.8 MPa, and even more preferably 5.4 - 5.6 MPa.

[0080] In the present invention, the temperature of the first-stage membrane separation is preferably 40 - 45 °C, more preferably 41 - 44 °C, and even more preferably 42 - 43 °C.

[0081] In the present invention, the flux of the clear phase after the first-stage membrane separation is preferably 30 - 35 LMH, more preferably 31 - 34 LMH, and even more preferably 32 - 33 LMH.

[0082] In the concentrated acid after the first-stage membrane separation in the present invention, by mass content, it preferably includes P2O5: 23.5% - 24.5%, Fe: 2000 - 2100 ppm, Al: 1900 - 2000 ppm, Ca: 3300 - 3400 ppm, Mg: 8000 - 8300 ppm, Na: 1500 - 1600 ppm, K: 500 - 550 ppm, and SO4 2- : 50 - 60 ppm, more preferably P2O5: 23.7% - 24.3%, Fe: 2020 - 2080 ppm, Al: 1920 - 1980 ppm, Ca: 3320 - 3380 ppm, Mg: 8050 - 8250 ppm, Na: 1520 - 1580 ppm, K: 510 - 540 ppm, and SO4 2- : 52 - 58 ppm, more preferably P2O5: 23.9% - 24.1%, Fe: 2040 - 2060 ppm, Al: 1940 - 1960 ppm, Ca: 3340 - 3360 ppm, Mg: 8100 - 8200 ppm, Na: 1540 - 1560 ppm, K: 520 - 530 ppm, and SO4 2- : 54 - 56 ppm.

[0083] In the present invention, the first-stage clear phase obtained in the above step is further subjected to second-stage membrane separation through an acid-resistant nanofiltration membrane to obtain a second-stage concentrated phase and purified acid.

[0084] In the present invention, the second-stage concentrated phase is preferably returned to the first-stage membrane separation process in step 1).

[0085] In the present invention, the operating pressure of the secondary membrane separation is preferably 4 to 5 MPa, more preferably 4.2 to 4.8 MPa, and even more preferably 4.4 to 4.6 MPa.

[0086] In the present invention, the temperature of the secondary membrane separation is preferably 25 to 30 °C, more preferably 26 to 29 °C, and even more preferably 27 to 28 °C.

[0087] In the present invention, the flux of the clear phase after the secondary membrane separation is preferably 50 to 60 LMH, more preferably 52 to 58 LMH, and even more preferably 54 to 56 LMH.

[0088] In the purified acid after the secondary membrane separation in the present invention, by mass content, it preferably includes P2O5: 21.5% to 22.5%, Fe: 0 ppm, Al: 0 ppm, Ca: 0 ppm, Mg: 0 to 1 ppm, Na: 10 to 25 ppm, K: 10 to 45 ppm, and SO4 2- : 0.5 to 1 ppm, more preferably P2O5: 21.7% to 22.3%, Fe: 0 ppm, Al: 0 ppm, Ca: 0 ppm, Mg: 0.2 to 0.8 ppm, Na: 13 to 22 ppm, K: 15 to 40 ppm, and SO4 2- : 0.6 to 0.9 ppm, more preferably P2O5: 21.9% to 22.1%, Fe: 0 ppm, Al: 0 ppm, Ca: 0 ppm, Mg: 0.4 to 0.6 ppm, Na: 16 to 19 ppm, K: 20 to 35 ppm, and SO4 2- : 0.7 to 0.8 ppm.

[0089] In the present invention, the mass ratio of the purified acid to the concentrated acid is preferably (5 to 7):(5 to 3), more preferably (5.5 to 6.5):(5 to 3), and even more preferably (5 to 7):(4.5 to 3.5). Specifically, if the mass ratio of the purified acid to the concentrated acid is too high (>7:3), it will lead to an increase in the impurity index of the purified acid and at the same time increase the difficulty of ammonia neutralization and impurity removal of the concentrated acid. If the mass ratio of the purified acid to the concentrated acid is too low (<5:5), it will reduce the yield of the purified acid and at the same time increase the treatment amount of the concentrated acid.

[0090] Finally, in the present invention, the purified acid, filter cake, and water obtained in the above steps are mixed for a crystal transformation reaction. After the obtained slurry is filtered and dried, the obtained solid phase is calcined to obtain anhydrous iron phosphate.

[0091] The preparation process of the filter cake in the present invention includes the following steps:

[0092] a) After subjecting the concentrated acid, ammonia gas, and co-precipitant to an ammonia neutralization reaction and filtration, a filtrate is obtained;

[0093] b) After performing a synthesis reaction on the filtrate, hydrogen peroxide, and ferrous sulfate solution obtained in the above steps, filter again to obtain a filter cake.

[0094] In the present invention, first, concentrated acid, ammonia gas, and a co-precipitating agent are subjected to an ammonia neutralization reaction and filtration to obtain a filtrate.

[0095] In the present invention, the ammonia neutralization reaction includes a two-stage ammonia neutralization reaction process;

[0096] The specific steps of step a) preferably include the following steps:

[0097] a1) After mixing the concentrated acid and the co-precipitating agent obtained in step 1), introduce ammonia gas to adjust the pH value, perform a first-stage ammonia neutralization reaction, and then filter to obtain a first-stage ammonia neutralization filtrate;

[0098] a2) After remixing the first-stage ammonia neutralization filtrate obtained in the above steps and an alkaline pH regulator, perform a second-stage ammonia neutralization reaction, and then filter to obtain a filtrate;

[0099] In the present invention, the alkaline pH regulator preferably includes ammonia gas and / or ammonia water, and more preferably ammonia gas or ammonia water.

[0100] In the present invention, the co-precipitating agent preferably includes fluoride and silica white.

[0101] In the present invention, the fluoride preferably includes one or more of ammonium fluoride, sodium fluoride, and potassium fluoride, and more preferably ammonium fluoride, sodium fluoride, or potassium fluoride.

[0102] In the present invention, the addition amount of the fluoride is preferably 0.003% - 0.005% of the concentrated acid, more preferably 0.0034% - 0.0046%, and even more preferably 0.0038% - 0.0042%. Here, the percentage is the mass content percentage.

[0103] In the present invention, the addition amount of the silica white is preferably 0.001% - 0.003% of the concentrated acid, more preferably 0.0014% - 0.0026%, and even more preferably 0.0018% - 0.0022%. Here, the percentage is the mass content percentage.

[0104] In the present invention, in step a1), the pH value is preferably adjusted to 4.0 - 4.5, more preferably 4.1 - 4.4, and even more preferably 4.2 - 4.3.

[0105] In the present invention, the temperature of the first-stage ammonia neutralization reaction is preferably 70 - 75°C, more preferably 71 - 74°C, and even more preferably 72 - 73°C.

[0106] In the present invention, the time of the first-stage ammonia neutralization reaction is preferably 20 - 30 min, more preferably 22 - 28 min, and even more preferably 24 - 26 min.

[0107] In the present invention, the filtrate of the first-stage ammonia neutralization obtained in the above step and the alkaline pH regulator are mixed again, and then a second-stage ammonia neutralization reaction is carried out. After filtration, a filtrate is obtained.

[0108] In the present invention, the pH value of the second-stage ammonia neutralization reaction is preferably controlled to be 6.0 - 6.5, more preferably 6.1 - 6.4, and even more preferably 6.2 - 6.3.

[0109] In the present invention, the temperature of the second-stage ammonia neutralization reaction is preferably 75 - 80 °C, more preferably 76 - 79 °C, and even more preferably 77 - 78 °C.

[0110] In the present invention, the time of the second-stage ammonia neutralization reaction is preferably 20 - 30 min, more preferably 22 - 28 min, and even more preferably 24 - 26 min.

[0111] In the present invention, in step a), the filter cake obtained after filtration is preferably ammonium phosphate compound fertilizer.

[0112] In the present invention, the ferrous sulfate is specifically preferably by-product ferrous sulfate heptahydrate produced in titanium dioxide production.

[0113] Finally, in the present invention, the filtrate, hydrogen peroxide, and ferrous sulfate solution obtained in the above steps are subjected to a synthesis reaction, and then filtered again to obtain a filter cake.

[0114] In the present invention, the added mass of the hydrogen peroxide is preferably 1 - 1.5 times the theoretical amount required for oxidizing ferrous sulfate, more preferably 1.1 - 1.4 times, and even more preferably 1.2 - 1.3 times.

[0115] In the present invention, the Fe element concentration of the ferrous sulfate solution is preferably 6 wt% - 9 wt%, more preferably 6.5 wt% - 8.5 wt%, and even more preferably 7 wt% - 8 wt%.

[0116] In the present invention, in step b), the nFe / nP of ferrous sulfate to the filtrate is preferably (1.03 - 1.18):1, more preferably (1.06 - 1.15):1, and even more preferably (1.09 - 1.12):1.

[0117] In the present invention, in step b), the temperature of the synthesis reaction is preferably 60 - 70 °C, more preferably 62 - 68 °C, and even more preferably 64 - 66 °C.

[0118] In the present invention, in the step b), the time of the synthesis reaction is preferably 0.5 to 2 h, more preferably 0.8 to 1.7 h, and even more preferably 1.1 to 1.4 h.

[0119] In the present invention, in the step b), after the secondary filtration, a washing step is further preferably included.

[0120] In the present invention, the mass ratio of the filter cake to water is preferably (0.8 to 1.2):2, more preferably (0.85 to 1.15):2, even more preferably (0.9 to 1.1):2, still more preferably (0.95 to 1.05):2, and specifically can be 1:2.

[0121] In the present invention, for the purified acid and the filter cake, in terms of nFe / nP, the ratio between the two is preferably (0.95 to 0.99):1, more preferably (0.955 to 0.985):1, even more preferably (0.96 to 0.98):1, and still more preferably (0.965 to 0.975):1.

[0122] In the present invention, during the mixing process in the step 3), a pH regulator is further preferably added.

[0123] In the present invention, the adjusted pH value is preferably 0.6 to 1, more preferably 0.65 to 0.95, even more preferably 0.7 to 0.9, and still more preferably 0.75 to 0.85.

[0124] In the present invention, the temperature of the crystal conversion reaction is preferably 90 to 95 °C, more preferably 91 to 94 °C, and even more preferably 92 to 93 °C.

[0125] In the present invention, the time of the crystal conversion reaction is preferably 2 to 4 h, more preferably 2.4 to 3.6 h, and even more preferably 2.8 to 3.2 h.

[0126] In the present invention, in the step 3), after the filtration, a washing step is further preferably included.

[0127] In the present invention, the drying temperature is preferably 120 to 150 °C, more preferably 125 to 145 °C, and even more preferably 130 to 140 °C.

[0128] In the present invention, the drying time is preferably 1.5 to 3 h, more preferably 1.8 to 2.7 h, and even more preferably 2.1 to 2.4 h.

[0129] In the present invention, the calcination process is preferably as follows: calcination at 250 - 300°C for 1 - 1.5 h, calcination at 350 - 400°C for 1 - 1.5 h, calcination at 500 - 580°C for 1 - 1.5 h; more preferably, calcination at 260 - 290°C for 1.1 - 1.4 h, calcination at 360 - 390°C for 1.1 - 1.4 h, calcination at 515 - 565°C for 1.1 - 1.4 h; even more preferably, calcination at 270 - 280°C for 1.2 - 1.3 h, calcination at 370 - 380°C for 1.2 - 1.3 h, calcination at 530 - 550°C for 1.2 - 1.3 h.

[0130] In the present invention, the purity of the anhydrous iron phosphate is preferably greater than or equal to 99.9%.

[0131] In the present invention, the anhydrous iron phosphate preferably includes battery-grade iron phosphate.

[0132] To complete and refine the overall technical solution of the present invention, better ensure the parameters and grades of the iron phosphate product, and further improve the stability and efficiency of the membrane separation wet-process phosphoric acid co-production of iron phosphate, the method for membrane separation wet-process phosphoric acid co-production of iron phosphate specifically may include the following:

[0133] See Figure 1 , Figure 1 which is a schematic diagram of the process flow for the membrane separation wet-process phosphoric acid co-production of iron phosphate provided by the present invention.

[0134] As shown in the process flow diagram Figure 1 , the wet-process phosphoric acid from the phosphoric acid workshop enters the first-stage acid-resistant nanofiltration membrane for membrane separation. After the first-stage membrane separation, concentrated acid and the first-stage clear phase are obtained. The first-stage clear phase enters the second-stage acid-resistant nanofiltration membrane, and the concentrated acid enters the second-stage ammonia neutralization as the phosphorus source for the synthesis reaction. After the second-stage membrane separation, the second-stage concentrated phase and purified acid are obtained. The second-stage concentrated phase returns to the first-stage feed port, and the purified acid serves as the raw material acid for the crystal transformation reaction.

[0135] Ammonia is added to the concentrated acid for second-stage ammonia neutralization, followed by filtration. The filter cake is used as ammonium phosphate compound fertilizer, and the filtrate is added to the ferrous sulfate solution together with hydrogen peroxide for the synthesis reaction. After the reaction, pressure filtration is carried out. The filter cake obtained from the pressure filtration is pulped and then added with the purified acid from the second-stage membrane separation and a small amount of sulfuric acid for the crystal transformation reaction. After the reaction, pressure filtration, drying, and calcination are carried out to obtain the battery-grade iron phosphate product.

[0136] Specifically, in the second-stage ammonia neutralization, it is optimal to use ammonia for pH adjustment, and other alkaline substances such as ammonia water can be used as the pH regulator.

[0137] Specifically, one of the optimal compositions of the co-precipitant is ammonium fluoride, and other fluorides such as sodium fluoride and potassium fluoride can be used for substitution.

[0138] Specifically, the acid-resistant NF membrane is a two-stage NF system. The concentrated phase of the second stage returns to the first stage, and the concentrated acid from the first stage enters the ammonia neutralization process of the second stage.

[0139] The operating pressure of the first-stage acid-resistant NF is 5 - 6 MPa, the temperature is 40 - 45 °C, and the clean flux is greater than 30 LMH. The operating pressure of the second-stage acid-resistant NF system is 4 - 5 MPa, the temperature is 25 - 30 °C, and the clean flux is greater than 50 LMH. Since the acid-resistant NF membrane has a high rejection rate for cations, especially polyvalent cations, after passing through the second-stage NF membrane system, impurity cations can hardly be detected in the purified acid. The purified acid passing through the second-stage acid-resistant NF membrane enters the crystal conversion reaction unit.

[0140] Specifically, ammonia neutralization is carried out in two stages. After each stage of ammonia neutralization reaction, filtration is carried out. The filtrate is used as the stock solution for the next stage of ammonia neutralization, and the filter cake is used as ammonium phosphate compound fertilizer.

[0141] Specifically, in the first-stage ammonia neutralization, a co-precipitant is added first, and then the pH is adjusted by adding ammonia gas. The pH is controlled at 4.0 - 4.5, the temperature is 70 - 75 °C, and the reaction time is 20 - 30 min.

[0142] Specifically, the co-precipitant formula is 0.003% - 0.005% ammonium fluoride and 0.001% - 0.003% silica white. Preferably, the pH of the second-stage ammonia neutralization is controlled at 6.0 - 6.5, the temperature is 75 - 80 °C, and the reaction time is 20 - 30 min.

[0143] Specifically, the filtrate after ammonia neutralization and hydrogen peroxide are added together to the by-product ferrous sulfate heptahydrate of titanium dioxide. The addition amount of hydrogen peroxide is 1 - 1.5 times the theoretical amount required for oxidizing ferrous sulfate. The Fe concentration of the ferrous sulfate solution is 6 - 9 wt%. The nFe / nP in the filtrate after ammonia neutralization and ferrous sulfate is 1.03 - 1.18. The reaction temperature is 60 - 70 °C, and the reaction time is 0.5 - 2 h.

[0144] Specifically, after the reaction, filtration and washing are carried out. The filter cake (wet basis) and pure water are pulped according to a mass ratio of 1:2. Membrane-purified acid is added, and the mass ratio of the purified acid to the filter cake is 0.2 - 0.3. The nFe / nP is adjusted to 0.95 - 0.99, and a small amount of sulfuric acid is added to adjust the solution pH to 0.6 - 1. The reaction temperature is 90 - 95 °C, and the reaction time is 2 - 4 h. After the reaction, the slurry is filtered, washed, dried, and calcined to obtain anhydrous iron phosphate products.

[0145] In the raw and auxiliary material costs of iron phosphate, the phosphorus source raw materials (ammonium phosphate and 85% phosphoric acid) account for the largest proportion. In the ammonium method process, 0.78 t of ammonium phosphate and 0.13 t of 85% phosphoric acid are consumed to produce one ton of iron phosphate, and the costs of ammonium phosphate and 85% phosphoric acid account for 81.1% of the raw and auxiliary material costs. By adopting the present invention, wet-process phosphoric acid with higher impurity content and lower cost can be used to completely replace ammonium phosphate and 85% phosphoric acid required for the production of iron phosphate.

[0146] In the present invention, the wet-process phosphoric acid is subjected to secondary membrane separation to separate the wet-process phosphoric acid with higher impurity content and lower cost into membrane-purified acid and concentrated acid. The concentrated acid is used as the phosphorus source for the synthesis reaction after being subjected to auxiliary two-stage ammonia neutralization for impurity removal, and the membrane-purified acid is used as the raw material for the crystal transformation reaction. The present invention makes full use of the wet-process phosphoric acid with higher impurity content, does not require expensive 85% phosphoric acid and monoammonium phosphate, and has lower raw material costs. Moreover, the first-stage acid-resistant NF system has high pressure and high temperature, which can ensure the recovery rate and production efficiency of the purified acid, and the second-stage acid-resistant NF system has a lower temperature, which can ensure the quality of the purified acid. Further, after the concentrated acid is subjected to a co-precipitant and two-stage ammonia neutralization, the impurity ions retained by the membrane in the concentrated acid can be precipitated as ammonium phosphate compound fertilizer, and the impurity content in the filtrate is very low and can be used as the phosphorus source for the production of iron phosphate. The produced ammonium phosphate compound fertilizer contains relatively high contents of phosphorus and nitrogen, as well as metal elements such as calcium, magnesium, sodium, and potassium, and has high fertilizer efficiency.

[0147] The specific membrane-purified acid raw material and auxiliary ammonia neutralization raw material of the present invention enable the produced battery-grade iron phosphate to have a higher tap density and a lower BET, which is more conducive to improving the electrochemical performance of the subsequent cathode material lithium iron phosphate. The present invention has the advantages of simple process, low operating cost, high product added value, and high economic benefits. The subsequent process of this application can directly use the existing iron phosphate production device for production, which can reduce the investment cost and rapidly expand the scale.

[0148] The above content of the present invention provides a method for co-producing iron phosphate by membrane separation of wet-process phosphoric acid. The specific route and steps designed by the present invention are based on the process route for preparing iron phosphate after membrane separation of wet-process phosphoric acid, and there are few reports in the previous iron phosphate production technologies. In the present invention, the wet-process phosphoric acid is subjected to secondary membrane separation to separate the wet-process phosphoric acid into membrane-purified acid and concentrated acid. The concentrated acid is subjected to auxiliary two-stage ammonia neutralization for impurity removal to obtain an ammonium phosphate solution, which is added to a ferrous sulfate solution together with hydrogen peroxide for a synthesis reaction and then pressure-filtered. The filter cake is washed and then made into a slurry with pure water, and the membrane-purified acid is added for a crystal transformation reaction. After pressure filtration, a crystal transformation filter cake is obtained. Finally, the filter cake is dried and calcined to obtain anhydrous battery-grade iron phosphate.

[0149] The present invention uses a secondary acid-resistant nanofiltration membrane to perform membrane separation on wet-process phosphoric acid into purified acid and concentrated acid. Both the purified acid and the concentrated acid are used as raw materials for producing battery-grade iron phosphate, without raw material waste, achieving atom economy. Moreover, the membrane separation operation is simple, with low cost and low energy consumption. The acid-resistant nanofiltration membrane has a high rejection rate for impurity ions. After separation by two-stage acid-resistant nanofiltration membrane, the purified acid basically contains no impurity ions, and most of the impurity ions are retained in the concentrated acid. At the same time, through the membrane separation method, the phosphorus combined with hydrogen ions is separated into the purified acid, and the phosphorus combined with impurity ions is separated into the concentrated acid. However, the concentrated acid still contains some phosphorus combined with hydrogen ions. Through the technology of auxiliary precipitation + ammonia neutralization, the phosphorus combined with impurity ions is precipitated into the filter cake and used as fertilizer. Through ammonia neutralization, the part of phosphorus combined with hydrogen ions is converted into ammonium phosphate salt as the phosphorus source for the iron phosphate synthesis reaction.

[0150] The present invention realizes the maximum utilization of the effective components of phosphoric acid in wet-process phosphoric acid with high added value. The part of phosphorus combined with impurity ions is used as fertilizer to exert its fertilizer effect. Moreover, the first-stage acid-resistant NF uses high temperature and high pressure to increase the membrane flux, thereby increasing the yield of the purified acid. The second-stage acid-resistant NF uses low temperature and high pressure to ensure the quality of the purified acid. In particular, the concentrated acid passes through the formulation combination of a co-precipitant, and then two-stage ammonia neutralization. Most of the impurity ions retained by the membrane in the concentrated acid are precipitated in the ammonium phosphate compound fertilizer after two-stage ammonia neutralization, making the impurity index of the filtrate lower, so that it can be used as a high-quality production raw material for the iron phosphate synthesis reaction.

[0151] The trial production results show that the unique membrane-purified acid raw material and auxiliary ammonia-neutralized raw material of the present invention make the prepared iron phosphate have a higher tap density and a lower BET, which is more conducive to improving the electrochemical performance of lithium iron phosphate. At the same time, the latter-stage iron phosphate production process of the present invention is similar to the conventional process of iron phosphate, so the existing iron phosphate production equipment can be directly used for production, which can reduce the investment cost and rapidly expand the scale.

[0152] To further illustrate the present invention, the following describes in detail a process for membrane separation of wet-process phosphoric acid and co-production of iron phosphate provided by the present invention in combination with embodiments. However, it should be understood that these embodiments are implemented on the premise of the technical solution of the present invention, and the detailed implementation methods and specific operation processes are given. It is only to further illustrate the features and advantages of the present invention, rather than a limitation on the claims of the present invention. The protection scope of the present invention is not limited to the following embodiments.

[0153] Example 1

[0154] The wet-process phosphoric acid from the phosphoric acid workshop enters the first-stage acid-resistant nanofiltration membrane for membrane separation. After the first-stage membrane separation, concentrated acid and the first-stage clear phase are obtained. The first-stage clear phase enters the second-stage acid-resistant nanofiltration membrane, and the concentrated acid enters the second-stage ammonia neutralization as the phosphorus source for the synthesis reaction. After the second-stage membrane separation, the second-stage concentrated phase and purified acid are obtained. The second-stage concentrated phase returns to the first-stage feed inlet, and the purified acid serves as the raw material acid for the crystal transformation reaction.

[0155] Ammonia gas is added to the concentrated acid for the second-stage ammonia neutralization, followed by filtration. The filter cake is used as ammonium phosphate compound fertilizer. The filtrate is added to the ferrous sulfate solution together with hydrogen peroxide for the synthesis reaction. After the reaction, pressure filtration is carried out. The filter cake obtained from the pressure filtration is pulped and then added with the purified acid from the second-stage membrane separation and a small amount of sulfuric acid for the crystal transformation reaction. After the reaction, pressure filtration, drying, and calcination are carried out to obtain the battery-grade iron phosphate product.

[0156] Refer to Table 1. Table 1 shows the indicators of the wet-process phosphoric acid from the phosphoric acid workshop in the embodiments of the present invention, which is the phosphoric acid composition table.

[0157] Table 1

[0158] <![CDATA[P2O5]]> ％ 23 Fe ppm 680 Al ppm 512 Ca ppm 1453 Mg ppm 4015 Na ppm 181 K ppm 295 <![CDATA[SO4 2- > ppm 43

[0159] The acid-resistant NF membrane is a two-stage NF system. The second-stage concentrated phase returns to the first stage, and the first-stage concentrated acid enters the second-stage ammonia neutralization process. The operating pressure of the first-stage acid-resistant NF is 5.5 MPa, the temperature is 43 °C, and the clear-phase flux is 35.2 LMH. The operating pressure of the second-stage acid-resistant NF system is 4.5 MPa, the temperature is 27 °C, and the clear-phase flux is 58 LMH. Since the acid-resistant NF membrane has a very high rejection rate for cations, especially polyvalent cations, after passing through the second-stage NF membrane system, impurity cations can hardly be detected in the purified acid. The purified acid passing through the second-stage acid-resistant NF membrane enters the crystal transformation reaction unit. Refer to Table 2. Table 2 shows the composition table of the purified acid after membrane separation in Example 1 of the present invention.

[0160] Table 2

[0161] <![CDATA[P2O5]]> ％ 22 Fe ppm Not detected Al ppm Not detected Ca ppm Not detected Mg ppm 1 Na ppm 16 K ppm 27 <![CDATA[SO4 2- > ppm 0.71

[0162] Refer to Table 3. Table 3 shows the composition table of the concentrated acid after the first-stage membrane separation in Example 1 of the present invention.

[0163] Table 3

[0164] <![CDATA[P2O5]]> ％ 24 Fe ppm 2078 Al ppm 1950 Ca ppm 3382 Mg ppm 8212 Na ppm 1516 K ppm 523 <![CDATA[SO4 2- > ppm 58

[0165] Refer to Table 4. Table 4 shows the rejection rate of the first-stage acid-resistant NF membrane in Example 1 of the present invention.

[0166] Table 4

[0167]

[0168]

[0169] Refer to Table 5. Table 5 shows the rejection rate of the secondary acid-resistant NF membrane in Example 1 of the present invention.

[0170] Table 5

[0171] Ionic components Rejection rate (%) <![CDATA[H3PO4]]> 3.1 Fe 100 Al 100 Ca 100 Mg 99.8 Na 75.3 K 76.2

[0172] Ammonia neutralization is carried out in two stages. After each stage of ammonia neutralization reaction, filtration is carried out. The filtrate is used as the stock solution for the next stage of ammonia neutralization, and the filter cake is used as ammonium phosphate compound fertilizer. In particular, in the first stage of ammonia neutralization, a co-precipitant is added first, and then the pH is adjusted by adding ammonia gas. The pH is controlled at 4.3, the temperature is 73 °C, and the reaction time is 26 min. The formula of the co-precipitant is 0.004% ammonium fluoride and 0.002% silica white. The pH of the second stage of ammonia neutralization is controlled at 6.4, the temperature is 78 °C, and the reaction time is 27 min.

[0173] Refer to Table 6. Table 6 shows the filter cake indexes of ammonia neutralization in Example 1 of the present invention.

[0174] Table 6

[0175] First-stage filter cake Second-stage filter cake <![CDATA[P2O5]]> ％ 31.19 32.15 Ammonia N ％ 6.61 9.86 Fe ppm 11219 672 As ppm 0.3 0.3 Pd ppm Not detected 0.2 Mn ppm 266 1216 Ni ppm 5 38 <![CDATA[SiO2]]> ppm 187 13 K ppm 1008 231 Na ppm 1782 1116 Mg ppm 13678 123870 Ti ppm 254 2 Cr ppm 110 6 Cu ppm 0.6 0.5 Al ppm 13205 358 Ca ppm 6018 19025 F ppm 6815 3788 <![CDATA[SO4 2- > ppm 228 573

[0176] Refer to Table 7. Table 7 shows the filtrate indexes of ammonia neutralization in Example 1 of the present invention.

[0177] Table 7

[0178] <![CDATA[P2O5]]> ％ 23 Fe ppm 0.4 As ppm 0.5 Pd ppm Not detected Mn ppm 0.6 Ni ppm 0.2 <![CDATA[SiO2]]> ppm 45.3 K ppm 348 Na ppm 890 Mg ppm 143 Ti ppm 0.06 Cr ppm 2 Cu ppm 0.02 Al ppm 8 Ca ppm 16 F ppm 46 <![CDATA[SO4 2- > ppm 12

[0179] The filtrate after ammonia neutralization and hydrogen peroxide are added together to the by-product ferrous sulfate heptahydrate of titanium dioxide. The addition amount of hydrogen peroxide is 1.3 times the theoretical amount required to oxidize ferrous sulfate. The Fe concentration of the ferrous sulfate solution is 7.8 wt%. The nFe / nP in the filtrate after ammonia neutralization and ferrous sulfate is 1.05. The reaction temperature is 65 °C, and the reaction time is 1 h.

[0180] After the reaction, filtration and washing are carried out. The filter cake (wet basis) and pure water are made into a slurry according to a mass ratio of 1:2. Membrane purification acid is added. The mass ratio of the purification acid to the filter cake is 0.25. The nFe / nP is adjusted to 0.97, and a small amount of sulfuric acid is added to adjust the pH of the solution to 0.8. The reaction temperature is 94 °C, and the reaction time is 3.5 h. After the reaction, the slurry is filtered, washed, dried, and calcined (calcined at 275 °C for 1.2 h, 376 °C for 1.2 h, and 540 °C for 1.3 h) to obtain anhydrous iron phosphate products.

[0181] Refer to Table 8. Table 8 shows the indexes of the anhydrous iron phosphate products prepared in Example 1 of the present invention.

[0182] Table 8

[0183]

[0184] The above has introduced in detail a process for co-producing iron phosphate by membrane separation wet-process phosphoric acid provided by the present invention. Specific examples are used in this article to elaborate on the principle and implementation manner of the present invention. The description of the above embodiments is only used to help understand the method and its core idea of the present invention, including the best mode, and also enables any person skilled in the art to practice the present invention, including manufacturing and using any device or system, and implementing any combined method. It should be noted that for those of ordinary skill in the art in this technical field, without departing from the principle of the present invention, several improvements and modifications can be made to the present invention, and these improvements and modifications also fall within the protection scope of the claims of the present invention. The protection scope of the present invention patent is defined by the claims and may include other embodiments that can be conceived by those skilled in the art. If these other embodiments have structural elements that are not different from the literal description of the claims, or if they include equivalent structural elements that have no substantial difference from the literal description of the claims, then these other embodiments should also be included within the scope of the claims.

Claims

1. A method for co-producing iron phosphate by membrane separation of wet-process phosphoric acid, characterized in that, It includes the following steps: 1) Using an acid-resistant nanofiltration membrane, after performing primary membrane separation on wet-process phosphoric acid, concentrated acid and a primary clear phase are obtained; 2) After subjecting the primary clear phase obtained in the above step to secondary membrane separation through an acid-resistant nanofiltration membrane, a secondary concentrated phase and purified acid are obtained; 3) Mixing the purified acid, filter cake and water obtained in the above step for a crystal transformation reaction, and after filtering and drying the resulting slurry, the obtained solid phase is calcined to obtain anhydrous iron phosphate; The preparation process of the filter cake includes the following steps: a) After subjecting concentrated acid, ammonia gas and a co-precipitant to an ammonia neutralization reaction and filtration, a filtrate is obtained; b) After subjecting the filtrate obtained in the above step, hydrogen peroxide and a ferrous sulfate solution to a synthesis reaction, and then filtering again, a filter cake is obtained.

2. The method according to claim 1, wherein In the wet-process phosphoric acid, by mass content, it includes P2O5: 22% - 23%, Fe: 650 - 700 ppm, Al: 500 - 520 ppm, Ca: 1430 - 1500 ppm, Mg: 3900 - 4100 ppm, Na: 170 - 200 ppm, K: 200 - 300 ppm, and SO4 2- : 35 - 45 ppm; The flux of the primary membrane separation of the wet-process phosphoric acid is 30 - 35 LMH; The operating pressure of the primary membrane separation is 5 - 6 MPa; The temperature of the primary membrane separation is 40 - 45 °C; The flux of the clear phase after the primary membrane separation is 30 - 35 LMH.

3. The method according to claim 1, wherein The secondary concentrated phase is returned to the primary membrane separation process in step 1); The operating pressure of the secondary membrane separation is 4 - 5 MPa; The temperature of the secondary membrane separation is 25 - 30 °C; The flux of the clear phase after the secondary membrane separation is 50 - 60 LMH.

4. The method according to claim 1, characterized in that In the concentrated acid after the first-stage membrane separation, by mass content, it includes P2O5: 23.5% - 24.5%, Fe: 2000 - 2100 ppm, Al: 1900 - 2000 ppm, Ca: 3300 - 3400 ppm, Mg: 8000 - 8300 ppm, Na: 1500 - 1600 ppm, K: 500 - 550 ppm, and SO4 2- : 50 - 60 ppm; In the purified acid after the secondary membrane separation, by mass content, it includes P2O5: 21.5% to 22.5%, Fe: 0 ppm, Al: 0 ppm, Ca: 0 ppm, Mg: 0 to 1 ppm, Na: 10 to 25 ppm, K: 10 to 45 ppm, and SO4 2- : 0.5 to 1 ppm; The mass ratio of the purified acid to the concentrated acid is (5 - 7):(5 - 3).

5. The method according to claim 1, wherein The ammonia neutralization reaction includes a two-stage ammonia neutralization reaction process; Step a) specifically includes the following steps: a1) After mixing the concentrated acid and the co-precipitant obtained in step 1), then introducing ammonia gas to adjust the pH value, performing a first-stage ammonia neutralization reaction, and then filtering to obtain a first-stage ammonia neutralization filtrate; a2) After remixing the first-stage ammonia neutralization filtrate obtained in the above step and an alkaline pH regulator, performing a second-stage ammonia neutralization reaction, and then filtering to obtain a filtrate; The alkaline pH regulator includes ammonia gas and / or ammonia water; The co-precipitant includes fluoride and silica white; The fluoride includes one or more of ammonium fluoride, sodium fluoride and potassium fluoride.

6. The method according to claim 5, wherein The addition amount of the fluoride is 0.003% - 0.005% of the concentrated acid; The addition amount of the silica white is 0.001% - 0.003% of the concentrated acid; In step a1), the pH value is adjusted to 4.0 - 4.5; The temperature of the first-stage ammonia neutralization reaction is 70 - 75 °C; The time of the first-stage ammonia neutralization reaction is 20 - 30 min.

7. The method according to claim 5, wherein The pH value of the second-stage ammonia neutralization reaction is controlled at 6.0 - 6.5; The temperature of the second-stage ammonia neutralization reaction is 75 - 80 °C; The time of the second-stage ammonia neutralization reaction is 20 - 30 min; In step a), the filter cake obtained after filtration is ammonium phosphate compound fertilizer; The ferrous sulfate is specifically the by-product heptahydrate ferrous sulfate from titanium dioxide production.

8. The method according to claim 1, wherein The added mass of the hydrogen peroxide is 1 - 1.5 times the theoretical amount required to oxidize the ferrous sulfate; The Fe element concentration of the ferrous sulfate solution is 6 wt% - 9 wt%; In step b), the nFe / nP of the ferrous sulfate to the filtrate is (1.03 - 1.18):1; In step b), the temperature of the synthesis reaction is 60 - 70 °C; In the step b), the time of the synthesis reaction is 0.5 to 2 h; In the step b), after the re-filtration, a washing step is further included.

9. The method according to claim 1, wherein The mass ratio of the filter cake to water is (0.8 to 1.2):2; For the purified acid and the filter cake, in terms of nFe / nP, the ratio of the two is (0.95 to 0.99):1; In the mixing process of the step 3), a pH regulator is further added; The adjusted pH value is 0.6 to 1; The temperature of the crystal transformation reaction is 90 to 95 °C; The time of the crystal transformation reaction is 2 to 4 h.

10. The method according to claim 1, characterized in that, In the step 3), after the filtration, a washing step is further included; The temperature of the drying is 120 to 150 °C; The time of the drying is 1.5 to 3 h; The calcination process is: calcination at 250 to 300 °C for 1 to 1.5 h, calcination at 350 to 400 °C for 1 to 1.5 h, and calcination at 500 to 580 °C for 1 to 1.5 h; The purity of the anhydrous iron phosphate is greater than or equal to 99.9%; The anhydrous iron phosphate includes battery-grade iron phosphate.

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