Salt separation and crystallization process and system for high-salinity wastewater

Abstract

The invention relates to the field of industrial wastewater treatment, in particular to a high-salinity wastewater salt separation and crystallization process and system.The process comprises the steps that hardness and silicon removal treatment, turbidity removal treatment, softening treatment, membrane treatment, ozone oxidation treatment, evaporation and concentration treatment and freezing crystallization treatment are conducted on wastewater, and mirabilite and freezing mother liquor are obtained; wherein the mirabilite is subjected to melt crystallization treatment to obtain a sodium sulfate product and a melt mother solution; the molten mother liquor is repeatedly subjected to melt crystallization treatment, and a sodium sulfate product is obtained; performing circulating sodium chloride crystallization treatment on the frozen mother liquor to obtain a sodium chloride product and sodium chloride mother liquor; the sodium chloride mother liquor is subjected to circulating carnallite crystallization treatment and drum drying, and a carnallite product is obtained. The process solves the problem of crystallization and salt separation of the coal chemical industry wastewater, and realizes cyclic utilization of the coal chemical industry wastewater, reduction of carnallite and resource cyclic economic utilization.

Description

Technical Field

[0001] The invention relates to the technical field of industrial wastewater treatment, and in particular to a high-salt wastewater salt crystallization process and system. Background Art

[0002] Wastewater from industries such as petrochemicals and coal chemicals has a high salt content and complex composition, which poses a considerable threat to the sewage treatment system. In order to achieve "near-zero discharge" of wastewater, strict treatment of these wastewaters is inevitable. However, the miscellaneous salts generated during the treatment process are regarded as solid hazardous wastes, and the treatment cost is high, which brings some burdens to enterprises.

[0003] In order to solve this problem, chemical companies that have achieved zero emissions in the early stage need to replace the original mixed salt crystallization technology with salt separation crystallization technology to realize the resource utilization of salt. Among the existing treatment methods, nanofiltration membrane method and traditional thermal crystallization salt separation technology are more commonly used. However, with the continuous advancement of technology, a new low-temperature critical freezing vacuum thermal salt separation crystallization process has emerged.

[0004] However, in actual application, this new process has some problems: First, the operating temperature of the frozen crystallizer cannot meet the design standard (-4.5℃~-5.0℃), resulting in a low extraction rate of sodium sulfate. Second, the frozen crystallizer will experience freezing and "crystal explosion" phenomenon, the operating cycle is short and unstable, and frequent thawing is required. Third, because the sodium sulfate extraction rate is low (only about 95%), a large amount of sodium sulfate will enter the sodium chloride crystallization system, resulting in the quality of the sodium chloride product salt being unqualified and difficult to meet the requirements of refined industrial first-class products. These problems not only affect the production efficiency of the enterprise, but also have an adverse impact on product quality. Summary of the invention

[0005] The purpose of the present invention is to solve the problems faced by the low-temperature critical freezing vacuum thermal salt separation crystallization process in the background technology, such as low sodium sulfate extraction rate, unqualified salt quality of sodium chloride product and short operation cycle of the freezing crystallizer, so as to provide a high-salt wastewater salt separation crystallization process and system.

[0006] To achieve the above-mentioned object of the invention, the first aspect of the present invention provides a process for salt crystallization of high-salt wastewater, the process comprising the following steps:

[0007] S1: removing hardness and silicon from high-salinity wastewater to obtain hardness and silicon-free effluent;

[0008] S2: performing turbidity removal treatment on the hardness and silicon removal effluent to obtain turbidity-removed effluent;

[0009] S3: subjecting the deturbidified effluent to ion exchange softening treatment to obtain softened effluent;

[0010] S4: performing membrane treatment on the softened effluent to obtain membrane treated effluent;

[0011] S5: subjecting the membrane treated effluent to ozone oxidation treatment to obtain ozone oxidized effluent;

[0012] S6: performing evaporation and concentration treatment on the ozone oxidation effluent to obtain evaporation and concentration effluent;

[0013] S7: subjecting the evaporated concentrated water to freezing crystallization treatment to obtain Glauber's salt and frozen mother liquor, wherein the frozen mother liquor is divided into two parts, 85% to 90% of which is sent to step S9 for sodium chloride crystallization treatment, and the remaining 10% to 15% is repeatedly subjected to freezing crystallization treatment;

[0014] S8; subjecting the mirabilite to melt crystallization to obtain a sodium sulfate product and a molten mother liquor, and repeatedly subjecting the molten mother liquor to melt crystallization;

[0015] S9: subjecting the frozen mother liquor obtained in step S7 to sodium chloride crystallization to obtain a sodium chloride product and a sodium chloride mother liquor, wherein the sodium chloride mother liquor is divided into two parts, 5% to 9% of which is sent to step S10 for impurity salt crystallization, and the remaining 91% to 95% is repeatedly subjected to sodium chloride crystallization;

[0016] S10: subjecting the sodium chloride mother liquor obtained in step S9 to a mixed salt crystallization treatment to obtain a mixed salt product.

[0017] The second aspect of the present invention provides a system for the above-mentioned high-salt wastewater salt crystallization process, and the system includes the following units according to the material flow:

[0018] The hardness and silicon removal unit is used to remove most of the hardness and silicon in the high-salinity wastewater to obtain hardness and silicon-free effluent;

[0019] Turbidity removal unit, used to remove suspended matter in the effluent from the hardness removal and silicon removal process to obtain turbidity removal effluent;

[0020] A softening unit is used to remove the remaining hardness in the deturbidified effluent water to obtain softened effluent water;

[0021] The membrane treatment unit is used to remove the precipitated colloids, suspended solids and reduced matter in the softened effluent to obtain membrane treated effluent;

[0022] Ozone oxidation unit, used to remove organic matter in membrane treated water to obtain ozone oxidized water;

[0023] An evaporation concentration unit is used to concentrate the ozone oxidation effluent to obtain evaporation concentrated effluent;

[0024] A freezing crystallization unit is used to freeze and crystallize the evaporation and concentration water to obtain sodium sulfate and frozen mother liquor;

[0025] A melt crystallization unit is used to melt crystallize the mirabilite to obtain a sodium sulfate product and a molten mother liquor;

[0026] A sodium chloride crystallization unit is used to crystallize sodium chloride from the frozen mother liquor to obtain a sodium chloride product and a sodium chloride mother liquor;

[0027] The mixed salt crystallization unit is used to process part of the mother liquor discharged from the sodium chloride crystallization unit to obtain mixed salt products.

[0028] The present invention has the following beneficial effects:

[0029] 1. The present invention solves the problem of salt separation and crystallization of high-salt wastewater from coal chemical industry, realizes the recycling of coal chemical wastewater, the reduction of impurity salts, and the resource-based circular economic utilization, and produces first-class sodium sulfate products that meet national standards and refined industrial salt first-class sodium chloride products.

[0030] 2. In the present invention, calcium agent, magnesium agent, coagulant and cation exchanger are used in sequence to treat high-salt wastewater, so that the hardness of the obtained de-hardened effluent is ≤5mg / L.

[0031] 3. In the membrane treatment of the present invention, the precipitated colloids and suspended matter in the wastewater are removed by ultrafiltration to reduce the turbidity and sludge density, and then the wastewater is reduced by reverse osmosis treatment to reduce the processing capacity of each crystallization unit during the subsequent salt separation crystallization.

[0032] Fourth, in the evaporation concentration treatment of the present invention, the ozone oxidation effluent is evaporated and concentrated by a TVR device and an MVR device, so that the soluble solid content in the brine is increased to 200-230 g / L, further reducing the scale of subsequent sodium sulfate crystallization and sodium chloride crystallization.

[0033] 5. In the freeze-crystallization of the present invention, in the case where the actual salt-nitrate ratio is significantly lower than the designed salt-nitrate ratio of the freeze-crystallizer, the chloride ion content in the pre-cooled concentrated water is increased by increasing the degree of evaporation concentration, replacing the step of adjusting the pH value with sulfuric acid with the step of using hydrochloric acid and returning the sodium chloride mother liquor to the freeze-crystallization unit, that is, the salt-nitrate ratio of the actual feed is increased to be close to the designed salt-nitrate ratio, so that the operating temperature of the freeze-crystallizer is reduced to -5.5°C to -4.5°C, and at the same time, the freezing and blocking of the freeze-crystallizer is avoided when the temperature of the pre-cooled concentrated water is reduced to 0°C, thereby ultimately improving the extraction rate of Glauber's salt (or sodium sulfate).

[0034] 6. In the freeze crystallization of the present invention, the heat exchange temperature difference between the pre-cooled concentrated water and the heat exchange medium (i.e., ethylene glycol solution) is controlled to be ≤3°C, and the temperature difference of the heat exchange medium before and after the heat exchange is controlled to be ≤1.5°C, so that the low-temperature crystallization process can be carried out continuously and stably, and ultimately the extraction rate of sodium sulfate is greatly improved, thereby improving the quality of the sodium chloride product.

[0035] 7. In the melt crystallization of the present invention, the sodium sulfate produced by the frozen crystallization is melt crystallized to obtain a sodium sulfate product that complies with GB / T 6009-2014.

[0036] 8. In the sodium chloride crystallization of the present invention, the frozen mother liquor is crystallized by removing volatiles, circulating flash crystallization and recrystallization, so that a first-grade sodium chloride product that meets the requirements of refined industrial salt can be produced.

[0037] 9. In the crystallization of the impurity salt of the present invention, the sodium chloride mother liquor is crystallized by circulating flash crystallization and drum drying, which can achieve zero discharge of wastewater. The generated impurity salt product can be dissolved and returned to the sodium sulfate crystallization unit, thereby further improving the recovery rate of sodium chloride and sodium sulfate crystal salts and reducing the amount of impurity salt generated by the device. BRIEF DESCRIPTION OF THE DRAWINGS

[0038] Figure 1 Schematic diagram of the high-salt wastewater desalination crystallization process according to an embodiment of the present invention. DETAILED DESCRIPTION

[0039] The present invention is described in detail below in conjunction with specific embodiments, but such description is not intended to limit the scope of protection of the present invention. Those skilled in the art should know that certain process steps and equipment used in the specific embodiments of the present invention can be properly adjusted or replaced without affecting the implementation of the present invention.

[0040] The first aspect of the present invention provides a process for salt crystallization of high-salt wastewater, the process comprising the following steps:

[0041] S1: removing hardness and silicon from high-salinity wastewater to obtain hardness and silicon-free effluent;

[0042] S2: performing turbidity removal treatment on the hardness and silicon removal effluent to obtain turbidity-removed effluent;

[0043] S3: subjecting the deturbidified effluent to ion exchange softening treatment to obtain softened effluent;

[0044] S4: performing membrane treatment on the softened effluent to obtain membrane treated effluent;

[0045] S5: subjecting the membrane treated effluent to ozone oxidation treatment to obtain ozone oxidized effluent;

[0046] S6: performing evaporation and concentration treatment on the ozone oxidation effluent to obtain evaporation and concentration effluent;

[0047] S7: subjecting the evaporated concentrated water to freezing crystallization treatment to obtain Glauber's salt and frozen mother liquor, wherein the frozen mother liquor is divided into two parts, 85% to 90% of which is sent to step S9 for sodium chloride crystallization treatment, and the remaining 10% to 15% is repeatedly subjected to freezing crystallization treatment;

[0048] S8; subjecting the mirabilite to melt crystallization to obtain a sodium sulfate product and a molten mother liquor, and repeatedly subjecting the molten mother liquor to melt crystallization;

[0049] S9: subjecting the frozen mother liquor obtained in step S7 to sodium chloride crystallization to obtain a sodium chloride product and a sodium chloride mother liquor, wherein the sodium chloride mother liquor is divided into two parts, 5% to 9% of which is sent to step S10 for impurity salt crystallization, and the remaining 91% to 95% is repeatedly subjected to sodium chloride crystallization;

[0050] S10: subjecting the sodium chloride mother liquor obtained in step S9 to a mixed salt crystallization treatment to obtain a mixed salt product.

[0051] In this scheme, the high-salt wastewater refers to wastewater with a total salt content of 1wt% (equivalent to 10g / L). Its main components are sulfate and sodium salt in inorganic salts. These wastewaters mainly come from petrochemical, coal chemical plants and natural gas collection and processing. These high-salt wastewaters are usually weakly alkaline and have the characteristics of high salt content, hardness, silicon and organic matter.

[0052] Furthermore, the high-salinity wastewater generally contains Na + , K + , Ca 2+ Mg 2+ , Fe 3+ , Mn 2+ 、Al 3+ NH4 + , Cu 2+ , Ba 2 + , Sr 2+ 、As 3+ Isocations and Cl - 、SO4 2- 、HCO3 - 、CO3 2- OH - 、NO3 - 、NO2 - PO4 3- 、F - , S 2- 、SiO3 2- Anions.

[0053] In step S1, the hardness and silicon removal process includes:

[0054] (S1a) adding a calcium agent to the high-salinity wastewater to obtain a first clarified water outlet;

[0055] (S1b) adding a magnesium agent to the first clarified water to obtain a second clarified water;

[0056] (S1c) adding a coagulant, sodium carbonate and a coagulant aid to the second clarified water to obtain a hardness- and silicon-free water.

[0057] Preferably, the high-salt wastewater can be subjected to a homogenization operation well known in the art before the hardness and silicon removal treatment. The homogenization operation has two main functions: one is to regulate the water volume and store the wastewater to ensure stable water inflow after the device is operated; the other is to regulate the water quality, adjust various water qualities and stabilize the inflow water quality.

[0058] It is well known to those skilled in the art that the hardness in the hardness and silicon removal refers to the sum of the concentrations of various metal ions (such as calcium and magnesium ions) in water, and the silicon in the hardness and silicon removal refers to silicon in the form of silicate, colloid or particles.

[0059] In some embodiments, in step (S1a), the calcium agent is calcium chloride and / or calcium hydroxide, the purpose of which is to remove F from high-salinity wastewater. - 、CO3 2- PO4 3- Calcium and other insoluble anions.

[0060] Furthermore, sodium hydroxide can be added to the calcium agent to make the pH of the calcium agent reacting with the high-salinity wastewater to be 11-11.5, thereby removing most of the hardness in the high-salinity wastewater, such as Fe 3+ , Mn 2+ 、Al 3+ , Cu 2+ , Sr 2+ Plasma.

[0061] Furthermore, the amount of the calcium agent added is 0.8 to 1.5 g / L; further, the reaction time of the calcium agent and high-salt wastewater is 5 to 60 minutes, preferably 15 to 30 minutes.

[0062] In some embodiments, the hardness of the first clarified water is ≤65 mg / L.

[0063] In some embodiments, in step (S1b), the magnesium agent is at least one of magnesium oxide, magnesium chloride and magnesium sulfate, and the purpose is to remove silicate in high-salt wastewater.

[0064] Furthermore, the added amount of the magnesium agent is 0.8 to 2.0 g / L.

[0065] In some embodiments, sulfuric acid or hydrochloric acid may be added to the magnesium agent, so that the pH value when the magnesium agent reacts with the first clarified water is 7 to 8;

[0066] Furthermore, considering that the freezing crystallizer used in the following freezing crystallization treatment needs to meet the feed salt-nitrate ratio (the ratio of sodium chloride to sodium sulfate) of 3:7 to 4:6 in design, otherwise the freezing crystallizer is prone to freezing and blocking. Therefore, when adjusting the pH, sulfuric acid or hydrochloric acid can be selected according to the Cl- and SO4 2- to judge by the relative content.

[0067] In some embodiments, in step (S1b), the time for the first clarified water to react with the magnesium agent is 5 to 60 minutes, preferably 15 to 30 minutes.

[0068] Those skilled in the art will appreciate that the addition of a magnesium agent to the first clarified water can generate new magnesium hydroxide precipitates, which can create a high suspended matter environment, maintain a high adsorption specific surface area, adsorb silicon in precipitated and colloidal forms in the wastewater, and convert magnesium hydroxide into magnesium silicate precipitates with lower solubility, thereby achieving the purpose of silicon removal.

[0069] In some embodiments, the silicon content of the second clarified water is ≤40 mg / L.

[0070] In some embodiments, in step (S1c), the coagulant is selected from 30% polyferric sulfate solution, 30% polyferrous sulfate solution or 30% polyferric chloride solution, and the added amount is 0.1 to 0.25 g / L;

[0071] The coagulant aid is PAM (polyacrylamide), and the addition amount is 0.2-0.5 mg / L;

[0072] The amount of sodium carbonate added is 0.1-0.5 g / L.

[0073] Those skilled in the art can understand that the coagulant can destabilize the colloid in the second clarified water; the coagulant aid can enhance the adsorption effect of the flocs and remove the suspended matter in the second clarified water; the role of the sodium carbonate is to remove the calcium hardness in the second clarified water.

[0074] Furthermore, the reaction time of the second clarified water with the coagulant, sodium carbonate and coagulant aid is 5 to 60 minutes, preferably 15 to 30 minutes.

[0075] In some embodiments, steps (S1a) to (S1c) can be carried out in a reaction sedimentation tank and a V-type filtration tank. After steps (S1a) to (S1c), calcium hardness, magnesium hardness, silicates, fluorides and their complexes in the high-salt wastewater are removed, and the hardness of the effluent after hardness and silicon removal is ≤50 mg / L.

[0076] In step S2, the purpose of the turbidity removal treatment is to remove the insoluble suspended matter in the hard and silicon-removed effluent to obtain turbidity-removed effluent.

[0077] Furthermore, the turbidity removal treatment method is to use a high-density sedimentation tank for sedimentation, and the conditions are: the rising flow rate of the pipe area water flow is 0.1-0.2 m / s, and the sludge solid load is 11-12 kg / (m 2 h), the outer edge linear speed of the scraper is 0.03~0.04m / s,

[0078] In some embodiments, the high-density sedimentation tank is provided with an inclined plate to increase the hydraulic retention time and ultimately improve the turbidity removal efficiency.

[0079] Furthermore, the turbidity of the deturbidity effluent is ≤5NTU.

[0080] In step S3, the softening treatment is to use a hydrogen-type weakly acidic cation exchanger to deeply remove the hardness of the deturbidified water to remove cations other than Na+ and K+ to obtain ion exchanged water.

[0081] Furthermore, the hardness of the softened water is ≤5 mg / L.

[0082] In some embodiments, the hydrogen-type weakly acidic cation exchanger comprises a primary weakly acidic cation exchanger and a secondary weakly acidic cation exchanger.

[0083] Furthermore, the regeneration period of the primary weak acid cation exchanger is 24 hours, and the regeneration period of the secondary weak acid cation exchanger is 48 hours.

[0084] In some embodiments, the ion exchange effluent obtained in step S3 can be adjusted to an acidic pH using sulfuric acid or hydrochloric acid, and then decarbonized to reduce the residual CO2 in the ion exchange effluent to below 5 mg / L, and then step S4 is performed.

[0085] Furthermore, the step of adjusting the pH to acidity can prevent the salt water from causing alkaline corrosion to the evaporation equipment of the evaporation concentration processing unit.

[0086] In some embodiments, the decarbonization treatment uses a decarbonizer in the art, which removes free CO2 in water by means of air blast degassing. Water is introduced from the top of the decarbonizer, flows through the surface of the packing layer through a spray device, and air enters from the lower air outlet and passes through the packing layer in reverse. Free CO2 in the water is rapidly decomposed into the air and discharged from the top of the decarbonizer.

[0087] Furthermore, the decarbonizer is arranged after the hydrogen-type weakly acidic cation exchanger, and the ion exchange effluent is degassed by the decarbonizer and then flows into the water production pool of the softening unit, at which time the residual CO2 in the water is reduced to ≤5 mg / L.

[0088] In some embodiments, the characteristics of the softened water after the decarbonization treatment are as follows: pH = 6-9, hardness ≤ 5 mg / L, SiO2 content ≤ 40 mg / L, alkalinity ≤ 80 mg / L.

[0089] In step S4, the membrane treatment includes:

[0090] (S4a) subjecting the softened water to ultrafiltration to obtain ultrafiltration water;

[0091] (S4b) The ultrafiltration effluent is subjected to reverse osmosis treatment to obtain membrane treated effluent.

[0092] In some embodiments, in step (S4a), the ultrafiltration treatment may use an ultrafiltration device or system well known in the art, such as a rolled membrane assembly made of PVDF.

[0093] Furthermore, the conditions of the ultrafiltration treatment are: water inlet pressure ≤ 0.3 MPa, transmembrane pressure difference ≤ 0.16 MPa, and recovery rate ≥ 95%.

[0094] In some embodiments, the ultrafiltration treatment can remove precipitated colloids and suspended solids in the softened effluent, reduce turbidity and sludge density, and thus improve the efficiency of the subsequent reverse osmosis treatment.

[0095] Furthermore, the soluble solid content (ie, salt content) of the ultrafiltration water is 10-14 g / L, the turbidity of the ultrafiltration water is ≤0.2 NTU, and the pollution index (SDI) of the ultrafiltration water is ≤3.

[0096] In some embodiments, in step (S4b), the reverse osmosis treatment may use a reverse osmosis device or system well known in the art, for example, a reverse osmosis spiral membrane module of DuPont, Toray or Hydraulic Energy may be used.

[0097] Furthermore, the reverse osmosis treatment conditions are: water inlet pressure ≤ 3.0 MPa, interstage pressure ≤ 0.3 MPa, and recovery rate ≥ 75%.

[0098] In some embodiments, the reverse osmosis treatment can concentrate and reduce the ultrafiltration water. After the ultrafiltration water is treated by the reverse osmosis membrane component, it is separated into fresh water (passed through the reverse osmosis membrane) and concentrated water (not passed through the reverse osmosis membrane, i.e., reverse osmosis effluent).

[0099] Furthermore, the dissolved solid content (i.e., salt content) of the fresh water is 0.2 g / L, accounting for 75% of the total water intake of the reverse osmosis device or system, and can be reused in the production device as high-quality regenerated water (for example, for dissolving Glauber's salt, or recrystallizing sodium chloride, or dissolving and reusing miscellaneous salts);

[0100] Furthermore, the concentrated water (ie, reverse osmosis effluent) has a dissolved solid content (ie, salt content) of 41-45 g / L, accounting for 25% of the total water inlet into the reverse osmosis device or system, and enters step S5 as the final membrane treated effluent.

[0101] Furthermore, the conductivity of the water treated by the membrane is ≤200us / cm 2 .

[0102] In some embodiments, in order to reduce the scale of subsequent crystallization units, the reverse osmosis recovery rate (fresh water output / total water intake) of the reverse osmosis treatment needs to be controlled to be ≥75%, and the system desalination rate (salinity of fresh water / salinity of total water intake) needs to be controlled to be ≥98%.

[0103] In some embodiments, the reverse osmosis treatment may add different agents (such as reducing agents, scale inhibitors, non-oxidizing bactericides, hydrochloric acid, sodium hydroxide, etc.) to the ultrafiltration effluent according to changes in the operating state to reduce the cleaning frequency of the reverse osmosis device and extend the life of the reverse osmosis membrane.

[0104] In step S5, the ozone oxidation treatment uses ozone catalytic oxidation technology to remove organic matter in the membrane treated water, and at the same time decomposes the macrocyclic molecules or long-chain molecules therein into easily decomposable small molecules, thereby reducing (or even removing) the chromaticity of the softened water to obtain ozone oxidized water.

[0105] Furthermore, the catalyst used in the ozone catalytic oxidation technology may be an aluminum-based ozone catalyst, a silicon-aluminum ozone catalyst or a silicon-aluminum-titanium ozone catalyst.

[0106] In some embodiments, the ozone oxidation treatment is carried out under the following conditions: the water inlet is 90 to 114 m 3 / h, the total ozone production is 60kg / h, the ozone concentration is 148mg / L (10wt%), and the reaction time is 20-50min.

[0107] Furthermore, the removal rate of the total amount of organic matter (COD) treated by ozone oxidation is above 40%, and the COD of the effluent from ozone oxidation is ≤200 mg / L.

[0108] In step S6, the evaporation concentration process comprises the following steps:

[0109] (S6a) adjusting the pH of the ozone oxidation effluent to 5.5-6.0;

[0110] (S6b) heating the solution obtained in (S6a) to a temperature close to the boiling point;

[0111] (S6c) deoxygenating the solution obtained in (S6b) to obtain deoxygenated effluent;

[0112] (S6d) Evaporating and concentrating the deoxygenated water to obtain evaporated and concentrated water.

[0113] In some embodiments, in step (S6a), the purpose of pH adjustment is to prevent scaling of the evaporator during evaporation concentration treatment.

[0114] Furthermore, the pH adjustment can be carried out using sulfuric acid or hydrochloric acid, and the ozone oxidation effluent can be adjusted according to the Cl - and SO4 2- The relative content is used to judge, so that the feed salt-nitrogen ratio during the final freezing crystallization is close to its designed value.

[0115] In some embodiments, the step (S6b) is a preheating step before the evaporation and concentration, which can reduce the solubility of oxygen and other non-condensable gases on the one hand, and reduce the steam load of the deaerator during the deoxygenation treatment on the other hand, thereby ensuring the operation effect of the deaerator.

[0116] In some embodiments, in step (S6c), the purpose of deoxygenation is to remove non-condensable gases in water, and also to reduce the corrosion of Cl- on the sodium chloride crystallization unit during subsequent sodium chloride crystallization.

[0117] Furthermore, the deoxygenation treatment is carried out in a deaerator, specifically, the solution obtained in (S6b) is sprayed through a nozzle and flows downward step by step in the deaerator, while contacting with steam flowing upward in countercurrent to obtain the deoxygenated effluent.

[0118] Furthermore, the deoxygenation treatment is carried out in a deaerator, and the solution obtained (S6b) is sprayed through a nozzle and flows downward step by step in the deaerator, while contacting with steam flowing upstream in the countercurrent, wherein the removed gas is discharged into the atmosphere together with a small amount of vent steam, and finally deoxygenated effluent is obtained.

[0119] In some embodiments, in step (S6d), when performing the evaporation concentration, the deoxygenated water is divided into two streams, which are evaporated and concentrated in an MVR evaporation device and a TVR evaporation device in the field respectively, and finally the obtained concentrated liquids are combined into one stream, which is the evaporation and concentration water of step (S6d).

[0120] Further, the operating temperature of the TVR evaporation device is 98°C to 99.5°C, and the pressure is ≤50kPa; the operating temperature of the MVR evaporation device is 98°C to 99.5°C, and the pressure is ≤50kPa;

[0121] Furthermore, the evaporation concentration method can concentrate the ozone oxidation effluent by more than 5 times, that is, the amount of evaporation concentration effluent is only 20% of the corresponding water intake, and reusable steam condensate can be produced, reducing the scale of subsequent sodium sulfate crystallization units and sodium chloride crystallization units.

[0122] Furthermore, the soluble solid content in the evaporation and concentration water is 200-230 g / L.

[0123] In step S7, the freezing crystallization comprises:

[0124] (S7a) mixing the evaporation concentrated water with the frozen mother liquor repeatedly subjected to freeze crystallization treatment, and precooling the mixture to obtain precooled concentrated water having a temperature of 28 to 40° C.;

[0125] (S7b) freezing and crystallizing the pre-cooled concentrated water to obtain a frozen slurry with a temperature of -5.5 to -4.5° C.;

[0126] (S7c) performing solid-liquid separation on the frozen slurry to obtain sodium sulfate and frozen mother liquor, wherein the frozen mother liquor is divided into two parts, 85% to 90% of which is sent to step S9 for sodium chloride crystallization treatment, and the remaining 10% to 15% is repeatedly subjected to the frozen crystallization treatment;

[0127] In some embodiments, in step (S7a), before the pre-cooling treatment, the mixed solution obtained after mixing may be concentrated to improve the efficiency of freeze crystallization.

[0128] In some embodiments, in step (S7a), during the precooling treatment, the temperature of the mixed solution obtained after mixing is slowly reduced to 28-40° C. to avoid rapid crystallization caused by rapid cooling and thus clogging the heat exchange equipment.

[0129] Furthermore, if the temperature of the pre-cooled concentrated water is too low, the freezing crystallizer will be blocked, and if the temperature is too high, the effect of freezing crystallization is not very ideal. The temperature of the pre-cooled concentrated water is most preferably 28±0.2°C.

[0130] In some embodiments, in step (S7b), the freeze-crystallization is carried out in a freeze-crystallizer in the art, such as a slurry internal circulation crystallizer. Since an internal guide tube is provided in the crystallizer, a circulation channel is formed, which is conducive to continuous crystallization.

[0131] Furthermore, the freezing crystallizer is preferably a DTB type crystallizer, in which a draft tube is provided, and the circulating slurry in the freezing crystallizer enters the freezing crystallizer from the upper part of the draft tube, and the circulating slurry contacts the crystals in the freezing crystallizer, and the crystals gradually grow. Large-grained crystallized salt will settle in the sedimentation area outside the annular baffle of the draft tube, while small-grained salt will enter the circulation pipe with the circulating slurry to continue crystallization, and preferably the soluble solids in the freezing crystallizer are controlled within 15%.

[0132] Furthermore, the heat exchange medium of the freezing crystallizer is ethylene glycol solution (also called chilled water in technology).

[0133] In some embodiments, in step (S7b), when the pre-cooled concentrated water is subjected to freeze crystallization, if the heat exchange temperature difference is too large (more than 3°C), rapid nucleation of crystals will occur, i.e., "crystal explosion", which results in a short operating cycle of the freezing crystallizer and requires frequent thawing. Therefore, it is necessary to control the temperature difference between the pre-cooled concentrated water and the ethylene glycol solution during heat exchange to ≤3°C.

[0134] Furthermore, if the temperature difference between the precooled concentrated water and the ethylene glycol solution is too large, the precooled concentrated water will be cooled rapidly, and the sodium sulfate crystals in the precooled concentrated water will precipitate quickly, resulting in a "crystal explosion" phenomenon, which will eventually block the cooler of the freezing crystallizer. On the other hand, if the temperature difference between the precooled concentrated water and the ethylene glycol solution is too small, a larger amount of ethylene glycol solution will be required for crystallization, which will increase the operating load of the refrigeration unit.

[0135] The inventors have found that by adding an internal circulation pump on one side of the ethylene glycol solution, the ethylene glycol solution after heat exchange is returned to the inlet of the freezing crystallizer, mixed with the ethylene glycol solution from the refrigeration unit, and then heat exchanged with the pre-cooled concentrated water, this can avoid excessive temperature difference between the pre-cooled concentrated water and the ethylene glycol solution.

[0136] Furthermore, the circulation volume of the internal circulation pump is 850-950m 3 / h.

[0137] On the one hand, the above measures appropriately increase the flow rate of the ethylene glycol solution entering the freezing crystallizer and reduce the temperature difference between the ethylene glycol solution and the pre-cooled concentrated water during heat exchange. On the other hand, the ethylene glycol solution is circulated to ensure the heat exchange effect with the pre-cooled concentrated water. These two aspects can ensure that the crystallization process of sodium sulfate is carried out in the metastable zone, provide good conditions for its crystal growth, realize a continuous and stable low-temperature crystallization process, and are not prone to the "crystal explosion" phenomenon.

[0138] In some embodiments, the temperature difference before and after the precooling concentrated water heat exchange is ≤1.5°C, preferably the temperature difference before and after the precooling concentrated water heat exchange is ≤0.5°C;

[0139] In some embodiments, in step (S7b), the present invention utilizes the change in solubility of sodium sulfate and sodium chloride at different temperatures in the solution, and cools down the solution by heat exchange with brine and ethylene glycol solution, so that sodium sulfate precipitates from the solution in the form of sodium sulfate (sodium sulfate decahydrate) crystals. When the temperature is reduced to -5°C, the precipitation rate of sodium sulfate is above 98%.

[0140] In some embodiments, the designed freezing crystallizer requires a salt-nitrate ratio (the ratio of sodium chloride to sodium sulfate) of 3:7 to 4:6. The inventors have found that when the actual feed salt-nitrate ratio is low (especially lower than 2.8:7.2), ice and freezing are likely to occur when the temperature of the pre-cooled concentrated water is reduced to below -2°C during freezing crystallization. However, at this time, the temperature for the precipitation of sodium sulfate crystals (-5°C) has not yet been reached, and effective separation of Glauber's salt and frozen mother liquor cannot be achieved.

[0141] That is to say, when the actual feed salt-nitrate ratio is too low, the operating temperature of the freezing crystallizer will not be able to be lowered to the design temperature (-5.0℃~-4.5℃), which will ultimately lead to a decrease in the extraction rate of Glauber's salt and an inability to produce qualified sodium chloride products.

[0142] In some embodiments, through the scheme of the present invention, the designed salt-nitrogen ratio of the freezing crystallizer that can perform freezing crystallization at -5.0°C to -4.5°C can be reduced from the original (3:7) to (4:6) to a minimum of 2.8:7.2.

[0143] The inventors have taken the following measures: first, the degree of evaporation concentration in step S6 is increased, and the TDS salt content in the concentrated water is controlled at 200-230 g / L (significantly higher than 140-150 g / L in the prior art); second, the water quality of the feed to the freezing crystallizer is sampled and analyzed. According to the degree to which the water quality deviates from the designed salt-nitrate ratio (2.8:7.2), the step of adjusting the pH value with sulfuric acid can be replaced by using hydrochloric acid to increase the Cl- ion content in the corresponding brine, or in step S9, part of the sodium chloride mother liquor is refluxed to the freezing crystallization unit to further increase the chloride ion content in the pre-cooled concentrated water.

[0144] Therefore, the above measures make the salt-nitrate ratio of the actual feed of the freezing crystallizer within the designed salt-nitrate ratio, and reduce the operating temperature of the freezing crystallizer to -5.5℃~-4.5℃. At the same time, it avoids the freezing blockage of the cooler of the freezing crystallizer after the brine temperature drops to 0℃, and improves the extraction rate of Glauber's salt.

[0145] In some embodiments, in step (S7c), the solid-liquid separation is preferably centrifugation, and the SO4 2- Ion content ≤5500mg / L.

[0146] In some embodiments, in step (S7c), the frozen slurry may be thickened using a cyclone separator before centrifugation to increase the concentration of the mirabilite and stabilize the operation of the centrifuge.

[0147] Furthermore, the supernatant obtained after thickening is returned to the freezing crystallizer, and the bottom thickened liquid obtained enters the centrifuge to achieve solid-liquid separation.

[0148] In some embodiments, in step (S7c), if the frozen mother liquor obtained by centrifugation contains a small amount of suspended solids (i.e., concentrated solution of Glauber's salt crystals), a sedimentation device can be provided to separate the suspended solids from the frozen mother liquor, and then the suspended solids are sent to step S7 to repeat the frozen crystallization treatment.

[0149] The above two methods can significantly reduce the content of sodium sulfate in the frozen mother liquor, thereby ensuring the recovery rate of the sodium chloride crystallization system and the quality of the sodium chloride product.

[0150] Furthermore, the frozen mother liquor is divided into two parts, 85%-90% of which is sent to step S9 for sodium chloride crystallization treatment, and the remaining 10%-15% is repeatedly subjected to freezing crystallization treatment.

[0151] In some embodiments, in step (S7c), the frozen mother liquor subjected to the freeze-crystallization treatment is repeatedly subjected to the reflux requirement on the one hand to prevent crystal explosion, and on the other hand to balance the pressure during the freeze-crystallization treatment.

[0152] In the existing technology of using nanofiltration membrane for salt separation and crystallization, the selective permeability of the nanofiltration membrane is used to allow Cl- to pass through the membrane, while SO4 2- Remaining on the feed side, however SO4 2- After all, it is impossible for all of them to be intercepted by the nanofiltration membrane, and they will inevitably pass through the membrane and enter the sodium chloride side, which will eventually lead to a short service life of the nanofiltration membrane. At the same time, because the water contains organic matter, the nanofiltration membrane will be blocked when separating it, and its life is short. However, the different low-temperature critical frozen crystal solubility of the present invention can separate the salt and sodium chloride in the brine, and a high recovery rate of the salt can be obtained, and the amount of foreign salt generated can be reduced to the maximum extent. At the same time, it has a strong resistance to water quality fluctuations, and overcomes the attenuation and instability of the salt separation efficiency of the nanofiltration membrane.

[0153] In step S8, the melt crystallization process includes:

[0154] (S8a) introducing water vapor into the sodium sulfate to dissolve the sodium sulfate into a supersaturated sodium sulfate solution;

[0155] (S8b) mixing the supersaturated sodium sulfate solution with a molten mother liquor subjected to repeated melt crystallization treatment, and obtaining a sodium sulfate concentrated salt slurry after flash evaporation;

[0156] (S8c) performing solid-liquid separation on the sodium sulfate concentrated salt slurry to obtain anhydrous sodium sulfate crystals and a molten mother liquor;

[0157] (S8d) drying the anhydrous sodium sulfate crystals to obtain a sodium sulfate product;

[0158] In some embodiments, in step (S8a), the supersaturated sodium sulfate solution contains 15 wt.% of undissolved sodium sulfate crystals.

[0159] In some embodiments, in step (S8b), before the flash evaporation, the mixed solution obtained after mixing is preferably heated to a temperature close to the boiling point.

[0160] Furthermore, the flash evaporation temperature is 96-110° C., and the flash evaporation pressure based on the atmosphere is 1.00-5.00 kPa.

[0161] In some embodiments, in step (S8c), the solid-liquid separation method is preferably centrifugation;

[0162] Further, the molten mother liquor is returned to step (S8b) to be mixed with the supersaturated sodium sulfate solution, and the melt crystallization treatment is repeated;

[0163] Furthermore, the moisture content of the anhydrous sodium sulfate crystals is ≤5%.

[0164] In some embodiments, in step (S8d), the sodium sulfate content of the sodium sulfate product is ≥99.0 g / 100 g, the water-insoluble matter content is ≤0.05 g / 100 g, the chloride content is ≤0.35 g / 100 g, the TOC content is ≤50 mg / kg, the calcium and magnesium ion content is ≤0.3 g / 100 g, and the moisture content is ≤0.2%, that is, the various indicators of the sodium sulfate product meet the requirements of Class II first-class products in "GBT 6009-2014 Industrial Anhydrous Sodium Sulfate".

[0165] In step S9, the sodium chloride crystallization process includes:

[0166] (S9a) mixing the frozen mother liquor obtained in step S7 with the sodium chloride mother liquor subjected to repeated sodium chloride crystallization treatment, and obtaining a sodium chloride concentrated salt slurry after flash evaporation;

[0167] (S9b) subjecting the sodium chloride concentrated salt slurry to solid-liquid separation to obtain crude sodium chloride wet salt and sodium chloride mother liquor;

[0168] (S9c) recrystallizing and drying the crude sodium chloride wet salt to obtain a sodium chloride product.

[0169] In some embodiments, the temperature of the frozen mother liquor obtained from step S7 is raised to 75°C±0.5°C before the sodium chloride crystallization treatment. The advantage of this is that the heat exchange load of the subsequent sodium chloride crystallizer during the sodium chloride crystallization treatment can be reduced, and the operating temperature of the sodium chloride crystallizer can be ensured to be stable at 110°C±1°C.

[0170] In some embodiments, in step (S9a), before the mixing, a volatile matter removal tower can be used to remove volatile small molecular organic matter in the feed, and carbon dioxide, oxygen and other non-condensable gases in the feed can also be removed. On the one hand, the effect of organic matter on the quality of sodium chloride crystal salt can be reduced, and on the other hand, the crystallization foaming caused by the enrichment of organic matter in the sodium chloride crystallizer can be reduced. At the same time, the corrosion of brine to equipment and pipelines can be reduced by deoxygenation.

[0171] Furthermore, the volatile matter is a small molecular organic matter obtained through ozone oxidation treatment, and the removal method can be to heat the frozen mother liquor with low-pressure steam and then separate the volatile organic matter by flash evaporation.

[0172] In some embodiments, in step (S9a), before the flash evaporation, the mixed liquid obtained by mixing can be heated to a temperature close to the boiling point.

[0173] In some embodiments, in step (S9a), the frozen mother liquor obtained in step S7 and the sodium chloride mother liquor repeatedly subjected to sodium chloride crystallization treatment are concentrated after flash evaporation, and the crystallized salt that is continuously precipitated due to supersaturation gradually falls into the salt legs from the cone of the sodium chloride crystallizer. The salt leg washing brine enters from the bottom, slowly rises along the leg wall to perform countercurrent washing on the salt crystals, and brings fine salt particles into the crystallizer to continue to participate in circulation and crystallization, while dissolving and removing soluble impurities, which is beneficial to improving the quality of the crystallized salt.

[0174] Furthermore, sight glass ports may be provided at different heights of the salt legs to facilitate observation of the settling of the crystallized salt particles.

[0175] In some embodiments, in step (S9a), the flash evaporation temperature is 95°C to 100°C, and the flash evaporation pressure based on atmospheric pressure is -30 to 5 kPa.

[0176] In some embodiments, in step (S9b), the sodium chloride concentrated salt slurry is preferably thickened using a cyclone separator before solid-liquid separation until the concentration of suspended matter in the sodium chloride concentrated salt slurry is increased to more than 40 wt.%;

[0177] Furthermore, the solid-liquid separation method is preferably centrifugation, the moisture content of the crude sodium chloride wet salt is 5wt.%, and the chloride ion content in the sodium chloride mother liquor is ≤120g / L.

[0178] In some embodiments, in step (S9b), the sodium chloride mother liquor is sent to step S10 for impure salt crystallization treatment, and / or is sent to step (S9a) for mixing with frozen mother liquor, and / or is sent to step S7 for frozen crystallization treatment.

[0179] Preferably, 5% to 9% of the sodium chloride mother liquor is sent to step S10 for impurity salt crystallization treatment, and the remaining 91% to 95% is sent to step (S9a) to be mixed with the frozen mother liquor and then repeatedly subjected to sodium chloride crystallization treatment.

[0180] In some embodiments, in step (S9c), the recrystallization mother liquor obtained by the recrystallization may also be sent to step (S7a) for mixing.

[0181] In step (S9b), the sodium chloride mother liquor to step S7 can increase the chloride ion content of the pre-cooled concentrated water, so that the salt-nitrate ratio of the feed to the freezing crystallizer is between 3:7 and 4:6, thereby reducing the freezing blockage of the freezing crystallizer and improving the extraction rate of sodium sulfate during the freezing crystallization process.

[0182] In some embodiments, in step (S9c), the sodium chloride content of the sodium chloride product is ≥98.5 g / 100 g, the water-insoluble matter content is ≤0.1 g / 100 g, the sulfate ion content is ≤0.5 g / 100 g, the moisture content is ≤0.5%, and the calcium and magnesium ion content is ≤0.4 g / 100 g, that is, the obtained sodium chloride product meets the requirements of the first grade refined industrial salt in the "GB / T 5462-2015 Industrial Salt" standard.

[0183] In step S10, the impure salt crystallization process includes:

[0184] (S10a) flashing the sodium chloride mother liquor obtained in step S9 to obtain a concentrated salt slurry of impure salts;

[0185] (S10b) performing solid-liquid separation on the impure salt concentrated slurry to obtain impure salt I and impure salt mother liquor;

[0186] (S10c) Drum drying the impure salt mother liquor to obtain impure salt II.

[0187] (S10d) mixing the impure salt I and the impure salt II and drying them to obtain an impure salt product.

[0188] In some embodiments, in step (S10a), the sodium chloride mother liquor obtained from step S9 can be heated to a temperature close to the boiling point before flash evaporation to reduce the heat exchange load of the subsequent impurity salt crystallization unit during the impurity salt crystallization treatment, and also to ensure the stability of the operating temperature of the sodium chloride crystallizer.

[0189] Furthermore, the flash evaporation temperature is 90° C. to 105° C., and the flash evaporation pressure based on the atmosphere is -5 to 5 kPa.

[0190] In some embodiments, in step (S10b), the solid-liquid separation method is preferably centrifugation, and the water content of the impurity salt I is ≤80%.

[0191] In some embodiments, in step (S10c), the drum drying is preferably double-drum drying, and the surface temperature of the drum is 115-120°C;

[0192] Furthermore, the drum drying can crystallize the miscellaneous salts with high solubility such as nitrates, which greatly reduces the energy consumption of the system, reduces the land occupation, and reduces the discharge of wastewater.

[0193] In some embodiments, in step (S10c), it is preferred to mix impure salt I and impure salt II to obtain 48% to 52% of the mixture for dissolution, and then return to step S7 for freeze crystallization, and the remaining 52% to 48% becomes the impure salt product, which can further improve the recovery rate of sodium sulfate and sodium chloride crystalline salts and reduce the production of impure salts.

[0194] Furthermore, the production volume of the impure salt product accounts for ≤12% of the production volume of the total crystalline salt (including sodium sulfate products and sodium chloride products), that is, the impure salt ratio is ≤12%.

[0195] The second aspect of the present invention provides a system for the above-mentioned high-salt wastewater salt crystallization process, and the system includes the following units according to the material flow:

[0196] The hardness and silicon removal unit is used to remove most of the hardness and silicon in the high-salinity wastewater to obtain hardness and silicon-free effluent;

[0197] Turbidity removal unit, used to remove suspended matter in the effluent from the hardness removal and silicon removal process to obtain turbidity removal effluent;

[0198] A softening unit is used to remove the remaining hardness in the deturbidified effluent water to obtain softened effluent water;

[0199] The membrane treatment unit is used to remove the precipitated colloids, suspended solids and reduced matter in the softened effluent to obtain membrane treated effluent;

[0200] Ozone oxidation unit, used to remove organic matter in membrane treated water to obtain ozone oxidized water;

[0201] An evaporation concentration unit is used to concentrate the ozone oxidation effluent to obtain evaporation concentrated effluent;

[0202] A freezing crystallization unit is used to freeze and crystallize the evaporation and concentration water to obtain sodium sulfate and frozen mother liquor;

[0203] A melt crystallization unit is used to melt crystallize the mirabilite to obtain a sodium sulfate product and a molten mother liquor;

[0204] A sodium chloride crystallization unit is used to crystallize sodium chloride from the frozen mother liquor to obtain a sodium chloride product and a sodium chloride mother liquor;

[0205] The impure salt crystallization unit is used to process part of the mother liquor discharged from the sodium chloride crystallization unit to obtain impure salt products; wherein

[0206] The hardness and silicon removal unit uses a reaction sedimentation tank and a V-type filter tank to treat the high-salt wastewater;

[0207] The turbidity removal unit uses a high-density sedimentation tank to treat the hardness and silicon removal effluent;

[0208] The softening unit uses a hydrogen-type weakly acidic cation exchanger to treat the deturbidified effluent;

[0209] The membrane treatment unit uses an ultrafiltration device and a reverse osmosis device to treat the softened effluent;

[0210] The ozone oxidation unit uses an ozone generator to treat the membrane treated water;

[0211] The evaporation concentration unit uses a TVR evaporation device and an MVR evaporation device to treat the ozone oxidation effluent;

[0212] The freezing crystallization unit uses a crystal slurry internal circulation crystallizer to treat the evaporation and concentration water;

[0213] The melt crystallization unit uses a forced circulation FC crystallizer to process the mirabilite;

[0214] The sodium chloride crystallization unit uses a negative pressure vacuum forced circulation FC crystallizer to treat the frozen mother liquor;

[0215] The mixed salt crystallization unit uses a forced circulation FC crystallizer and a drum dryer to treat the sodium chloride mother liquor.

[0216] In order to make the purpose, technical scheme and effect of the present invention clearer and more specific, the present invention will be described in detail below through specific embodiments and comparative examples.

[0217] In the following examples and comparative examples, the process flow diagram of the treatment is as follows Figure 1 As shown,

[0218] The water quality characteristics of the coal chemical high-salt wastewater are: the concentration of total dissolved solids (TDS) is 10625 mg / L, Cl - The concentration is 1910mg / L, SO4 2- The concentration is 3859 mg / L, Mg 2+ The concentration is 27 mg / L, Ca 2+ The concentration of F is 185.7 mg / L, the concentration of dissolved silicon is 154 mg / L, - The concentration is 54mg / L, HCO3 - The concentration of NH4O2 is 537mg / L, and the concentration of COD is 170mg / L.

[0219] The following sodium sulfate extraction rate and sodium chloride extraction rate are commonly used terms in this field and can be calculated through sampling and analysis by the quality inspection and analysis department; among them,

[0220] Sodium sulfate extraction rate: SO4 in water 2- The amount of salt converted entirely to sodium sulfate product;

[0221] Sodium chloride extraction rate: Cl in water - Salt content of the product converted into sodium chloride;

[0222] Trash salt rate: Trash salt output / (sodium chloride output + sodium sulfate output + trash salt output);

[0223] Example 1

[0224] Step S1: Coal chemical high-salt wastewater enters the hardness and silicon removal unit for treatment:

[0225] (S1a) The coal chemical high-salt wastewater (flow rate is 420m 3 / h) into the primary reaction tank, add calcium hydroxide (flow rate is 1.0g / L), and after reacting for 25min, the first clarified water is overflowed into the secondary reaction tank; wherein,

[0226] Under this condition, the hardness of the first clarified water obtained is 52 mg / L.

[0227] (S1b) adding magnesium oxide (flow rate of 0.9 g / L) to the secondary reaction tank and adjusting the pH in the tank to 7.5 using hydrochloric acid (concentration of 32%), and after reacting for 30 minutes, overflowing the obtained second clarified water into the tertiary reaction tank; wherein,

[0228] The silicon content of the obtained second clarified water was 38 mg / L.

[0229] (S1c) Adding polyferric sulfate solution (concentration of 30%, flow rate of 0.15 g / L) as coagulant, PAM (flow rate of 0.0003 g / L) as coagulant aid and sodium carbonate (flow rate of 0.2 g / L) into the tertiary reaction tank, after reacting for 25 minutes, obtain the de-hardening and de-siliconizing effluent; wherein,

[0230] Under this condition, the hardness of the de-siliconized effluent is 45 mg / L.

[0231] Step S2: The effluent from the hardness and silicon removal enters the turbidity removal unit for treatment:

[0232] The de-hardening and de-siliconizing effluent is passed into a high-density sedimentation tank, and after reacting for 145 minutes, a de-turbidifying effluent is obtained; wherein,

[0233] The operating conditions of the high-density sedimentation tank are: the water flow rate in the inclined tube area is 0.167m / s, the hydraulic retention time is 60min, and the sludge solid load is 12kg / (m 2 ·h), the outer edge linear speed of the scraper is 0.038m / s.

[0234] Under this condition, the turbidity of the effluent obtained was 3.5 NTU.

[0235] Step S3: The turbidity-removed water enters the softening unit for treatment:

[0236] (S3a) The turbidity-removed effluent enters the primary weak cation exchanger (using hydrogen-type ion exchange resin) and the secondary weak cation exchanger (using hydrogen-type ion exchange resin) in sequence to obtain ion exchange effluent; wherein,

[0237] The theoretical exchange capacity of the primary and secondary weak cation exchangers is 1.2 eq / l.

[0238] (S3b) adding hydrochloric acid to the ion exchange effluent to adjust its pH to 5.8, and then blowing it through a decarbonizer to obtain softened effluent; wherein,

[0239] The concentration of the hydrochloric acid is 32%.

[0240] Under this condition, the hardness of the softened water is 0 mg / L and the alkalinity (OH - The concentration of ions) is 65 mg / L.

[0241] Step S4: The softened effluent enters the membrane treatment unit for treatment:

[0242] (S4a) The softened effluent enters an ultrafiltration device (a wound membrane module made of PVDF membrane material) for cross-flow filtration to obtain ultrafiltration effluent; wherein,

[0243] The operating conditions of the ultrafiltration device are: water inlet pressure of 0.2MPa, transmembrane pressure difference of 0.11MPa, and recovery rate of 97%;

[0244] The ultrafiltration effluent obtained had a dissolved solid content (ie, salt content) of 10.5 g / L, a turbidity of 0.1 NTU, and a pollution index (SDI) of 2.

[0245] (S4b) the ultrafiltration effluent enters a high-pressure reverse osmosis device for reverse osmosis treatment to obtain membrane-treated effluent; wherein,

[0246] The operating conditions of the high-pressure reverse osmosis device are: water inlet pressure is 2.2 MPa, interstage pressure is 0.24 MPa, and recovery rate is 75.5%.

[0247] Under this condition, the conductivity of the membrane treated water was 115us / cm 2 .

[0248] Step S5, the membrane treated water enters the ozone oxidation unit for treatment:

[0249] Low-pressure oxygen enters the ozone generator, and after high-voltage discharge, it becomes ozone and enters the ozone dosing tank. Aluminum-based catalyst is added to the ozone dosing tank, and then the membrane treated water enters the ozone dosing tank to obtain ozone effluent.

[0250] The operating conditions of the ozone dosing pool are: water inflow of 95m 3 / h, the total ozone production is 28kg / h, the ozone concentration is 148mg / L (10wt%), and the ozone reaction time is 45min.

[0251] Under this condition, the COD removal rate of the ozone oxidation unit is 43%, and the effluent COD of the ozone oxidation effluent is 187 mg / L.

[0252] Step S6, the ozone oxidized water enters the evaporation concentration unit for treatment:

[0253] (S6a) Use hydrochloric acid (concentration of 32%) to adjust the pH of the ozone oxidation effluent to 5.8.

[0254] (S6b) The solution obtained in (S6a) is heated to 90°C using a plate heat exchanger.

[0255] (S6c) The solution obtained in (S6b) is sent to a deaerator to remove non-condensable gases in the solution. In the deaerator, the solution is sprayed through nozzles and flows downward step by step, and contacts with steam flowing upstream in the countercurrent direction. The removed gas is discharged into the atmosphere together with a small amount of vent steam, and finally deoxygenated effluent is obtained.

[0256] (S6d) the deoxygenated water is evenly divided into two streams, one of which is sent to an MVR evaporation device for evaporation and concentration, and the other is sent to a TVR evaporation device for evaporation and concentration, and finally the obtained concentrated liquids are combined into one stream to obtain evaporation and concentration water with a weight only 1 / 5 of that before concentration, and at the same time obtain water vapor condensate; wherein,

[0257] The operating conditions of the TVR evaporation device are: temperature 98.7°C, pressure 38kPa;

[0258] The operating conditions of the MVR evaporation device are: temperature of 98.7°C and pressure of 38 kPa.

[0259] Under these conditions, the evaporation concentration water obtained has a dissolved solid content of 225g / L and SO4 2- The ion content is 129g / L, Cl - The ion content is 35.08g / L.

[0260] Step S7, the evaporated concentrated brine enters the freezing crystallization unit for treatment:

[0261] (S7a) The evaporated concentrated water is heat exchanged in a precooler to obtain precooled concentrated water at 28°C.

[0262] (S7b) pre-cooling concentrated water is transported to a DTB type crystallizer for heat exchange with a circulating ethylene glycol solution (i.e., a heat exchange medium) to obtain a sodium sulfate slurry with a temperature of -5.0°C; wherein,

[0263] During the heat exchange process, the internal circulation flow rate of the ethylene glycol solution side is 850m 3 / h, and the rate of heat exchange of precooled concentrated water is 0.27℃ / time.

[0264] (S7c) sodium sulfate slurry is thickened by a cyclone separator and then centrifuged to obtain mirabilite and frozen mother liquor; wherein,

[0265] The frozen mother liquor still contains some suspended solids, which are separated by setting up a settling tank. The concentrated liquid at the bottom of the settling tank is used as a circulating sodium sulfate slurry and sent to (S5b) for reheating (i.e., repeating the freezing crystallization operation). The supernatant of the settling tank can enter the following step S9 for sodium chloride crystallization.

[0266] Under this condition, the obtained frozen mother liquor contains SO4 2- The ion content is 4580 mg / L.

[0267] Step S8, the mirabilite enters the melting crystallization unit for processing:

[0268] (S8a) passing the water vapor condensate (flow rate of 4.8 t / h) obtained in step (S6d) into a sodium sulfate hot melt tank containing the sodium sulfate to obtain a supersaturated sodium sulfate solution; wherein,

[0269] The content of anhydrous sodium sulfate in the supersaturated sodium sulfate solution is 15 wt.%.

[0270] (S8b) heating the supersaturated sodium sulfate solution to 110° C. and flash evaporating the solution in a sodium sulfate crystallizer to obtain a sodium sulfate concentrated salt slurry; wherein:

[0271] The flash temperature was 98.5° C., and the flash pressure based on the atmosphere was 3.8 kPa.

[0272] (S8c) the sodium sulfate concentrated salt slurry is concentrated by a cyclone thickener and then centrifuged to obtain sodium sulfate crystals and sodium sulfate mother liquor; wherein,

[0273] The sodium sulfate mother liquor is returned to step (S8b) to repeat melt crystallization.

[0274] (S8d) The anhydrous sodium sulfate crystals are dried in a fluidized bed to obtain a sodium sulfate product (flow rate is 2.83 t / h).

[0275] Under this condition, the parameters of the obtained sodium sulfate product are: sodium sulfate content 99.1g / 100g, water-insoluble matter 0.02g / 100g, chloride 0.25g / 100g, TOC content 25mg / kg, moisture content 0.04%, calcium and magnesium content 0, which meets the requirements of Class II first-class products in "GBT 6009-2014 Industrial Anhydrous Sodium Sulfate".

[0276] The extraction rate of the sodium sulfate product is 93.19%.

[0277] Step S9, the frozen mother liquor enters the sodium chloride crystallization unit for treatment:

[0278] (S9a) heating the frozen mother liquor to 75°C and then passing it into a volatile matter removal tower for flash evaporation; wherein,

[0279] The flash temperature was 95° C., and the flash pressure based on the atmosphere was 10.5 kPa.

[0280] (S9b) heating the solution obtained in (S9a) to 97.5°C, flash evaporating the solution in a sodium chloride crystallizer to obtain a sodium chloride concentrated salt slurry; wherein:

[0281] The flash temperature was 95° C., and the flash pressure based on the atmosphere was −28 kPa.

[0282] (S9c) thickening the sodium chloride concentrated salt slurry through a cyclone separator to increase the suspended matter concentration of the sodium chloride concentrated salt slurry to 42%, and obtaining crude sodium chloride wet salt and sodium chloride mother liquor after centrifugation; wherein,

[0283] 5% of the sodium chloride mother liquor is sent to step S10 for impurity salt crystallization treatment, and 95% of it enters step (S9b) and repeats the sodium chloride crystallization treatment;

[0284] The moisture content of the crude sodium chloride wet salt is 5%.

[0285] (S9d) dissolving the crude sodium chloride wet salt with the water vapor condensate obtained in step (S6d), and obtaining refined sodium chloride wet salt after recrystallization in a sodium chloride recrystallizer;

[0286] (S9e) The refined sodium chloride wet salt is dried in a fluidized bed to obtain a sodium chloride product (flow rate is 1 t / h).

[0287] Under this condition, the parameters of the sodium chloride product are: sodium chloride content 99.8g / 100g, water insoluble matter 0.04g / 100g, sulfate ion 0.16g / 100g, moisture content 0.02%, calcium and magnesium ion content 0g / 100g, which meets the requirements of the first-grade refined industrial salt in the "GB / T5462-2015 Industrial Salt" standard;

[0288] The extraction rate of the sodium chloride product is 81.59%.

[0289] Step S10, the sodium chloride mother liquor enters the impurity salt crystallization unit for treatment:

[0290] (S10a) heating the sodium chloride mother liquor to 95° C. and then flashing it in a mixed salt crystallizer to obtain a mixed salt concentrated slurry;

[0291] (S10b) Centrifuging the impure salt concentrated slurry to obtain impure salt I and impure salt mother liquor; wherein,

[0292] The water content of the impurity salt I is 65%;

[0293] (S10c) drying the impure salt mother liquor in a double drum dryer to obtain impure salt II; wherein,

[0294] The surface temperature of the drum is 118°C;

[0295] (S10d) Mix the impure salt I and the impure salt II, dissolve 50% of them in the steam condensate obtained in step (S6d) and return to step S7, and the other 50% becomes impure salt brine.

[0296] (S11d) The mixed salt brine is concentrated and dehydrated to obtain a mixed salt product with a flow rate of 0.28 t / h.

[0297] In summary, the sodium sulfate extraction rate of the coal chemical high-salt wastewater is 93.19%, and the sodium chloride extraction rate is 81.59%; the sodium sulfate product flow rate is 2.83 t / h, the sodium chloride product flow rate is 1 t / h, and the miscellaneous salt product flow rate is 0.28 t / h, that is, the miscellaneous salt rate is 6.81%.

[0298] Comparative Example 1 (Replacing hydrochloric acid with sulfuric acid and reducing the circulation amount of ethylene glycol solution when adjusting pH)

[0299] The only difference from Example 1 is that:

[0300] In step (S1b), step (S3b) and step (S6a), 32% hydrochloric acid is replaced by 98% sulfuric acid.

[0301] In step (S6), the evaporation concentrates the water, SO4 2- The ion content increased from 129g / L to 142g / L, Cl - The ion content decreased from 35.08g / L to 28.46g / L.

[0302] In step (S7b), the temperature of the sodium sulfate slurry is increased from -5°C to -2.0°C, and the internal circulation flow rate is increased from 850m 3 / h reduced to 250m 3 / h, the heat exchange rate is increased from 0.27℃ / time to 0.5℃ / time;

[0303] Under this condition,

[0304] In the frozen mother liquor obtained, SO4 2- The ion content increased from 4580mg / L to 22500mg / L;

[0305] After the operating temperature of the crystallizer reaches -2.0°C, it cannot be further reduced to the design temperature (-5.0°C to -4.5°C), and when the temperature of the pre-cooled concentrated water drops to 0.5°C, a "crystal explosion" phenomenon occurs in the freeze-crystallization cooler, causing the cooler to be blocked.

[0306] In step (S8d), the flow rate of the sodium sulfate product is reduced from 2.83 t / h to 1.91 t / h;

[0307] The parameters of the sodium sulfate product are: sodium sulfate content 98.5g / 100g, water-insoluble matter 0.35g / 100g, chloride 0.45g / 100g, moisture content 0.6%, calcium and magnesium content 0g / 100g, which meets the requirements of qualified products in "GB / T 6009-2014 Industrial Anhydrous Sodium Sulfate";

[0308] The extraction rate of the sodium sulfate product is reduced from 93.19% to 70.56%;

[0309] In step (S9e), the flow rate of the sodium chloride product is reduced from 1 t / h to 0.99 t / h;

[0310] The parameters of the sodium chloride product are: sodium chloride content is 72g / 100g, water insoluble content is 3.5g / 100g, SO4 2- The ion content is 18.2g / 100g, the moisture content is 4.3%, and the calcium and magnesium ion content is 2.0g / 100g, which means that it does not meet the requirements of the first-grade refined industrial salt in the "GB / T 5462-2015 Industrial Salt" standard, and the sodium chloride product salt is unqualified;

[0311] The extraction rate of the sodium chloride product was reduced from 81.59% to 80.12%.

[0312] In step (S11d), the flow rate of the mixed salt product is increased from 0.28 t / h to 0.55 t / h.

[0313] In summary, the sodium sulfate extraction rate of the coal chemical high-salt wastewater is 70.56%, and the sodium chloride extraction rate is 80.12%; the sodium sulfate product flow rate is 1.91t / h, the sodium chloride product flow rate is 0.99t / h, and the miscellaneous salt product flow rate is 0.55t / h, that is, the miscellaneous salt rate is 15.94%.

[0314] In the process production process of Comparative Example 1, when sulfuric acid was used to adjust the pH and a lower circulation rate of ethylene glycol solution was used, the freezer experienced crystal explosion and low extraction rates of sodium sulfate and sodium chloride. Therefore, these two variables were tested in the following Comparative Examples 2 and 3, respectively.

[0315] Comparative Example 2 (Reducing the Circulation Amount of Ethylene Glycol Solution)

[0316] The only difference from Example 1 is that:

[0317] In step (S7b), the temperature of the sodium sulfate slurry is increased from -5°C to -2.0°C, and the internal circulation flow rate is increased from 850m 3 / h reduced to 250m 3 / h, the heat exchange rate is increased from 0.27℃ / time to 0.5℃ / time;

[0318] Under this condition,

[0319] In the frozen mother liquor, SO4 2- The ion content increased from 4580 mg / L to 24500 mg / L.

[0320] After the operating temperature of the crystallizer reaches -2.0°C, it cannot be further reduced to the design temperature (-5.0°C to -4.5°C), and when the temperature of the pre-cooled concentrated water drops to 0.5°C, a "crystal explosion" phenomenon occurs in the freeze-crystallization cooler, causing the cooler to be blocked.

[0321] In step (S8d), the flow rate of the sodium sulfate product is reduced from 2.83 t / h to 1.86 t / h.

[0322] The parameters of the sodium sulfate product are as follows: sodium sulfate content is 98.6g / 100g, water-insoluble matter content is 0.38g / 100g, chloride content is 0.47g / 100g, moisture content is 0.5%, and calcium and magnesium ion content is 0g / 100g, which meets the requirements of qualified products in "GB / T 6009-2014 Industrial Anhydrous Sodium Sulfate";

[0323] The extraction rate of the sodium sulfate product was reduced from 93.19% to 70.56%, and a large amount of sodium sulfate was lost to the subsequent sodium chloride system.

[0324] In step (S9e), the flow rate of the sodium chloride product is reduced from 1 t / h to 0.99 t / h,

[0325] The parameters of the sodium chloride product are: sodium chloride content is 71g / 100g, water insoluble content is 3.4g / 100g, SO4 2- The ion content is 20.1g / 100g, the moisture content is 3.5%, and the calcium and magnesium ion content is 2.0g / 100g, which does not meet the requirements of the first-grade refined industrial salt in the "GB / T 5462-2015 Industrial Salt" standard;

[0326] The extraction rate of the sodium chloride product was reduced from 81.59% to 80.12%.

[0327] In step (S11d), the flow rate of the mixed salt product is increased from 0.28 t / h to 1.25 t / h.

[0328] In summary, the sodium sulfate extraction rate of the coal chemical high-salt wastewater is 70.56%, and the sodium chloride extraction rate is 80.12%; the sodium sulfate product flow rate is 1.86 t / h, the sodium chloride product flow rate is 0.99 t / h, and the miscellaneous salt product flow rate is 1.25 t / h, that is, the miscellaneous salt rate is 30.49%.

[0329] Comparative Example 3 (when adjusting pH, hydrochloric acid is replaced by sulfuric acid)

[0330] The only difference from Example 1 is that:

[0331] In step (S1b), step (S3b) and step (S6a), the 32% hydrochloric acid is replaced with 98% sulfuric acid, and the others remain unchanged;

[0332] Under this condition,

[0333] In step S6, the SO4 in the water is evaporated and concentrated. 2- The ion content increased from 129g / L to 142g / L, Cl - The ion content decreased from 35.08 g / L to 27.46 g / L;

[0334] In step S7, the obtained frozen mother liquor contains SO4 2- The ion content increased from 4580mg / L to 24500mg / L;

[0335] The extraction rate of the sodium sulfate product was reduced from 93.19% to 76.12%.

[0336] In step (S8d), the flow rate of sodium sulfate product changes from 2.83 t / h to 1.99 t / h,

[0337] The parameters of the obtained sodium sulfate product are as follows: sodium sulfate content 98.0 g / 100 g, water insoluble matter 0.30 g / 100 g, chloride 0.46 g / 100 g, moisture content 0.3%, calcium and magnesium content 0 g / 100 g, which meets the requirements of qualified products in "GB / T 6009-2014 Industrial Anhydrous Sodium Sulfate";

[0338] In step (S9e), the flow rate of the sodium chloride product is reduced from 1 t / h to 0.95 t / h,

[0339] The parameters of the sodium chloride product obtained in step S9 are: sodium chloride content is 72.5g / 100g, water-insoluble matter content is 2.2g / 100g, SO4 2- The ion content is 19.2g / 100g, the moisture content is 5.1%, and the calcium and magnesium ion content is 1.0g / 100g, which does not meet the requirements of the first-grade refined industrial salt in the "GB / T 5462-2015 Industrial Salt" standard;

[0340] The extraction rate of the sodium chloride product was reduced from 81.59% to 76.12%.

[0341] In step (S11d), the flow rate of the impurity salt product is increased from 0.28 t / h to 1.17 t / h.

[0342] In summary, the sodium sulfate extraction rate of the coal chemical high-salt wastewater is 76.12%, and the sodium chloride extraction rate is 76.12%; the sodium sulfate product flow rate is 1.99 t / h, the sodium chloride product flow rate is 0.95 t / h, and the miscellaneous salt product flow rate is 1.17 t / h, that is, the miscellaneous salt rate is 28.47%.

[0343] Example 1 Comparative Example 1 Comparative Example 2 Comparative Example 3 Sodium sulfate extraction rate 93.19％ 70.56％ 70.56％ 76.12％ Sodium chloride extraction rate 81.59％ 80.12％ 80.12％ 76.12％ Sodium Chloride Product Grades First-class products Non-first-class products Non-first-class products Non-first-class products Miscellaneous salt (product) rate 6.81％ 15.94％ 30.49％ 28.47％ Frozen crystallizer operation cycle More than 300 days 30-40 days Less than 10-15 days 12-15 days

[0344] The present invention has described in detail the preferred embodiments of the present invention, but the embodiments are only for illustrative purposes and are not intended to limit the scope of protection of the present invention. Those skilled in the art may make various other simple modifications and combinations within the scope of the present invention, and these simple modifications and combinations should also be regarded as the contents disclosed by the present invention and belong to the scope of protection of the present invention.

Claims

1. A process for salt crystallization of high-salt wastewater, characterized in that: The process comprises the following steps: S1: removing hardness and silicon from high-salinity wastewater to obtain hardness and silicon-free effluent; S2: performing turbidity removal treatment on the hardness and silicon removal effluent to obtain turbidity-removed effluent; S3: subjecting the deturbidified effluent to ion exchange softening treatment to obtain softened effluent; S4: performing membrane treatment on the softened effluent to obtain membrane treated effluent; S5: subjecting the membrane treated effluent to ozone oxidation treatment to obtain ozone oxidized effluent; S6: performing evaporation and concentration treatment on the ozone oxidation effluent to obtain evaporation and concentration effluent; S7: subjecting the evaporated concentrated water to freezing crystallization treatment to obtain Glauber's salt and frozen mother liquor, wherein the frozen mother liquor is divided into two parts, 85% to 90% of which is sent to step S9 for sodium chloride crystallization treatment, and the remaining 10% to 15% is repeatedly subjected to freezing crystallization treatment; S8: subjecting the sodium sulfate to melt crystallization to obtain a sodium sulfate product and a molten mother liquor, and repeatedly subjecting the molten mother liquor to melt crystallization; S9: subjecting the frozen mother liquor obtained in step S7 to sodium chloride crystallization to obtain a sodium chloride product and a sodium chloride mother liquor, wherein the sodium chloride mother liquor is divided into two parts, 5% to 9% of which is sent to step S10 for impurity salt crystallization, and the remaining 91% to 95% is repeatedly subjected to sodium chloride crystallization; S10: subjecting the sodium chloride mother liquor obtained in step S9 to a mixed salt crystallization treatment to obtain a mixed salt product.

2. The process according to claim 1, characterized in that In step S1, the hardness and silicon removal process includes: (S1a) adding a calcium agent to the high-salinity wastewater to obtain a first clarified water outlet; (S1b) adding a magnesium agent to the first clarified water to obtain a second clarified water; (S1c) adding a coagulant, sodium carbonate and a coagulant aid to the second clarified water to obtain a hardness- and silicon-free effluent; Preferably, in step (S1a), the calcium agent is calcium chloride and / or calcium hydroxide, the amount of the calcium agent added is 0.8-1.5 g / L, the time for the calcium agent to react with the high-salt wastewater is 5-60 min, the pH when the calcium agent reacts with the high-salt wastewater is 11-11.5, and the hardness of the first clarified water is ≤65 mg / L; Preferably, in step (S1b), the magnesium agent is at least one of magnesium oxide, magnesium chloride and magnesium sulfate, the amount of the magnesium agent added is 0.8-2.0 g / L, the time for the magnesium agent to react with the first clarified water is 5-60 min, the pH of the magnesium agent when reacting with the first clarified water is 7-8, and the silicon content of the second clarified water is ≤40 mg / L; Preferably, in step (S1c), the coagulant is selected from 30% polyferric sulfate solution, 30% polyferrous sulfate solution or 30% polyferric chloride solution, the amount of the coagulant added is 0.1-0.25 g / L, the coagulant aid is polyacrylamide, the amount of the coagulant added is 0.2-0.5 mg / L, the amount of sodium carbonate added is 1.0-1.8 g / L, the reaction time of the second clarified water with the coagulant, sodium carbonate and coagulant aid is 5-60 min, and the hardness of the hardness and silicon removal water is ≤50 mg / L.

3. The process according to claim 1 or 2, characterized in that In step S2, the turbidity removal treatment is carried out by using a high-density sedimentation tank for sedimentation, and the conditions are: the rising flow rate of the pipe area water flow is 0.1-0.2 m / s, and the sludge solid load is 11-12 kg / (m 2 h), the outer edge linear speed of the scraper is 0.03-0.04 m / s; Preferably, the turbidity of the deturbidified water is ≤5NTU.

4. The process according to any one of claims 1 to 3, characterized in that In step S3, a hydrogen-type weakly acidic cation exchanger is used to soften the turbidity removal effluent, wherein the pH of the softened effluent is 6-9, the hardness is ≤5 mg / L, the SiO2 content is ≤40 mg / L, and the alkalinity is ≤80 mg / L; Preferably, the hydrogen-type weakly acidic cation exchanger comprises a primary weakly acidic cation exchanger having a regeneration period of 24 hours and a secondary weakly acidic cation exchanger having a regeneration period of 48 hours.

5. The process according to any one of claims 1 to 4, characterized in that In step S4, the membrane treatment includes: (S4a) subjecting the softened effluent to ultrafiltration to obtain ultrafiltration effluent; (S4b) subjecting the ultrafiltration effluent to reverse osmosis treatment to obtain membrane-treated effluent; In step (S4a), the conditions of the ultrafiltration treatment are: water inlet pressure ≤ 0.3 MPa, transmembrane pressure difference of ultrafiltration treatment ≤ 0.16 MPa, and recovery rate ≥ 95%; In step (S4a), the ultrafiltration water has a dissolved solid content of 10 to 14 g / L, a turbidity of ≤0.2 NTU, and a pollution index of ≤3; In step (S4b), the reverse osmosis treatment conditions are: water inlet pressure ≤ 3.0 MPa, interstage pressure ≤ 0.3 MPa, recovery rate ≥ 75%; In step (S4b), the dissolved solid content of the membrane treated water is 41-43 g / L, and the conductivity of the membrane treated water is ≤200 us / cm 2 .

6. The process according to any one of claims 1 to 5, characterized in that In step S5, The ozone oxidation treatment uses an aluminum-based ozone catalyst, a silicon-aluminum ozone catalyst or a silicon-aluminum-titanium ozone catalyst; The conditions for the ozone oxidation treatment are: the water inflow is 90-114 m 3 / h, the total ozone production is 60kg / h, the ozone concentration is 148mg / L, and the reaction time is 20-50min; The COD of the ozone oxidation effluent is ≤200 mg / L.

7. The process according to any one of claims 1 to 6, characterized in that In step S6, the evaporation concentration process includes: (S6a) adjusting the pH of the ozone effluent to 5.5-6.0 using sulfuric acid or hydrochloric acid; (S6b) heating the solution obtained in (S6a) to a temperature close to the boiling point; (S6c) deoxygenating the solution obtained in (S6b) to obtain deoxygenated effluent; (S6d) subjecting the deoxygenated effluent to evaporation and concentration treatment to obtain evaporation and concentration effluent; In step (S6c), the deoxygenation treatment is carried out in a deaerator, specifically, the solution obtained in (S6b) is sprayed through a nozzle and flows downward step by step in the deaerator, while contacting with steam flowing upward in countercurrent, to obtain the deoxygenated effluent; In step (S6d), when the evaporation concentration is performed, the deoxygenated effluent is divided into two streams, which are evaporated and concentrated in an MVR evaporation device and a TVR evaporation device respectively, and finally the obtained concentrated liquids are combined into one stream to obtain the evaporation concentration effluent; Further, the operating temperature of the TVR evaporation device is 98°C to 99.5°C, and the pressure is ≤50kPa; the operating temperature of the MVR evaporation device is 98°C to 99.5°C, and the pressure is ≤50kPa; In step (S6d), the soluble solid content in the evaporation and concentration water is 200-230 g / L.

8. The process according to any one of claims 1 to 7, characterized in that In step S7, the freeze crystallization process includes: (S7a) mixing the evaporation concentrated water with the frozen mother liquor repeatedly subjected to freeze crystallization treatment, and precooling the mixture to obtain precooled concentrated water having a temperature of 28° C. to 40° C.; (S7b) freezing and crystallizing the pre-cooled concentrated water to obtain a frozen slurry with a temperature of -5.5°C to -4.5°C; (S7c) performing solid-liquid separation on the frozen slurry to obtain sodium sulfate and frozen mother liquor, wherein the frozen mother liquor is divided into two parts, 85% to 90% of which is sent to step S9 for sodium chloride crystallization treatment, and the remaining 10% to 15% is repeatedly subjected to the frozen crystallization treatment; In step (S7b), the freezing crystallization uses ethylene glycol solution as a heat exchange medium to cool and freeze the pre-cooled concentrated water; Preferably, during freeze crystallization, the heat exchange temperature difference between the pre-cooled concentrated water and the ethylene glycol solution is ≤3°C, and the temperature difference of the ethylene glycol solution before and after the heat exchange is ≤1.5°C; In step (S7c), the solid-liquid separation is preferably centrifugal, and the SO4 2- Ion content ≤5500mg / L.

9. The process according to any one of claims 1 to 8, characterized in that In step S8, the melt crystallization process includes: (S8a) introducing water vapor into the sodium sulfate to dissolve the sodium sulfate into a supersaturated sodium sulfate solution; (S8b) mixing the supersaturated sodium sulfate solution with a molten mother liquor subjected to repeated melt crystallization treatment, and obtaining a sodium sulfate concentrated salt slurry after flash evaporation; (S8c) performing solid-liquid separation on the sodium sulfate concentrated salt slurry to obtain anhydrous sodium sulfate crystals and a molten mother liquor; (S8d) drying the anhydrous sodium sulfate crystals to obtain a sodium sulfate product; In step (S8a), the supersaturated sodium sulfate solution contains 15 wt.% of undissolved sodium sulfate crystals; In step (S8b), before the flash evaporation, the mixed solution is heated to a temperature close to the boiling point; In step (S8c), the solid-liquid separation method is centrifugation, and the water content of the anhydrous sodium sulfate crystals is ≤5%; In step (S8c), the molten mother liquor is returned to step (S8b) to be mixed with the supersaturated sodium sulfate solution.

10. The process according to any one of claims 1 to 9, characterized in that In step S9, the sodium chloride crystallization process includes: (S9a) mixing the frozen mother liquor obtained in step S7 with the sodium chloride mother liquor subjected to repeated sodium chloride crystallization treatment, and obtaining a sodium chloride concentrated salt slurry after flash evaporation; (S9b) subjecting the sodium chloride concentrated salt slurry to solid-liquid separation to obtain crude sodium chloride wet salt and sodium chloride mother liquor; (S9c) recrystallizing and drying the crude sodium chloride wet salt to obtain a sodium chloride product; Before performing the step (S9a), the frozen mother liquor is heated to 75°C ± 0.5°C; In step (S9a), the flash evaporation temperature is 95°C to 100°C, and the flash evaporation pressure based on the atmosphere is -30 to 5 kPa; In step (S9b), the sodium chloride concentrated salt slurry is preferably thickened using a cyclone separator before solid-liquid separation until the concentration of the suspended matter therein is 40wt%; In step (S9b), the solid-liquid separation method is preferably centrifugation, the moisture content of the crude sodium chloride wet salt is 5wt.%, and the chloride ion content in the sodium chloride mother liquor is ≤120g / L. In step (S9b), the sodium chloride mother liquor is sent to step S10 for impurity salt crystallization treatment, and / or sent to step (S9a) to be mixed with the frozen mother liquor, and / or sent to step S7 for freeze crystallization treatment.

11. The process according to any one of claims 1 to 10, characterized in that In step S10, the impure salt crystallization process includes: (S10a) flashing the sodium chloride mother liquor obtained in step S9 to obtain a concentrated salt slurry of impure salts; (S10b) performing solid-liquid separation on the impure salt concentrated slurry to obtain impure salt I and impure salt mother liquor; (S10c) drum drying the impure salt mother solution to obtain impure salt II; (S10d) mixing the impure salt I and the impure salt II and drying them to obtain an impure salt product; In step (S10a), before the flash evaporation, the sodium chloride mother liquor is heated to a temperature close to the boiling point, preferably the flash evaporation temperature is 90 to 105°C, and the flash evaporation pressure based on the atmosphere is -5 to 5 kPa; In step (S10b), the solid-liquid separation method is centrifugation; In step (S10c), during the drum drying, the surface temperature of the drum is 115°C to 120°C; In step (S10c), impure salt I and impure salt II are mixed, 50% of the mixture is dissolved, and then returned to step S5 for freeze crystallization, and the remaining 50% becomes an impure salt product.

12. A system for the process according to any one of claims 1 to 11, the system comprising the following units according to the material flow: The hardness and silicon removal unit is used to remove most of the hardness and silicon in the high-salinity wastewater to obtain hardness and silicon-free effluent; Turbidity removal unit, used to remove suspended matter in the effluent from the hardness removal and silicon removal process to obtain turbidity removal effluent; A softening unit is used to remove the remaining hardness in the deturbidified effluent water to obtain softened effluent water; The membrane treatment unit is used to remove the precipitated colloids, suspended solids and reduced matter in the softened effluent to obtain membrane treated effluent; Ozone oxidation unit, used to remove organic matter in membrane treated water to obtain ozone oxidized water; An evaporation concentration unit is used to concentrate the ozone oxidation effluent to obtain evaporation concentrated effluent; A freezing crystallization unit is used to freeze and crystallize the evaporation and concentration water to obtain sodium sulfate and frozen mother liquor; A melt crystallization unit is used to melt crystallize the mirabilite to obtain a sodium sulfate product and a molten mother liquor; A sodium chloride crystallization unit is used to crystallize sodium chloride from the frozen mother liquor to obtain a sodium chloride product and a sodium chloride mother liquor; The mixed salt crystallization unit is used to process part of the mother liquor discharged from the sodium chloride crystallization unit to obtain mixed salt products.

13. The system according to claim 12, characterized in that The hardness and silicon removal unit uses a reaction sedimentation tank and a V-type filter tank to treat the high-salt wastewater; The turbidity removal unit uses a high-density sedimentation tank to treat the hardness and silicon removal effluent; The softening unit uses a hydrogen-type weakly acidic cation exchanger to treat the de-turbidified water. The membrane treatment unit uses an ultrafiltration device and a reverse osmosis device to treat the softened effluent; The ozone oxidation unit uses an ozone generator to treat the membrane treated water; The evaporation concentration unit uses a TVR evaporation device and an MVR evaporation device to concentrate the ozone oxidation effluent; The freezing crystallization unit uses a crystal slurry internal circulation crystallizer to treat the evaporation and concentration water; The melt crystallization unit uses a forced circulation FC crystallizer to process the mirabilite; The sodium chloride crystallization unit uses a negative pressure vacuum forced circulation FC crystallizer to treat the frozen mother liquor; The mixed salt crystallization unit uses a forced circulation FC crystallizer and a drum dryer to treat the mother liquor discharged from the sodium chloride crystallization unit.

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