CN109422407B

CN109422407B - Method for treating catalyst production wastewater

Info

Publication number: CN109422407B
Application number: CN201710752799.3A
Authority: CN
Inventors: 殷喜平; 顾松园; 李叶; 王涛
Original assignee: China Petroleum and Chemical Corp; Sinopec Catalyst Co
Current assignee: China Petroleum and Chemical Corp; Sinopec Catalyst Co
Priority date: 2017-08-28
Filing date: 2017-08-28
Publication date: 2023-04-07
Anticipated expiration: 2037-08-28
Also published as: CN109422407A

Abstract

The invention relates to the field of sewage treatment, and discloses a method for treating catalyst production wastewater, wherein the catalyst production wastewater contains NH ₄ ⁺ 、SO ₄ ^2‑ 、Cl ^‑ And Na ⁺ The method comprises the following steps of 1) carrying out first evaporation on wastewater to be treated to obtain first ammonia-containing steam and first concentrated solution, wherein the wastewater to be treated contains the catalyst production wastewater; 2) Cooling and crystallizing the first concentrated solution to obtain a crystal liquid containing sodium sulfate crystals; 3) Carrying out first solid-liquid separation on the crystallization liquid containing the sodium sulfate crystals, and introducing a liquid phase obtained by the first solid-liquid separation into a second MVR evaporation device for second evaporation to obtain second ammonia-containing steam and a second concentrated solution containing sodium chloride crystals; 4) And carrying out second solid-liquid separation on the second concentrated solution containing the sodium chloride crystals. The method can respectively recover the ammonium, the sodium sulfate and the sodium chloride in the wastewater, and furthest recycle resources in the wastewater.

Description

Method for treating catalyst production wastewater

Technical Field

The invention relates to the field of sewage treatment, in particular to a method for treating catalyst production wastewater, and especially relates to a catalyst containing NH ₄ ⁺ 、SO ₄ ^2- 、Cl ^- And Na ⁺ The method for treating wastewater from catalyst production.

Background

In the production process of the oil refining catalyst, a large amount of inorganic acid alkali salts such as sodium hydroxide, hydrochloric acid, sulfuric acid, ammonium salts, sulfates, hydrochlorides and the like are needed, and a large amount of mixed sewage containing ammonium, sodium sulfate, sodium chloride and aluminosilicate is generated. For such sewage, the common practice in the prior art is that the pH value is adjusted to be within the range of 6-9, most of suspended matters are removed, then the biochemical method, the blow-off method or the steam stripping method is adopted to remove ammonium ions, then the salt-containing sewage is subjected to pH value adjustment, most of suspended matters are removed, hardness, silicon and part of organic matters are removed, most of organic matters are removed through ozone biological activated carbon adsorption oxidation or other advanced oxidation methods, then the salt-containing sewage enters an ion exchange device for further hardness removal, enters an enrichment device (such as reverse osmosis or electrodialysis) for concentration, and then MVR evaporative crystallization or multiple-effect evaporative crystallization is adopted to obtain mixed miscellaneous salt of sodium sulfate and sodium chloride containing a small amount of ammonium salt; or is; firstly, adjusting the pH value to be within the range of 6.5-7.5, removing most suspended matters, then removing hardness, silicon and part of organic matters, removing most organic matters through ozone biological activated carbon adsorption oxidation or other advanced oxidation methods, then entering an ion exchange device for further removing hardness, entering a thickening device (such as reverse osmosis and/or electrodialysis) for concentration, and then adopting MVR (mechanical vapor recompression) evaporative crystallization or multi-effect evaporative crystallization to obtain the mixed salt of sodium sulfate and sodium chloride containing ammonium salt. However, these ammonium-containing mixed salts are currently difficult to treat or expensive to treat, and the process of removing ammonium ions at the early stage additionally increases the cost of wastewater treatment.

In addition, the biochemical deamination can only treat wastewater with low ammonium content, and can not directly carry out biochemical treatment due to insufficient COD content in the catalyst sewage, and organic matters such as glucose or starch and the like are additionally added in the biochemical treatment process, so that the ammoniacal nitrogen can be treated by the biochemical method. The most important problems are that the total nitrogen of the wastewater after the biochemical deamination treatment is not up to the standard (the contents of nitrate ions and nitrite ions exceed the standard), advanced treatment is needed, in addition, the salt content in the wastewater is not reduced (20-30 g/L), the wastewater cannot be directly discharged, and further desalination treatment is needed.

In order to remove ammoniacal nitrogen in wastewater by gas stripping deamination, a large amount of alkali is required to adjust the pH value, the alkali consumption is high, the alkali in the deaminated wastewater cannot be recovered, the pH value of the treated wastewater is high, the treatment cost is high, the COD content in the catalyst wastewater after gas stripping is not greatly changed, the salt content in the wastewater is not reduced (20-30 g/L), the wastewater cannot be directly discharged, the desalting treatment is required to be further performed, the wastewater treatment operation cost is high, a large amount of alkali remains in the treated wastewater, the pH value is high, the waste is large, and the treatment cost is up to 50 yuan/ton.

Disclosure of Invention

The invention aims to overcome the defect of NH content in the prior art ₄ ⁺ 、SO ₄ ^2- 、Cl ^- And Na ⁺ The wastewater treatment cost is high, and only mixed salt crystals can be obtained, and a low-cost and environment-friendly NH-containing catalyst is provided ₄ ⁺ 、SO ₄ ^2- 、Cl ^- And Na ⁺ The method for treating the wastewater generated in the catalyst production can respectively recover ammonium, sodium sulfate and sodium chloride in the wastewater, and furthest recycle resources in the wastewater.

In order to achieve the above object, the present invention provides a method for treating wastewater from catalyst production containing NH ₄ ⁺ 、SO ₄ ^2- 、Cl ^- And Na ⁺ The method comprises the following steps of,

1) Performing first evaporation on wastewater to be treated to obtain first ammonia-containing steam and first concentrated solution, wherein the wastewater to be treated contains the catalyst production wastewater;

2) Cooling and crystallizing the first concentrated solution to obtain a crystal liquid containing sodium sulfate crystals;

3) Carrying out first solid-liquid separation on the crystallization liquid containing the sodium sulfate crystals, and introducing a liquid phase obtained by the first solid-liquid separation into a second MVR evaporation device for second evaporation to obtain second ammonia-containing steam and a second concentrated solution containing sodium chloride crystals;

4) Carrying out second solid-liquid separation on the second concentrated solution containing the sodium chloride crystals;

before the wastewater to be treated is subjected to first evaporation, adjusting the pH value of the wastewater to be treated to be more than 9; SO in the first concentrated solution ₄ ^2- Has a concentration of 0.01mol/L or more and Cl ^- The concentration of (b) is less than 5.2 mol/L; the second evaporation did not crystallize sodium sulfate.

By the technical scheme, the method aims at the content of NH ₄ ⁺ 、SO ₄ ^2- 、Cl ^- And Na ⁺ The pH value of the wastewater to be treated is adjusted to a specific range in advance, the wastewater is subjected to first evaporation to obtain concentrated ammonia water, then sodium sulfate crystals are obtained by cooling crystallization separation, and then sodium chloride crystals and dilute ammonia water are obtained by second evaporation. The method can respectively obtain high-purity sodium sulfate and sodium chloride, avoids the difficulty in the processes of mixed salt treatment and recycling, simultaneously completes the process of separating ammonia and salt, simultaneously heats the wastewater and cools the ammonia-containing steam by adopting a heat exchange mode without a condenser, reasonably utilizes the heat in the evaporation process, saves energy, reduces the wastewater treatment cost, recovers the ammonium in the wastewater in the form of ammonia water, recovers the sodium chloride and the sodium sulfate in the form of crystals respectively, does not generate waste residues and waste liquid in the whole process, and achieves the purpose of changing waste into valuable.

Drawings

FIG. 1 is a schematic flow diagram of a method for treating wastewater from catalyst production according to an embodiment of the present invention.

Description of the reference numerals

1. Second MVR evaporation plant 36, sixth heat transfer device

2. Cooling crystallization device 38 and eighth heat exchange device

3. First MVR evaporation plant 30, tenth heat transfer device

31. First heat exchange device 51 and first ammonia water storage tank

32. Second heat exchange device 52 and second ammonia water storage tank

33. Third heat exchange device 53 and first mother liquor tank

34. Fourth heat exchange device 54 and second mother liquor tank

61. First pH value measuring device 78 and eighth circulating pump

62. Second pH value measuring device 79 and ninth circulating pump

60. Third pH value measuring device 80 and tenth circulating pump

70. Eleventh circulating pump 81, vacuum pump

71. First circulating pump 82 and circulating water tank

72. Second circulating pump 83 and tail gas absorption tower

73. Third circulating pump 91 and first solid-liquid separation device

74. Fourth circulating pump 92 and second solid-liquid separation device

75. Fifth circulating pump 101, first compressor

76. Sixth circulation pump 102, second compressor

77. Seventh circulating pump

Detailed Description

The endpoints of the ranges and any values disclosed herein are not limited to the precise range or value, and such ranges or values should be understood to encompass values close to those ranges or values. For ranges of values, between the endpoints of each of the ranges and the individual points, and between the individual points may be combined with each other to give one or more new ranges of values, and these ranges of values should be considered as specifically disclosed herein.

The present invention will be described below with reference to fig. 1, but the present invention is not limited to fig. 1.

The invention provides a method for treating wastewater generated in catalyst production, which contains NH ₄ ⁺ 、SO ₄ ^2- 、Cl ^- And Na ⁺ The method comprises the following steps of,

2) Cooling and crystallizing the first concentrated solution to obtain a crystallization solution containing sodium sulfate crystals;

before the wastewater to be treated is subjected to first evaporation, adjusting the pH value of the wastewater to be treated to be more than 9; SO in the first concentrated solution ₄ ^2- Has a concentration of 0.01mol/L or more and Cl ^- The concentration of (A) is less than 5.2 mol/L; the second evaporation did not crystallize sodium sulfate.

Preferably, the wastewater to be treated is the catalyst production wastewater; or the wastewater to be treated contains the catalyst production wastewater and a liquid phase obtained by the second solid-liquid separation.

More preferably, the wastewater to be treated is a mixed solution of the catalyst production wastewater and at least part of a liquid phase obtained by the second solid-liquid separation.

The method provided by the invention can be used for the treatment of the ammonia-containing gas containing NH ₄ ⁺ 、SO ₄ ^2- 、Cl ^- And Na ⁺ Except that it contains NH ₄ ⁺ 、SO ₄ ^2- 、Cl ^- And Na ⁺ In addition, the catalyst production wastewater is not particularly limited.

In the present invention, the SO contained in the first concentrated solution is considered to improve the efficiency of wastewater treatment ₄ ^2- Is 0.01mol/L or more, more preferably 0.07mol/L or more, further preferably 0.1mol/L or more, further preferably 0.2mol/L or more, particularly preferably 0.3mol/L or more, and may be, for example, 0.5 to 1mol/L, and may be 0.6 to 0.7mol/L. And, cl in the first concentrated solution ^- The concentration of (B) is 5.2mol/L or less, preferably 4.5mol/L or less, more preferably 3mol/L or less, preferably 0.01mol/L or more, more preferably 0.05mol/L or more, more preferably 0.1mol/L or more, further preferably 0.5mol/L or more, further preferably 1mol/L or more, further preferably 2mol/L or more, and may be, for example, 2 to 3mol/L. By adding SO in the first concentrated solution ₄ ^2- 、Cl ^- The concentration of (3) is controlled in the above range, so that sodium sulfate can be separated out from the cooled crystal without separating out sodium chloride, thereby achieving the purpose of efficiently separating sodium sulfate.

As SO contained in the first concentrated solution ₄ ^2- Specific examples of (a) may be: 0.01mol/L, 0.03mol/L, 0.05mol/L, 0.1mol/L, 0.3mol/L, 0.4mol/L, 0.5mol/L, 0.6mol/L, 0.7mol/L, 0.8mol/L, 0.9mol/L, 1mol/L, 1.2mol/L, 1.4mol/L, 1.5mol/L, 1.8mol/L, 2mol/L, 2.2mol/L, 2.5mol/L, 2.8mol/L, or 3mol/L, etc.

As Cl contained in the first concentrated solution ^- Specific examples of (a) may be: 0.05mol/L, 0.1mol/L, 0.3mol/L, 0.5mol/L, 0.8mol/L, 1mol/L, 1.3mol/L, 1.5mol/L, 1.6mol/L, 1.7mol/L, 1.8mol/L, 2mol/L, 2.2mol/L, 2.4mol/L, 2.5mol/L, 2.8mol/L, 3mol/L, 3.2mol/L, 3.5mol/L, 3.8mol/L, 4mol/L, 4.2mol/L, 4.5mol/L, 4.8mol/L or 5mol/L, and the like.

In the present invention, the sequence of the first heat exchange, the adjustment of the pH value of the wastewater to be treated, and the preparation of the wastewater to be treated (in the case where the wastewater to be treated contains a liquid phase obtained by the separation of the catalyst production wastewater and the second solid-liquid, the preparation of the wastewater to be treated needs to be performed) is not particularly limited, and may be appropriately selected as needed, and may be completed before the first evaporation of the wastewater to be treated.

In the present invention, it is understood that the first ammonia-containing steam and the second ammonia-containing steam are so-called secondary steam in the art. The pressures are all pressures in gauge pressure.

In the present invention, the first evaporation is to concentrate the wastewater to be treated, obtain relatively concentrated ammonia water, increase the concentration of ions, and increase the precipitation rate of cooling crystals, and the degree of the first evaporation is not particularly limited, and may be selected according to the requirements of the wastewater to be treated and the components of the wastewater to be treated, so as to meet the requirements of the cooling crystals on the first concentrated solution. For example, evaporation can be controlled to obtain only a small amount of ammonia-containing steam, so as to obtain ammonia water with higher concentration; and evaporation can be controlled to be fully performed, so that the wastewater to be treated is fully concentrated, and subsequent cooling crystallization is facilitated. Preferably, sodium chloride is not crystallized out in the first evaporation.

In the present invention, the first evaporation may be performed using an evaporation apparatus conventional in the art, such as an MVR evaporation apparatus, a single-effect evaporation apparatus, or a multi-effect evaporation apparatus. As the MVR evaporation means, for example, one or more selected from the group consisting of an MVR falling film evaporator, an MVR forced circulation evaporator, an MVR-FC continuous crystallization evaporator, and an MVR-OSLO continuous crystallization evaporator may be mentioned. Among them, preferred are an MVR forced circulation evaporator and an MVR-FC continuous crystallization evaporator, and more preferred is a falling film + forced circulation two-stage MVR evaporation crystallizer. The evaporator as the single-effect evaporator or the multiple-effect evaporator may be, for example, one or more selected from falling-film evaporators, rising-film evaporators, wiped-plate evaporators, central-circulation-tube-type multiple-effect evaporators, basket-suspended evaporators, external-heat evaporators, forced-circulation evaporators and lien evaporators. Among them, a forced circulation evaporator and an external heating evaporator are preferable. The respective evaporators of the multi-effect evaporation apparatus are composed of a heating chamber and an evaporation chamber, and may further include other evaporation auxiliary components as necessary, such as a demister for further separating liquid foam, a condenser for condensing all secondary steam, and a vacuum apparatus for pressure reduction operation. The number of evaporators included in the multi-effect evaporation apparatus is not particularly limited, and may be 2 or more, and more preferably 3 to 5. According to a preferred embodiment of the present invention, said first evaporation is carried out by means of a first MVR evaporation device 3. In the invention, under the condition that the multi-effect evaporation device is used for carrying out first evaporation, when concurrent or countercurrent feeding is adopted, the condition of the first evaporation refers to the evaporation condition of the last evaporator of the multi-effect evaporation device; when advection feeding is employed, the conditions of the first evaporation include evaporation conditions of each effect evaporator of the multi-effect evaporation apparatus.

According to the invention, the conditions of the first evaporation can be properly selected according to requirements, and the purpose of concentrating the wastewater to be treated can be achieved. For example, the conditions of the first evaporation may include: the temperature is above 35 ℃ and the pressure is above-98 kPa. In order to improve the efficiency of evaporation, preferably, the first evaporation condition may include: the temperature is 75-130 ℃, and the pressure is-73 kPa-117 kPa; preferably, the conditions of the first evaporation include: the temperature is 85-130 ℃, and the pressure is-58 kPa-117 kPa; preferably, the conditions of the first evaporation include: the temperature is 95-110 ℃, and the pressure is-37 kPa-12 kPa; preferably, the conditions of the first evaporation include: the temperature is 100-107 ℃ and the pressure is-23 kPa-0 kPa.

In the present invention, the operation pressure of the first evaporation is preferably the saturated vapor pressure of the evaporated feed liquid. In addition, the evaporation amount of the first evaporation may be appropriately selected depending on the capacity of the apparatus to treat and the amount of wastewater to be treated, and may be, for example, 0.1m ³ More than h (e.g. 0.1 m) ³ /h～500m ³ /h)。

By appropriately controlling the conditions of the first evaporation, 80 mass% or more, preferably 90 mass% or more, of the ammonia contained in the wastewater to be treated can be obtained by evaporation, and for example, 80 mass%, 83 mass%, 85 mass%, 86 mass%, 87 mass%, 88 mass%, 89 mass%, 90 mass%, 91 mass%, 93 mass%, 95 mass%, or 98 mass% can be obtained, and the first aqueous ammonia can be directly recycled in the production process of the catalyst, or can be recycled after being neutralized with an acid to obtain an ammonium salt, or can be used by blending with water and a corresponding ammonium salt or aqueous ammonia.

In the present invention, the degree of progress of the first evaporation is performed by monitoring the concentration of the liquid obtained by the first evaporation, and specifically, the concentration of the liquid obtained by the first evaporation is controlled within the above range so that the first evaporation does not cause crystallization of sodium chloride. The concentration of the liquid resulting from the first evaporation is monitored by measuring the density, which may be carried out using a densitometer.

According to the present invention, before the wastewater to be treated is subjected to the first evaporation, the pH of the wastewater to be treated is adjusted to be more than 9, preferably more than 10.8, and the upper limit of the adjustment of the pH is not limited, and may be, for example, 14 or less, preferably 13.5 or less, more preferably 13 or less, further preferably 12 or less, and still further preferably 11.5 or less. By carrying out the first evaporation at the above pH, the evaporation of ammonia can be promoted, aqueous ammonia of higher concentration can be obtained, and high purity sodium sulfate and sodium chloride crystals can be conveniently obtained in the subsequent crystallization.

Specific examples of adjusting the pH of the wastewater to be treated before subjecting the wastewater to be treated to the first evaporation include: 9. 9.5, 9.6, 9.7, 9.8, 9.9, 10, 10.1, 10.2, 10.3, 10.4, 10.5, 10.6, 10.7, 10.8, 10.9, 11, 11.1, 11.2, 11.3, 11.4, 11.5, 11.6, 11.7, 11.8, 11.9, 12, 12.2, 12.4, 12.6, 12.8, 13, 13.5, or 14, etc.

In the present invention, the method of the pH adjustment is not particularly limited, and for example, the pH of the wastewater to be treated may be adjusted by adding an alkaline substance. The alkaline substance is not particularly limited, and the purpose of adjusting the pH value may be achieved. The alkaline substance is preferably NaOH in order not to introduce new impurities in the wastewater to be treated, and to increase the purity of the crystals obtained. In addition, the second mother liquor (i.e., the liquid phase obtained by the second solid-liquid separation) contains NaOH at a relatively high concentration, and it is also preferable to use the second mother liquor as the basic substance, and further NaOH may be additionally added.

The manner of adding the basic substance may be any manner known in the art, but it is preferable to mix the basic substance with the wastewater to be treated in the form of an aqueous solution, and for example, an aqueous solution containing the basic substance may be introduced into a pipe through which the wastewater to be treated is introduced and mixed. The content of the alkaline substance in the aqueous solution is not particularly limited as long as the above-mentioned purpose of adjusting the pH value can be achieved. However, in order to reduce the amount of water used and further reduce the cost, it is preferable to use a saturated aqueous solution of an alkaline substance or a second mother liquor. In order to monitor the pH value of the wastewater to be treated, the pH value of the wastewater to be treated may be measured after the above-mentioned pH value adjustment.

According to a preferred embodiment of the present invention, the first evaporation is performed in the first MVR evaporation device 3, pH adjustment is performed by introducing and mixing the aqueous solution containing the alkaline substance in the pipe that feeds the wastewater to be treated to the first MVR evaporation device 3 before feeding the wastewater to be treated to the first MVR evaporation device 3, and the adjusted pH is measured by the first pH measurement device 61 and the third pH measurement device 60 after the adjustment.

According to the present invention, in order to fully utilize the heat of the first ammonia-containing steam, it is preferable that before the wastewater to be treated is subjected to the first evaporation, the wastewater to be treated is subjected to the first heat exchange with the first ammonia-containing steam to obtain the first ammonia water, and at the same time, the temperature of the wastewater to be treated is raised to facilitate the evaporation.

According to a preferred embodiment of the present invention, the first heat exchange between the wastewater to be treated and the first ammonia-containing steam is performed by the first heat exchange device 31 and the eighth heat exchange device 38, specifically, the ammonia-containing steam is sequentially passed through the eighth heat exchange device 38 and the first heat exchange device 31, and simultaneously the wastewater to be treated is heat-exchanged with the first ammonia-containing steam condensate by the first heat exchange device 31, and then is heat-exchanged with the first ammonia-containing steam by the eighth heat exchange device 38. Through the first heat exchange, the obtained first ammonia water is stored in the first ammonia water storage tank 51, and meanwhile, the temperature of the wastewater to be treated is raised to 74-129 ℃, preferably 94-109 ℃, so that the evaporation is convenient to carry out.

According to the present invention, in order to fully utilize the heat of the first concentrated solution, it is preferable that before the wastewater to be treated is subjected to the first evaporation, the wastewater to be treated is subjected to the first heat exchange with the first concentrated solution, so that the temperature of the first concentrated solution is lowered to facilitate cooling crystallization, and the temperature of the wastewater to be treated is raised to facilitate evaporation.

According to a preferred embodiment of the present invention, the first heat exchange between the wastewater to be treated and the first concentrated solution is performed by the eleventh heat exchange device 30.

The first heat exchange device 31, the eleventh heat exchange device 30 and the eighth heat exchange device 38 are not particularly limited, and various heat exchangers conventionally used in the art may be used to perform heat exchange. Specifically, a jacketed heat exchanger, a plate heat exchanger, a shell-and-tube heat exchanger, or the like may be mentioned, with the plate heat exchanger being preferred. The material of the heat exchanger can be specifically selected according to the needs, for example, in order to resist the corrosion of chloride ions, the heat exchanger made of duplex stainless steel, titanium alloy and hastelloy can be selected, and the heat exchanger made of plastic can be selected when the temperature is lower.

In the present invention, in order to increase the concentration of the liquid in the first MVR evaporation device 3 and reduce the ammonia content in the liquid, it is preferable that a part of the liquid evaporated by the first MVR evaporation device 3 (i.e. the liquid located inside the first MVR evaporation device, also referred to as the first circulation liquid) is heated and then returned to the first MVR evaporation device 3 for evaporation. The above process of returning the first circulation liquid to the first MVR evaporation device 3 is preferably to return the first circulation liquid to the first MVR evaporation device 3 after mixing the first circulation liquid with the wastewater to be treated before the first heat exchange is completed, for example, the first circulation liquid may be returned to the wastewater conveying pipeline between the first heat exchange device 31 and the eighth heat exchange device 38 by the fifth circulation pump 75 to be mixed with the wastewater to be treated, and then after the second pH adjustment, heat exchange is performed by the eighth heat exchange device 38, and the pH is monitored by the third pH measurement device 60, and finally sent to the first MVR evaporation device 3. Here, the first reflux ratio means: the ratio of the amount of reflux to the total amount of liquid fed to the first MVR evaporator 3 minus the amount of reflux. The reflux ratio may be set appropriately according to the evaporation amount to ensure that the first MVR evaporation device 3 can evaporate the required amount of water and ammonia at a given evaporation temperature. The reflux ratio of the evaporation may be, for example, 10 to 200, preferably 50 to 100.

According to the present invention, preferably, the method further comprises compressing the first ammonia-containing vapor before the first heat exchange. The compression of the first ammonia-containing vapor may be performed by a first compressor 101. Through right first ammonia vapor that contains compresses, for input energy among the MVR vaporization system, guarantee that waste water heaies up-evaporation-cooling's process and goes on in succession, need input start-up steam when MVR vaporization process starts, only through first compressor 101 energy supply after reaching continuous running state, no longer need input other energy. The first compressor 101 may employ various compressors conventionally used in the art, such as a centrifugal fan, a turbine compressor, or a roots compressor. After being compressed by the first compressor 101, the temperature of the first ammonia-containing vapor is raised by 5 to 20 ℃.

In the present invention, the purpose of the cooling crystallization is to precipitate sodium sulfate, but not sodium chloride, so that sodium sulfate can be separated from the catalyst production wastewater well. The cooling crystallization only precipitates sodium sulfate and does not exclude sodium chloride entrained by the sodium sulfate crystals or adsorbed on the surface. In the present invention, the content of sodium sulfate in the obtained sodium sulfate crystals is preferably 92% by mass or more, more preferably 96% by mass or more, and still more preferably 98% by mass or more, and it is understood that the amount of the obtained sodium sulfate crystals is based on a dry basis. When the content of sodium sulfate in the obtained sodium sulfate crystal is within the above range, it is considered that only sodium sulfate is precipitated.

The pH value of the wastewater to be treated is more than 9 by adjusting the pH value before the first evaporation, wherein the NH is ₄ ⁺ Most of the ammonia molecules are evaporated during the first evaporation, so that ammonium sulfate and/or ammonium chloride are not separated out in the cooling crystallization process, and the separation rate of sodium sulfate can be improved due to the increase of the concentration of the sodium sulfate and the sodium chloride.

In the present invention, the conditions for the cooling crystallization are not particularly limited and may be appropriately selected as necessary, and the effect of crystallizing the sodium sulfate may be obtained. The cooling crystallization conditions may include: the temperature is-21.7-17.5 ℃, preferably-20-5 ℃, more preferably-10-5 ℃, further-10-0 ℃, and particularly preferably-4-0 ℃; the time (in terms of the residence time in the cooling crystallization apparatus 2) is 5min or more, preferably 60min to 180min, more preferably 90min to 150min, and still more preferably 120min to 150min. By controlling the cooling crystallization conditions within the above range, sodium sulfate can be sufficiently precipitated.

Specific examples of the temperature for cooling and crystallizing include: -21 ℃, -20 ℃, -19 ℃, -18 ℃, -17 ℃, -16 ℃, -15 ℃, -14 ℃, -13 ℃, -12 ℃, -11 ℃, -10 ℃, -9 ℃, -8 ℃, -7 ℃, -6 ℃, -5 ℃, -4 ℃, -3 ℃, -2 ℃, -1 ℃ or 0 ℃.

Specific examples of the time for cooling crystallization include: 5min, 6min, 7min, 8min, 10min, 15min, 20min, 25min, 30min, 35min, 40min, 45min, 50min, 52min, 54min, 56min, 58min, 60min, 65min, 70min, 75min, 80min, 85min, 90min, 95min, 100min, 105min, 110min, 115min, 120min, 130min, 140min, 150min, or 160min.

According to the present invention, the cooling crystallization is carried out in a continuous or batch manner, and the temperature of the first concentrated solution is lowered to precipitate sodium sulfate crystals, and the continuous cooling crystallization is preferably carried out. The cooling crystallization of sodium sulfate can be carried out by various cooling crystallization devices conventionally used in the art, for example, by using a continuous cooling crystallizer with an external cooling heat exchanger, or by using a crystallization tank having a cooling means, such as the cooling crystallization device 2. The cooling part can lead the first concentrated solution in the cooling crystallization device to be cooled to the condition required by cooling crystallization by introducing a cooling medium. The cooling crystallization device is preferably provided with a mixing part, such as a stirrer, and the like, so that the first concentrated solution is mixed to achieve the effect of uniform cooling, sodium sulfate in the first concentrated solution can be fully precipitated, and the size of crystal grains can be increased. The cooling crystallization device is preferably provided with a circulating pump, so as to avoid generating a large amount of fine crystal nuclei and prevent crystal grains in circulating crystal slurry from colliding with the impeller at a high speed to generate a large amount of secondary crystal nuclei, and the circulating pump is preferably a low-rotation-speed centrifugal pump, more preferably a high-flow and low-rotation-speed guide pump impeller or a high-flow, low-lift and low-rotation-speed axial pump.

By carrying out the cooling crystallization under the above-mentioned conditions, sodium sulfate can be sufficiently precipitated in the cooling crystallization without precipitating sodium chloride, thereby achieving the object of separating and purifying sodium sulfate.

According to the present invention, before the first concentrated solution is cooled and crystallized, the concentration of sodium chloride in the first concentrated solution may be adjusted as necessary so that the concentration of sodium chloride in the crystal solution obtained by cooling and crystallization is X or less, where X is the concentration of sodium chloride at which both sodium sulfate and sodium chloride in the crystal solution are saturated under the condition of cooling and crystallization. By adjusting the concentration of sodium chloride in the first concentrated solution to the above range, it can be ensured that sodium chloride is not precipitated in the cooled crystals. Preferably, the concentration of sodium chloride in the crystallization liquid is made to be 0.95X-0.999X. The method of adjusting the sodium chloride concentration may be performed by mixing the second mother liquor, water, the catalyst production wastewater and/or other waste liquid generated during the treatment, etc. at an appropriate concentration, as long as the concentration of sodium chloride in the first concentrated solution is adjusted to be within the above range. To avoid introducing more liquid, preferably, adjusting the concentration of sodium chloride in the first concentrate is performed by mixing the second mother liquor, the catalyst production wastewater and/or a wash liquor rinsing the sodium sulfate crystals. By adjusting the concentration of the sodium chloride in the first concentrated solution to the range, the precipitation of the sodium chloride in the cooling crystallization process is avoided, the purity of the sodium sulfate precipitated in the cooling crystallization process is improved, and the cooling crystallization efficiency is improved.

In the present invention, in order to control the crystal size distribution in the cooling crystallization device 2 and reduce the content of fine crystal grains, it is preferable that a part of the liquid crystallized by the cooling crystallization device 2 (i.e., the liquid located inside the cooling crystallization device 2, hereinafter also referred to as cooling circulation liquid) is mixed with the wastewater to be treated and then returned to the cooling crystallization device 2 for cooling crystallization again. The process of returning the cooling circulation liquid to the cooling crystallization device 2 for crystallization can be performed by, for example, returning the cooling circulation liquid to the sixth heat exchange device 36 through the second circulation pump 72, mixing the cooling circulation liquid with the first concentrated liquid, and then returning the mixed cooling circulation liquid to the cooling crystallization device 2 for cooling crystallization. The return amount of the cooling circulation liquid can be defined by the circulation ratio of the cooling crystallization, and the circulation ratio of the cooling crystallization is as follows: the ratio of the circulating amount to the total amount of the liquid fed to the cooling crystallization device minus the circulating amount. The circulation ratio may be appropriately set according to the supersaturation degree of sodium sulfate in the cooling crystallization apparatus 2 to ensure the particle size of sodium sulfate crystals. In order to control the particle size distribution of crystals obtained by cooling crystallization and to reduce the content of fine crystal grains, it is preferable to control the supersaturation degree to less than 1.5g/L, more preferably to less than 1g/L.

In the invention, the sodium sulfate crystal and the first mother liquor (i.e. the liquid phase obtained by the first solid-liquid separation) are obtained after the first solid-liquid separation is carried out on the crystallization liquid containing the sodium sulfate crystal. The method of the first solid-liquid separation is not particularly limited, and may be selected from, for example, one or more of centrifugation, filtration, and sedimentation.

According to the present invention, the first solid-liquid separation may be performed using a first solid-liquid separation device (e.g., a centrifuge, a belt filter, a plate filter, etc.) 91. After the first solid-liquid separation, the first mother liquor obtained by the first solid-liquid separation device 91 is temporarily stored in the first mother liquor tank 53, and can be sent to the second MVR evaporation device 1 by the sixth circulation pump 76 for second evaporation. In addition, it is difficult to avoid that the obtained sodium sulfate crystals adsorb impurities such as chlorine ions, free ammonia, and hydroxide ions to a certain extent, and in order to remove the adsorbed impurities, reduce the odor of solid salts, reduce corrosiveness, and improve the purity of the crystals, the sodium sulfate crystals are preferably subjected to a first washing with water or a sodium sulfate solution, and may be dried when anhydrous sodium sulfate is required to be obtained.

The form of the first solid-liquid separation and the first washing is not particularly limited, and may be carried out, for example, by using a solid-liquid separation apparatus which is conventional in the art, or may be carried out in a staged solid-liquid separation apparatus. The washing is not particularly limited and may be carried out by a method conventional in the art. The first washing method is preferably rinsing, and rinsing is preferably performed after solid-liquid separation. The number of washing is not particularly limited, and may be 1 or more, and is preferably 2 to 4 times in order to obtain sodium sulfate crystals of higher purity. The first washing is preferably carried out using an aqueous sodium sulfate solution (the concentration of which is preferably such that the sodium chloride and the sodium sulfate reach the concentration of sodium sulfate in the saturated aqueous solution at the same time at the temperature corresponding to the sodium sulfate crystals to be washed). The liquid generated by washing is preferably returned to the cooling crystallization device 2, and for example, may be mixed with the first concentrated solution before being returned to the sixth heat exchange device 36 by the eighth circulation pump 78, and then returned to the cooling crystallization device 2.

According to a preferred embodiment of the present invention, after cooling and crystallizing the obtained crystal liquid containing sodium sulfate, solid-liquid separation is performed by a solid-liquid separation apparatus, and the crystal obtained by the solid-liquid separation is rinsed again with an aqueous sodium sulfate solution (the concentration of the aqueous sodium sulfate solution is the concentration of sodium sulfate in an aqueous solution in which sodium chloride and sodium sulfate are simultaneously saturated at a temperature corresponding to the sodium sulfate crystal to be washed), and the rinsed liquid is returned to the cooling and crystallizing apparatus 2. By the above washing process, the purity of the obtained sodium sulfate crystal can be improved.

According to the present invention, in order to fully utilize the refrigeration capacity of the first mother liquor, it is preferable that the first mother liquor and the first concentrated solution are subjected to the second heat exchange before the first concentrated solution is subjected to the cooling crystallization.

According to a preferred embodiment of the present invention, the second heat exchange is performed by a second heat exchange device 32, and specifically, the first mother liquor and the first concentrated liquor are respectively passed through the second heat exchange device 32 and heat exchanged, so that the temperature of the first concentrated liquor is lowered to facilitate the cooling crystallization, and the temperature of the first mother liquor is raised to facilitate the second evaporation. After the second heat exchange is performed by the second heat exchange device 32, the temperature of the first concentrated solution is-20.7 ℃ to 16.5 ℃, preferably-5 ℃ to 10 ℃, and is close to the temperature of cooling crystallization.

According to the invention, in order to facilitate the cooling crystallization, the first concentrated solution and the refrigerating fluid are subjected to second heat exchange. According to a preferred embodiment of the present invention, the second heat exchange between the first concentrated solution and the refrigerating fluid is performed by the sixth heat exchange device 36, and specifically, the refrigerating fluid and the first concentrated solution are respectively passed through the sixth heat exchange device 36 to perform heat exchange therebetween, so that the temperature of the first concentrated solution is lowered to facilitate the cooling crystallization. The refrigerating fluid can be the refrigerating fluid which is used for cooling conventionally in the field, as long as the temperature of the first concentrated solution can meet the requirement of cooling crystallization.

The second heat exchanger 32 and the sixth heat exchanger 36 are not particularly limited, and various heat exchangers conventionally used in the art may be used to perform heat exchange. Specifically, it may be a jacketed heat exchanger, a plate heat exchanger, a shell-and-tube heat exchanger, or the like, with the plate heat exchanger being preferred. The material of the heat exchanger can be specifically selected according to the needs, for example, in order to resist the corrosion of chloride ions, the heat exchanger of duplex stainless steel, titanium and titanium alloy, hastelloy can be selected as the material, and the heat exchanger containing plastic material can be selected when the temperature is lower. The second heat exchange device 32 is preferably a heat exchanger made of plastic.

In the present invention, the second MVR evaporation device 1 is not particularly limited, and may be various MVR evaporation devices conventionally used in the art. For example, it may be one or more selected from the group consisting of an MVR falling film evaporator, an MVR forced circulation evaporator, an MVR-FC continuous crystallization evaporator, and an MVR-OSLO continuous crystallization evaporator. Among them, preferred are an MVR forced circulation evaporator and an MVR-FC continuous crystallization evaporator, and more preferred is a falling film + forced circulation two-stage MVR evaporation crystallizer.

In the present invention, the evaporation conditions of the second evaporation may be appropriately selected as needed, and the purpose of precipitating crystals may be achieved. The conditions of the second evaporation may include: the temperature is above 17.5 ℃, and the pressure is above-101 kPa; preferably, the conditions of evaporation include: the temperature is 35-110 ℃, and the pressure is-98 kPa-12 kPa; preferably, the conditions of the evaporation include: the temperature is 45-110 ℃, and the pressure is-95 kPa-12 kPa; preferably, the conditions of evaporation include: the temperature is 50-100 ℃, and the pressure is-93 kPa to-22 kPa.

In the present invention, the operation pressure of the second evaporation is preferably the saturated vapor pressure of the evaporated feed liquid. Further, the evaporation amount of the second evaporation may be appropriately selected depending on the capacity of the apparatus to treat and the amount of the waste water to be treated, and may be, for example, 0.1m ³ More than h (e.g. 0.1 m) ³ /h～500m ³ /h)。

In order to ensure that sodium chloride crystals can be obtained in the second evaporation process, it is necessary to further satisfy the requirement that 1mol of SO contained in the liquid phase obtained by the first solid-liquid separation is contained ₄ ^2- Cl contained in the liquid phase obtained by the first solid-liquid separation ^- 9.5mol or more, preferably 10mol or more, for example, 10 to 20mol. By reacting SO ₄ ^2- And Cl ^- The molar ratio of sodium chloride is controlled in the range, pure sodium chloride crystals can be obtained through secondary evaporation, and separation of sodium sulfate and sodium chloride is realized.

According to the present invention, from the viewpoint of improving the efficiency of wastewater treatment, the higher the degree of progress of the second evaporation, the better; however, if the second evaporation exceeds a certain level, a second concentrated solution containing only sodium chloride crystals cannot be obtained, and in this case, the crystals may be dissolved by adding water to the second concentrated solution, but the efficiency of wastewater treatment is impaired. Therefore, the second evaporation is preferably performed to such an extent that sodium sulfate crystals are not crystallized, that is, the second evaporation is performed so that the concentration of sodium sulfate in the second concentrated solution becomes Y or less (where Y is the concentration of sodium sulfate at which both sodium sulfate and sodium chloride in the second concentrated solution are saturated under the conditions of the second evaporation). In the second evaporation step, the concentration of sodium sulfate in the second concentrated solution is preferably 0.9Y to 0.99Y, more preferably 0.95Y to 0.98Y, from the viewpoint of precipitating sodium chloride as much as possible without precipitating sodium sulfate. By controlling the degree of the second evaporation within the above range, it is possible to ensure that sodium chloride is precipitated as much as possible during the second evaporation, and sodium sulfate is not precipitated, and pure sodium chloride crystals are finally separated. By crystallizing sodium chloride in the second evaporation as much as possible, the wastewater treatment efficiency can be improved, and the energy can be saved.

In the present invention, the degree of progress of the second evaporation is performed by monitoring the concentration of the liquid obtained by the second evaporation, and specifically, the concentration of the liquid obtained by the second evaporation is controlled to be in the above range so that the second evaporation does not cause precipitation of sodium sulfate crystals. The concentration of the liquid resulting from the second evaporation is monitored by measuring the density, which may be carried out using a densitometer.

In the present invention, in order to increase the solids content in the second MVR evaporation device 1 and reduce the ammonia content in the liquid, it is preferred that part of the liquid evaporated by the second MVR evaporation device 1 (i.e. the liquid located inside the second MVR evaporation device, hereinafter also referred to as circulation liquid) is returned to the second MVR evaporation device 1 for evaporation, preferably heated and then returned to the second MVR evaporation device 1 for evaporation. The above-mentioned process of returning the circulating liquid to the second MVR evaporation plant 1 can be, for example, returned to the third heat exchange process by the seventh circulating pump 77, for example, returned to the wastewater delivery line between the third heat exchange device 33 and the fourth heat exchange device 34 to be mixed with the first mother liquid, and then heat exchanged in the fourth heat exchange device 34, and then returned to the second MVR evaporation plant 1. The second reflux ratio of the second evaporation is: the ratio of the amount of reflux to the total amount of liquid fed to the second MVR evaporator 1 minus the amount of reflux. The second reflux ratio can be set appropriately according to the evaporation amount to ensure that the second MVR evaporation device can evaporate the required amount of water and ammonia at the given second evaporation temperature. The second reflux ratio of the second evaporation may be, for example, 10 to 200, preferably 40 to 150.

According to the present invention, preferably, the method further comprises compressing the second ammonia-containing vapor before the third heat exchange. The compression of the second ammonia-containing vapor may be performed by a second compressor 102. The second ammonia-containing steam is compressed, energy is input into the MVR evaporation system, the continuous process of waste water heating, evaporation and cooling is guaranteed, starting steam needs to be input when the MVR evaporation process is started, energy is supplied only through the second compressor 102 after the continuous running state is achieved, and other energy does not need to be input. The second compressor 102 may employ various second compressors conventionally used in the art, such as a centrifugal fan, a turbine compressor, a roots compressor, or the like. After compression by the second compressor 102, the temperature of the second ammonia-containing vapor is raised by 5 to 20 ℃.

According to the present invention, in order to fully utilize the heat in the second ammonia-containing vapor obtained by the second evaporation, it is preferable to subject the first mother liquor to a third heat exchange with the second ammonia-containing vapor before sending the first mother liquor to the second MVR evaporation device 1.

According to a preferred embodiment of the present invention, the third heat exchange of the first mother liquor with the second ammonia-containing steam is carried out by a third heat exchange means 33 and a fourth heat exchange means 34, respectively. Specifically, the first mother liquor sequentially passes through the third heat exchange device 33 and the fourth heat exchange device 34, and the second ammonia-containing steam sequentially passes through the fourth heat exchange device 34 and the third heat exchange device 33, so that the temperature of the first mother liquor is raised, the second evaporation is facilitated, and the second ammonia-containing steam is condensed to obtain ammonia water. After heat exchange is carried out by the third heat exchange device 33, the temperature of the first mother liquor is raised to 34-109 ℃, preferably 44-109 ℃; after heat exchange by the fourth heat exchange device 34, the temperature of the first mother liquor is raised to 42 ℃ to 117 ℃, preferably 52 ℃ to 117 ℃.

The third heat exchange device 33 and the fourth heat exchange device 34 are not particularly limited, and various heat exchangers conventionally used in the art may be used to exchange heat between the second ammonia-containing steam and the first mother liquor. Specifically, a jacketed heat exchanger, a plate heat exchanger, a shell-and-tube heat exchanger, or the like may be mentioned, with the plate heat exchanger being preferred. The material of the heat exchanger can be specifically selected according to the needs, for example, in order to resist the corrosion of chloride ions, the heat exchanger of duplex stainless steel, titanium and titanium alloy, hastelloy can be selected as the material, and the heat exchanger containing plastic material can be selected when the temperature is lower. Preferably, a duplex stainless steel plate heat exchanger is used.

According to the present invention, the pH of the first mother liquor is preferably monitored before it is passed into the second MVR evaporator 1, for example by the second pH measuring device 62.

According to the present invention, the second solid-liquid separation may be performed by a second solid-liquid separation device (e.g., a centrifuge, a belt filter, a plate filter, etc.) 92. After the second solid-liquid separation, the second mother liquor obtained by the second solid-liquid separation device 92 (i.e., the liquid phase obtained by the second solid-liquid separation) returns to the first MVR device 3 for evaporation again, and specifically, the second mother liquor may be returned to the first MVR device by the ninth circulation pump 79 before the second pH adjustment. In addition, it is difficult to avoid that the obtained sodium chloride crystals adsorb certain impurities such as sulfate ions, free ammonia, hydroxide ions, etc., and in order to remove the adsorbed impurities, reduce the odor of solid salts, reduce corrosiveness, and improve the purity of the crystals, the sodium chloride crystals are preferably subjected to secondary washing with water, the catalyst production wastewater, or a sodium chloride solution and dried. In order to avoid dissolution of the sodium chloride crystals during washing, preferably the sodium chloride crystals are washed with an aqueous solution of sodium chloride. More preferably, the concentration of the aqueous sodium chloride solution is preferably the concentration of sodium chloride in an aqueous solution in which sodium chloride and sodium sulphate reach saturation simultaneously at the temperature corresponding to the sodium chloride crystals to be washed.

The form of the second solid-liquid separation and the second washing is not particularly limited, and may be carried out, for example, by using a combination of an elutriation apparatus and a solid-liquid separation apparatus which are conventional in the art, or may be carried out on a staged solid-liquid separation apparatus such as a belt filter. The second washing comprises elutriation and/or rinsing, and the second washing mode is preferably that elutriation is carried out before rinsing. The above-mentioned elutriation and rinsing are not particularly limited and may be carried out by a method conventional in the art. The number of elutriation and rinsing is not particularly limited, and may be 1 or more, and is preferably 2 to 4 times in order to obtain sodium chloride crystals of higher purity. In the elutriation process, the washing liquid recovered by the second washing can be used in a countercurrent manner when used as the elutriation liquid. Before the elutriation, it is preferable to perform preliminary solid-liquid separation by sedimentation to obtain a slurry containing sodium chloride crystals (the liquid content may be 35% by mass or less). In the elutriation process, 1 to 20 parts by weight of a liquid is used for elutriation with respect to 1 part by weight of a slurry containing sodium chloride crystals. The rinsing is preferably carried out using an aqueous sodium chloride solution, the concentration of which is preferably the concentration of sodium chloride in an aqueous solution in which sodium chloride and sodium sulfate are simultaneously saturated at the temperature corresponding to the sodium chloride crystals to be rinsed. In order to further enhance the elutriation effect and obtain sodium chloride crystals with higher purity, it is preferable to perform elutriation using the eluted solution. For the liquid generated by washing, the catalyst production wastewater elutriation liquid is returned to the second MVR evaporation device 1 before the pH adjustment before the first evaporation.

According to a preferred embodiment of the present invention, after the initial solid-liquid separation by settling of the second concentrate containing sodium chloride crystals, the second concentrate is subjected to elutriation in another elutriation tank with the liquid obtained in the subsequent washing of the sodium chloride crystals, the slurry is sent to a solid-liquid separation device for solid-liquid separation, the crystals obtained by solid-liquid separation are washed with an aqueous sodium chloride solution (the concentration of the aqueous sodium chloride solution is the concentration of sodium chloride in an aqueous solution in which sodium chloride and sodium sulfate are simultaneously saturated at the temperature corresponding to the sodium chloride crystals to be washed), and the liquid obtained by washing is returned to the elutriation as an elutriation liquid. Through the washing process combining elutriation and leaching, the purity of the obtained sodium chloride crystal is improved, washing liquid cannot be introduced too much, and the efficiency of wastewater treatment is improved.

According to a preferred embodiment of the invention, the tail gas generated by cooling crystallization is discharged after ammonia removal; discharging the tail gas which is remained by the third heat exchange condensation after ammonia removal; and discharging the tail gas which is remained by the condensation of the first heat exchange after ammonia removal. The tail gas generated by the cooling crystallization is the tail gas discharged by the cooling crystallization device 2, and the third heat exchange condenses the residual tail gas which is the non-condensable gas discharged by the fourth heat exchange device 34; the first heat exchange ring condenses the remaining tail gas, i.e., the tail gas discharged from the eighth heat exchange device 38. By removing ammonia from the tail gas, the content of pollutants in the discharged tail gas can be further reduced, so that the tail gas can be directly discharged.

As the method of removing ammonia, absorption may be performed by the off-gas absorption tower 83. The off-gas absorption column 83 is not particularly limited, and may be any of various absorption columns conventionally used in the art, such as a plate-type absorption column, a packed absorption column, a falling film absorption column, or an empty column. Circulating water is arranged in the tail gas absorption tower 83, the circulating water circulates in the tail gas absorption tower 83 under the action of the fourth circulating pump 74, water can be supplemented into the tail gas absorption tower 83 from the circulating water tank 82 through the third circulating pump 73, fresh water can be supplemented into the circulating water tank 82, and meanwhile the temperature of working water of the vacuum pump 81 and the ammonia content can be reduced. The flow of the off-gas and the circulating water in the off-gas absorption tower 83 may be in a counter-current or co-current flow, preferably in a counter-current flow. The circulating water can be supplemented by additional fresh water. In order to ensure the sufficient absorption of the tail gas, dilute sulfuric acid may be further added to the tail gas absorption tower 83 to absorb a small amount of ammonia and the like in the tail gas. The circulating water can be used as ammonia water or ammonium sulfate solution for production or direct sale after absorbing tail gas. The off gas may be introduced into the off gas absorption tower 83 by a vacuum pump 81.

In the present invention, the catalyst production wastewater is not particularly limited as long as it contains NH ₄ ⁺ 、SO ₄ ^2- 、Cl ^- And Na ⁺ The wastewater is obtained. In addition, the method is particularly suitable for treating high-salt ammonium-containing wastewater. The wastewater from the catalyst production of the present invention may be specifically wastewater from the production of a molecular sieve, alumina or an oil refining catalyst, or wastewater from the production of a molecular sieve, alumina or an oil refining catalyst after the following impurity removal and concentration. It is preferable that the wastewater from the production of molecular sieves, alumina or refinery catalysts is subjected to the following impurity removal and concentration.

As NH in the catalyst production wastewater ₄ ⁺ May be 8mg/L or more, preferably 300mg/L or more.

As Na in the wastewater from the catalyst production ⁺ May be 510mg/L or more, preferably 1g/L or more, more preferably 2g/L or more, further preferably 4g/L or more, further preferably 8g/L or more, further preferably 16g/L or more, further preferably 32g/L or more,more preferably 40g/L or more, still more preferably 50g/L or more, and still more preferably 60g/L or more.

As SO in wastewater from the production of said catalyst ₄ ^2- May be 1g/L or more, preferably 2g/L or more, more preferably 4g/L or more, further preferably 8g/L or more, further preferably 16g/L or more, further preferably 32g/L or more, further preferably 40g/L or more, further preferably 50g/L or more, further preferably 60g/L or more, further preferably 70g/L or more.

As Cl in the catalyst production wastewater ^- May be 970mg/L or more, more preferably 2g/L or more, further preferably 4g/L or more, further preferably 8g/L or more, further preferably 16g/L or more, further preferably 32g/L or more, further preferably 40g/L or more, further preferably 50g/L or more, further preferably 60g/L or more.

NH contained in the catalyst production wastewater ₄ ⁺ 、SO ₄ ^2- 、Cl ^- And Na ⁺ The upper limit of (3) is not particularly limited. SO in the wastewater from catalyst production from the viewpoint of easy wastewater treatment ₄ ^2- 、Cl ^- And Na ⁺ The upper limit of (b) is 200g/L or less, preferably 150g/L or less, respectively; NH in catalyst production wastewater ₄ ⁺ Is 50g/L or less, preferably 30g/L or less.

From the viewpoint of reducing energy consumption in the treatment process, SO contained in the catalyst production wastewater ₄ ^2- With Cl ^- The higher the molar ratio of (A), (relative to 1mol of SO), the better ₄ ^2- ，Cl ^- Preferably 10mol or less, more preferably 8mol or less, and for example, may be 1.3 to 6.5 mol/L), from the viewpoint of improving the purity of the sodium sulfate product, cl contained in the wastewater from the catalyst production ^- Is 5.23mol/L or less (preferably 5.09mol/L or less, more preferably 4.5mol/L or less, and may be, for example, 1.8 to 3.9 mol/L). By adding SO contained in the catalyst production wastewater ₄ ^2- And Cl ^- The concentration of (B) is limited to the above range, and sodium sulfate having an extremely high purity can be obtained by crystallization in the course of cooling crystallizationAnd energy is saved, so that the treatment process is more economical.

In the present invention, the inorganic salt ions contained in the catalyst production wastewater are other than NH ₄ ⁺ 、SO ₄ ^2- 、Cl ^- And Na ⁺ In addition, it may contain Mg ²⁺ 、Ca ²⁺ 、K ⁺ 、Fe ³⁺ Inorganic salt ions such as rare earth element ions, mg ²⁺ 、Ca ²⁺ 、K ⁺ 、Fe ³⁺ The content of each inorganic salt ion such as a rare earth element ion is preferably 100mg/L or less, more preferably 50mg/L or less, still more preferably 10mg/L or less, and particularly preferably no other inorganic salt ion is contained. By controlling the other inorganic salt ions within the above range, the purity of the sodium sulfate crystals and ammonium chloride crystals finally obtained can be further improved. In order to reduce the content of other inorganic salt ions in the catalyst production wastewater, the following impurity removal is preferably performed.

The TDS of the catalyst production wastewater may be 1.6g/L or more, preferably 4g/L or more, more preferably 8g/L or more, further preferably 16g/L or more, further preferably 32g/L or more, further preferably 40g/L or more, further preferably 50g/L or more, further preferably 60g/L or more, further preferably 100g/L or more, further preferably 150g/L or more, further preferably 200g/L or more.

In the present invention, the pH of the catalyst production wastewater is preferably 4 to 8, such as 6 to 8, for example 6.6 to 7.2.

In addition, since the COD of the catalyst production wastewater may block a membrane at the time of concentration, affect the purity and color of a salt at the time of evaporative crystallization, etc., the COD of the catalyst production wastewater is preferably as small as possible (preferably 20mg/L or less, more preferably 10mg/L or less), and is preferably removed by oxidation at the time of pretreatment, and specifically, it may be carried out by, for example, a biological method, an advanced oxidation method, etc., and it is preferably oxidized by an oxidizing agent such as fenton's reagent at the time of very high COD content.

In the invention, in order to reduce the concentration of impurity ions in the wastewater, ensure the continuous and stable treatment process and reduce the equipment operation and maintenance cost, the catalyst production wastewater is preferably subjected to impurity removal before being treated by the treatment method. Preferably, the impurity removal is selected from one or more of solid-liquid separation, chemical precipitation, adsorption, ion exchange and oxidation.

As the solid-liquid separation, filtration, centrifugation, sedimentation, or the like may be mentioned; the chemical precipitation may be pH adjustment, carbonate precipitation, magnesium salt precipitation, or the like; the adsorption can be physical adsorption and/or chemical adsorption, and the specific adsorbent can be selected from activated carbon, silica gel, alumina, molecular sieve, natural clay and the like; as the ion exchange, any one of a strongly acidic cation resin and a weakly acidic cation resin can be used; as the oxidation, various oxidizing agents conventionally used in the art, such as ozone, hydrogen peroxide, and potassium permanganate, can be used, and in order to avoid introduction of new impurities, ozone, hydrogen peroxide, and the like are preferably used.

The specific impurity removal mode can be specifically selected according to the types of impurities contained in the catalyst production wastewater. Aiming at suspended matters, a solid-liquid separation method can be selected for removing impurities; for inorganic matters and organic matters, chemical precipitation, ion exchange and adsorption methods can be selected for removing impurities, such as weak acid cation exchange, activated carbon adsorption and the like; for organic matters, impurities can be removed by adopting an adsorption and/or oxidation mode, wherein an ozone biological activated carbon adsorption oxidation method is preferred. According to a preferred embodiment of the invention, the catalyst production wastewater is subjected to impurity removal by filtration, a weak acid cation exchange method and an ozone biological activated carbon adsorption oxidation method in sequence. Through the impurity removal process, most suspended matters, hardness, silicon and organic matters can be removed, the scaling risk of the device is reduced, and the continuous and stable operation of the wastewater treatment process is ensured.

In the present invention, the catalyst production wastewater having a low salt content may be concentrated to have a salt content within a range required for the wastewater of the present invention before the treatment by the treatment method of the present invention (preferably, after the above-mentioned removal of impurities). Preferably, the concentration is selected from ED membrane concentration and/or reverse osmosis; more preferably, the concentration is performed by ED membrane concentration and reverse osmosis, and the order of performing the ED membrane concentration and reverse osmosis is not particularly limited. The ED membrane concentration and reverse osmosis treatment apparatus and conditions may be performed in a manner conventional in the art, and may be specifically selected according to the condition of wastewater to be treated. Specifically, as the concentration of the ED membrane, a one-way electrodialysis system or a reversed electrodialysis system can be selected for carrying out; as the reverse osmosis, a roll membrane, a plate membrane, a disc-tube membrane, a vibrating membrane or a combination thereof can be selected for use. Through the concentration, the efficiency of wastewater treatment can be improved, and energy waste caused by a large amount of evaporation is avoided.

In a preferred embodiment of the invention, the catalyst production wastewater is wastewater obtained by performing chemical precipitation, filtration, weak acid cation exchange and ozone biological activated carbon adsorption oxidation on wastewater generated in a molecular sieve production process to remove impurities, and performing ED membrane concentration and reverse osmosis concentration on the wastewater.

The conditions for the above chemical precipitation are preferably: sodium carbonate is used as a treating agent, 1.2-1.4mol of sodium carbonate is added relative to 1mol of calcium ions in the wastewater, the pH value of the wastewater is adjusted to be more than 7, the reaction temperature is 20-35 ℃, and the reaction time is 0.5-4h.

The conditions for the filtration are preferably: the filtering unit adopts a double-layer filtering material multi-medium filter consisting of anthracite and quartz sand, the grain diameter of the anthracite is 0.7-1.7mm, the grain diameter of the quartz sand is 0.5-1.3mm, and the filtering speed is 10-30m/h. After the filter material is used, the regeneration method of 'gas back flushing-gas and water back flushing-water back flushing' is adopted to regenerate the filter material, and the regeneration period is 10-15h.

The conditions for the weak acid cation exchange method are preferably: the pH value range is 6.5-7.5; the temperature is less than or equal to 40 ℃, the height of the resin layer is 1.5-3.0m, the HCl concentration of the regeneration liquid is as follows: 4.5-5 mass%; the dosage of the regenerant (calculated by 100%) is 50-60kg/m ³ Wet resin; the flow rate of the regeneration liquid HCl is 4.5-5.5m/h, and the regeneration contact time is 35-45min; the forward washing flow rate is 18-22m/h, and the forward washing time is 2-30min; the running flow rate is 15-30m/h; as the acidic cation exchange resin, for example, there can be used a Gallery Senno chemical Co., ltd, SNT brand D113 acidic cation exchange resin.

The conditions of the above-mentioned ozone biological activated carbon adsorption oxidation method are preferably: the retention time of the ozone is 50-70min, and the empty bed filtration rate is 0.5-0.7m/h.

The conditions for the concentration of the ED membrane are preferably: the current is 145-155A, and the voltage is 45-65V. As the ED membrane, for example, an ED membrane manufactured by easton corporation, japan can be used.

The conditions for the reverse osmosis are preferably: the operation pressure is 5.4-5.6MPa, the water inlet temperature is 25-35 ℃, and the pH value is 6.5-7.5. The reverse osmosis membrane is, for example, a seawater desalination membrane TM810C manufactured by Dongli corporation of Lanxingdong.

According to the invention, when the wastewater treatment is started, the catalyst production wastewater can be used for directly starting operation, and if the ion content of the catalyst production wastewater meets the conditions of the invention, the first evaporation, the cooling crystallization and the second evaporation can be carried out according to the conditions of the invention; if the ion content of the catalyst production wastewater does not meet the conditions of the invention, second evaporation can be carried out to obtain a second concentrated solution, solid-liquid separation is carried out to obtain sodium chloride crystals and a second mother liquor, then the second mother liquor and the catalyst production wastewater are mixed to adjust the ion content of the wastewater to be treated to be in the range required by the invention, and then cooling crystallization is carried out to obtain sodium sulfate crystals. Of course, it is also possible to adjust the ion content of the wastewater to be treated in the initial stage by using sodium sulfate or sodium chloride, provided that the wastewater to be treated satisfies the SO content of the wastewater to be treated in the present invention ₄ ^2- 、Cl ^- The requirements are met.

The present invention will be described in detail below by way of examples.

In the following examples, the catalyst production wastewater is wastewater from a molecular sieve production process, which is subjected to chemical precipitation, filtration, weak acid cation exchange and ozone biological activated carbon adsorption oxidation in sequence to remove impurities, and is subjected to ED membrane concentration and reverse osmosis concentration in sequence.

Example 1

As shown in figure 1, the catalyst production wastewater (containing NaCl 38g/L, na) ₂ SO ₄ 70g/L、NH ₄ Cl 33g/L、(NH ₄ ) ₂ SO ₄ 61.8g/L, pH of 7.2) at a feed rate of 10m ³ Delivery rate of/hIntroducing a sodium hydroxide aqueous solution with the concentration of 45.16 mass% into a pipeline of a treatment system for carrying out primary pH value adjustment, monitoring the adjusted pH value through a first pH value measuring device 61 (a pH meter) (the measured value is 7.6), then sending one part of the catalyst production wastewater into a first heat exchange device 31 (a duplex stainless steel plate heat exchanger) for heat exchange with a first ammonia-containing steam condensate, mixing with a second mother liquor returned by a ninth circulating pump 79, sending the other part of the catalyst production wastewater into an eleventh heat exchange device 30 (a duplex stainless steel plate heat exchanger) for heat exchange with the first condensate, combining the two parts of wastewater to obtain wastewater to be treated, sending the wastewater to be treated into an eighth heat exchange device 38 (the duplex stainless steel plate heat exchanger) for heat exchange with first ammonia-containing steam to heat the wastewater to be treated to 102 ℃, then sending the wastewater to a pipeline of a second MVR evaporation device 3, introducing a sodium hydroxide aqueous solution with the concentration of 45.16 mass% for carrying out secondary pH value adjustment, and monitoring the pH value through a third pH value measuring device 60 (a pH meter) for monitoring the measured value of 11).

The first evaporation is carried out in a first MVR evaporation plant 3 at a temperature of 100 deg.C under a pressure of-22.82 kPa with an evaporation capacity of 3.13m ³ And/h, obtaining first ammonia-containing steam and first concentrated solution. After the first ammonia-containing steam is compressed by the first compressor 101 (the temperature is raised by 18 ℃), the first ammonia-containing steam exchanges heat with the wastewater to be treated and the catalyst production wastewater respectively in the eighth heat exchange device 38 and the first heat exchange device 31 in sequence to obtain first ammonia water, and the first ammonia water is stored in the first ammonia water storage tank 51. In addition, in order to increase the liquid concentration in the first MVR evaporation device 3, part of the liquid evaporated in the first MVR evaporation device 3 is circulated as a first circulation liquid to the eighth heat exchange device 38 by the fifth circulation pump 75, and then enters the first MVR evaporation device 3 again for first evaporation (the first reflux ratio is 27.2). Controlling the flow rate of the eleventh circulating pump 70, controlling the degree of the first evaporation by a densimeter arranged on the first MVR evaporation device 3, controlling the concentration of the first concentrated solution, and measuring the concentration of NaCl in the obtained first concentrated solution to be 135g/L (2.307 mol/L) and Na ₂ SO ₄ The concentration of (2) was 177.4g/L (1.249 mol/L).

Feeding the first concentrated solution intoThe eleventh heat exchange device 30 exchanges heat with the catalyst production wastewater, then the cooled first concentrated solution is sent to the second heat exchange device 32 (a heat exchanger made of plastic materials) to carry out second heat exchange with the first mother solution so as to cool the first concentrated solution to 19.2 ℃, then the cooled first concentrated solution is mixed with the cooling circulating solution of the cooling crystallization device 2 conveyed by the second circulating pump 72, and the cooled first concentrated solution is further cooled after exchanging heat with the refrigerating solution in the sixth heat exchange device 36, and then the cooled first concentrated solution is sent to the cooling crystallization device 2 (a freezing crystallization tank) to carry out cooling crystallization, so that the crystallization solution containing sodium sulfate crystals is obtained. Wherein the cooling crystallization temperature is-4 deg.C, the time is 120min, and the circulation amount of cooling crystallization is controlled to be 1176m ³ And h, controlling the supersaturation degree of sodium sulfate in the cooling crystallization process to be not more than 1.1g/L.

The sodium sulfate crystal-containing crystal liquid obtained in the cooling crystallization device 2 was sent to a first solid-liquid separation device 91 (centrifuge) to carry out solid-liquid separation, and 3.74m per hour was obtained ³ Contains NaCl 288.4g/L, na ₂ SO ₄ 14.6g/L、NaOH 4.7g/L、NH ₃ 0.41g/L of the first mother liquor was temporarily stored in the first mother liquor tank 53, and sodium sulfate decahydrate crystal cake 5555.76kg having a purity of 98.2 mass% and a water content of 75 mass% was obtained per hour.

The second evaporation process is carried out in a second MVR evaporation plant 1 (falling film + forced circulation two-stage MVR evaporator crystallizer). Sending the first mother liquor into a second heat exchange device 32 to exchange heat with the first concentrated solution, then sending the first mother liquor into a third heat exchange device 33 (a duplex stainless steel plate type heat exchanger) to exchange heat with the compressed second ammonia-containing steam condensate, then sending the first mother liquor into a fourth heat exchange device 34 (a duplex stainless steel plate type heat exchanger) to exchange heat with the second ammonia-containing steam, monitoring the pH value through a second pH value measuring device 62 (a pH meter) (the measured value is 11.3), and carrying out second evaporation in a second MVR evaporation device 1 to obtain a second concentrated solution containing sodium chloride crystals and second ammonia-containing steam. The second evaporation temperature is 100 deg.C, pressure is-22.83 kPa, and evaporation capacity is 2.74m ³ H is used as the reference value. After the second ammonia-containing steam is compressed by the second compressor 102 (the temperature rises by 17 ℃), the second ammonia-containing steam exchanges heat with the first mother liquor in the fourth heat exchange device 34 and the third heat exchange device 33 in sequence to obtain second ammonia water, and the second ammonia water is stored in the second ammonia water storage tank 52. In addition, in order to improve the second MVR evaporation device 1The second evaporated part of the liquid in the second MVR evaporation device 1 is sent again to the second MVR evaporation device 1 as a circulating liquid through the seventh circulating pump 77 for second evaporation (the second reflux ratio is 72.5). The degree of the second evaporation was monitored by a densimeter provided in the MVR evaporation apparatus 1, and the concentration of sodium sulfate in the second concentrated solution after the second evaporation was controlled to be 0.9603Y (48.4 g/L).

Sending the second concentrated solution containing sodium chloride crystals into a second solid-liquid separation device 92 (centrifugal machine) for second solid-liquid separation and leaching to obtain 1.13m per hour ³ Contains NaCl 300.2g/L, na ₂ SO ₄ 48.4g/L、NaOH 15.58g/L、NH ₃ 0.00036g/L of second mother liquor is temporarily stored in the second mother liquor tank 54 and returns to the wastewater conveying pipeline through a ninth circulating pump 79 to be mixed with the catalyst production wastewater to obtain wastewater to be treated. The obtained sodium chloride solid (sodium chloride crystal cake with a water content of 14 mass% 833.31kg, wherein the sodium sulfate content is 1.5 mass% or less, obtained per hour) was subjected to spray washing with 300g/L sodium chloride solution equivalent to the dry basis mass of sodium chloride, and then dried in a dryer to obtain 716.64kg (purity of 99.5 mass%) per hour, and the washed second washing liquid was circulated to the second MVR evaporation apparatus 1 by the tenth circulation pump 80.

In this example, 3.13m of aqueous ammonia having a concentration of 7.0 mass% was obtained per hour in the first aqueous ammonia tank 51 ³ (ii) a 2.74m of 0.05 mass% ammonia water was obtained per hour in the second ammonia water tank 52 ³ 。

In addition, the tail gas discharged by the eighth heat exchange device 38, the cooling crystallization device 2 and the fourth heat exchange device 34 is introduced into a tail gas absorption tower 83 through a vacuum pump 81 for absorption, circulating water is introduced into the tail gas absorption tower 83, the circulating water circulates in the tail gas absorption tower 83 under the action of the fourth circulating pump 74, water is supplemented into the tail gas absorption tower 83 from a circulating water tank 82 through a third circulating pump 73, and fresh water is supplemented into the circulating water tank 82, so that the temperature and the ammonia content of the working water of the vacuum pump 81 are reduced. Dilute sulfuric acid is further introduced into the tail gas absorption tower 83 to absorb ammonia and the like in the tail gas. The starting phase of MVR evaporation was initiated by steam at a temperature of 143.3 ℃.

Example 2

The treatment of the catalyst production wastewater was carried out in the same manner as in example 1 except that: for the sample containing NaCl 65g/L, na ₂ SO ₄ 66g/L、NH ₄ Cl 29g/L、(NH ₄ ) ₂ SO ₄ 29.9g/L, pH is 6.8, and the temperature of the wastewater to be treated after heat exchange by the eighth heat exchange device 38 is 114 ℃.

The temperature of the first evaporation is 107 ℃, the pressure is 0Pa, and the evaporation capacity is 3.45m ³ H; cooling and crystallizing at-2 deg.C for 120min; the temperature of the second evaporation is 75 ℃, the pressure is-72.75 Pa, and the evaporation capacity is 3.48m ³ H is used as the reference value. The first concentrated solution obtained contained 174.8g/L of NaCl and Na ₂ SO ₄ 132.6g/L (i.e., cl) ^- Has a concentration of 2.987mol/L, SO ₄ ^2- The concentration of (2) is 0.933 mol/L).

The first solid-liquid separation device 91 obtains 76 mass% of water-containing sodium sulfate decahydrate crystal cake 4150.80kg with a purity of 98.1 mass%) per hour; yield 4.77m per hour ³ The concentration of NaCl is 292.8g/L, na ₂ SO ₄ 16.5g/L、NaOH 4.4g/L、NH ₃ 0.2g/L of the first mother liquor.

The second solid-liquid separation device 92 obtains 1121.10kg of sodium chloride crystal filter cake with the water content of 15 mass% per hour, and finally obtains 952.93kg of sodium chloride (the purity is 99.5 mass%) per hour; 1.45 m/hr ³ The concentration of NaCl 296.6g/L, na ₂ SO ₄ 54.2g/L、NaOH 0.36g/L、NH ₃ 0.00014g/L of second mother liquor.

3.45m of ammonia water having a concentration of 4.2 mass% is obtained per hour in the first ammonia water tank 51 ³ (ii) a The second ammonia water tank 52 was filled with 3.48m of 0.02 mass% ammonia water per hour ³ 。

Example 3

The treatment of the catalyst production wastewater was carried out in the same manner as in example 1, except that: for NaCl-containing 88g/L, na ₂ SO ₄ 50g/L、NH ₄ Cl 38g/L、(NH ₄ ) ₂ SO ₄ 21.9g/L, pH of 6.6The catalyst production wastewater is treated, and the temperature of the wastewater to be treated after heat exchange is carried out by the eighth heat exchange device 38 is 112 ℃.

The temperature of the first evaporation was 105 deg.C, the pressure was-7.01 Pa, and the evaporation capacity was 3.59m ³ H; cooling and crystallizing at 0 deg.C for 120min; the second evaporation temperature is 50 deg.C, pressure is-92.67 Pa, and evaporation capacity is 4.40m ³ H is used as the reference value. The first concentrated solution thus obtained contained 219.0g/L of NaCl, na ₂ SO ₄ 104.4g/L (i.e., cl) ^- Has a concentration of 3.744mol/L, SO ₄ ^2- The concentration of (b) is 0.735 mol/L).

The first solid-liquid separation device 91 obtains 74.5 mass% of water-containing sodium sulfate decahydrate crystal cake 2914.06kg with a purity of 98.2 mass%) per hour; 5.76 m/hr ³ The concentration of NaCl is 304.1g/L, na ₂ SO ₄ 17.2g/L、NaOH 3.6g/L、NH ₃ 0.17g/L of the first mother liquor.

The second solid-liquid separation device 92 obtains 1515.97kg of sodium chloride crystal cake with the water content of 15 mass% per hour, and finally obtains 1288.58kg of sodium chloride (the purity is 99.5 mass%) per hour; 1.59 m/h ³ The concentration of NaCl 287g/L, na ₂ SO ₄ 62.3g/L、NaOH 13.1g/L、NH ₃ 0.0001g/L of second mother liquor.

3.59m of ammonia water having a concentration of 4.2 mass% was obtained per hour in the first ammonia water tank 51 ³ (ii) a 4.40m of ammonia water having a concentration of 0.02 mass% was obtained in the second ammonia water tank 52 every hour ³ 。

The preferred embodiments of the present invention have been described above in detail, but the present invention is not limited thereto. Within the scope of the technical idea of the invention, many simple modifications can be made to the technical solution of the invention, including combinations of various technical features in any other suitable way, and these simple modifications and combinations should also be regarded as the disclosure of the invention, and all fall within the scope of the invention.

Claims

1. Method for treating wastewater generated in catalyst production, wherein the wastewater generated in catalyst production contains NH ₄ ⁺ 、SO ₄ ^2- 、Cl ^- And Na ⁺ Characterized in that the method comprises the following steps,

1) Performing first evaporation on wastewater to be treated to obtain first ammonia-containing steam and first concentrated solution;

before the wastewater to be treated is subjected to first evaporation, adjusting the pH value of the wastewater to be treated to be more than 9;

SO in the first concentrated solution ₄ ^2- The concentration of (A) is more than 0.01 mol/L;

the second evaporation does not crystallize sodium sulfate out;

the wastewater to be treated contains the catalyst production wastewater and a liquid phase obtained by the second solid-liquid separation; NH in the catalyst production wastewater ₄ ⁺ Is more than 8mg/L, SO ₄ ^2- Is more than 1g/L, cl ^- Over 970mg/L of Na ⁺ Is more than 510 mg/L.

2. The method of claim 1, wherein the SO contained in the first concentrate is ₄ ^2- The concentration of (B) is 0.1mol/L or more.

3. The method of claim 2, wherein the first concentrate contains SO ₄ ^2- The concentration of (B) is 0.2mol/L or more.

4. The method of claim 1, wherein the first concentrate comprises Cl ^- The concentration of (B) is 5mol/L or less.

5. The method of claim 4, wherein the first concentrate comprises Cl ^- The concentration of (b) is 4.5mol/L or less.

6. The method according to claim 1, wherein the SO contained in the liquid phase obtained by the first solid-liquid separation is 1mol relative to the SO contained in the liquid phase ₄ ^2- Cl contained in the liquid phase obtained by the first solid-liquid separation ^- 9.5mol or more.

7. The method according to claim 1, wherein the pH of the wastewater to be treated is adjusted to be greater than 10.8 before the wastewater to be treated is subjected to the first evaporation.

8. The method of claim 1, wherein adjusting the pH is performed with NaOH.

9. The method of claim 1, wherein the SO in the first concentrated solution is present prior to the cooling crystallization ₄ ^2- Has a concentration of 0.01mol/L or more and Cl ^- The concentration of (b) is 5.2mol/L or less.

10. The method of claim 1, wherein the cooling crystallization does not crystallize sodium chloride.

11. The process of claim 1, wherein the concentration of sodium chloride in the first concentrate is adjusted prior to the cooling crystallization.

12. The process according to claim 1, wherein the adjustment of the concentration of sodium chloride in the first concentrated solution is performed by mixing a liquid phase obtained by the second solid-liquid separation, the catalyst production wastewater and/or a washing liquid for washing sodium sulfate crystals.

13. The process of any one of claims 1-12, wherein the second evaporation results in a sodium sulfate concentration in the second concentrated solution that is less than Y, where Y is the sodium sulfate concentration at which both sodium sulfate and sodium chloride are saturated in the second concentrated solution under the conditions of the second evaporation.

14. The process of claim 13, wherein the second evaporation provides a sodium sulfate concentration in the second concentrated solution of 0.9Y to 0.99Y.

15. The method of any one of claims 1-12, wherein the conditions of the first evaporation comprise: the temperature is above 35 ℃ and the pressure is above-98 kPa.

16. The method of claim 15, wherein the conditions of the first evaporation comprise: the temperature is 75-130 ℃, and the pressure is-73 kPa-117 kPa.

17. The method of claim 16, wherein the conditions of the first evaporation comprise: the temperature is 85-130 ℃, and the pressure is-58-117 kPa.

18. The method of claim 17, wherein the conditions of the first evaporation comprise: the temperature is 95-110 ℃, and the pressure is-37 kPa-12 kPa.

19. The method of claim 15, wherein the first evaporation is performed using an MVR evaporation device.

20. The process according to any one of claims 1 to 12, wherein the temperature of the cooling crystallization is from-21.7 ℃ to 17.5 ℃.

21. The method according to claim 20, wherein the temperature of the cooling crystallization is from-20 ℃ to 5 ℃.

22. The method of claim 21, wherein the temperature of the cooling crystallization is from-10 ℃ to 5 ℃.

23. The method of claim 22, wherein the temperature of the cooling crystallization is from-10 ℃ to 0 ℃.

24. The method according to claim 20, wherein the cooling crystallization time is 5min or more.

25. The method according to claim 24, wherein the cooling crystallization time is 60min to 180min.

26. The method of claim 25, wherein the cooling crystallization time is 90min to 150min.

27. The method of any one of claims 1-12, wherein the conditions of the second evaporation comprise: the temperature is above 17.5 ℃ and the pressure is above-101 kPa.

28. The method of claim 27, wherein the conditions of evaporation comprise: the temperature is 35-110 ℃, and the pressure is-98 kPa-12 kPa.

29. The method of claim 28, wherein the conditions of evaporation comprise: the temperature is 45-110 ℃, and the pressure is-95 kPa-12 kPa.

30. The method of claim 29, wherein the conditions of evaporation comprise: the temperature is 50-100 ℃, and the pressure is-93 kPa to-22 kPa.

31. The method according to claim 1, wherein the wastewater to be treated is subjected to a first heat exchange with the first ammonia-containing steam and a first ammonia water is obtained before the wastewater to be treated is subjected to a first evaporation.

32. The process as claimed in claim 31, wherein the first concentrated solution is subjected to a second heat exchange with a liquid phase obtained by the first solid-liquid separation before the first concentrated solution is subjected to cooling crystallization.

33. The process of claim 31, wherein the second ammonia-containing vapor is subjected to a third heat exchange with the first solid-liquid separated liquid phase and a second aqua ammonia is obtained before passing the first solid-liquid separated liquid phase to a second MVR evaporation unit.

34. The method according to any one of claims 1 to 12, further comprising subjecting the sodium sulfate crystal-containing crystal liquid to a first solid-liquid separation to obtain sodium sulfate crystals.

35. The method of claim 34, further comprising washing the resulting sodium sulfate crystals.

36. The method according to any one of claims 1 to 12, further comprising subjecting the second concentrated solution containing sodium chloride crystals to a second solid-liquid separation to obtain sodium chloride crystals.

37. The method of claim 36, further comprising washing the obtained sodium chloride crystals.

38. The process of any of claims 1-12, wherein the catalyst production wastewater is wastewater from a molecular sieve, alumina, or refinery catalyst production process.

39. The method of claim 38, further comprising removing impurities and concentrating the catalyst process wastewater.