CN108726755B

CN108726755B - Treatment method of catalyst production wastewater

Info

Publication number: CN108726755B
Application number: CN201710263235.3A
Authority: CN
Inventors: 殷喜平; 李叶; 苑志伟
Original assignee: China Petroleum and Chemical Corp; Sinopec Catalyst Co
Current assignee: China Petroleum and Chemical Corp; Sinopec Catalyst Co
Priority date: 2017-04-21
Filing date: 2017-04-21
Publication date: 2021-09-21
Anticipated expiration: 2037-04-21
Also published as: CN108726755A

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 NH₄ ⁺、SO₄ ^2‑、Cl^‑And Na⁺The method comprises the following steps of 1) introducing wastewater to be treated into a first MVR evaporation device for first evaporation to obtain first concentrated solution containing ammonia vapor and sodium sulfate crystals, wherein the wastewater to be treated contains the catalyst production wastewater; 2) carrying out first solid-liquid separation on the first concentrated solution 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 the sodium chloride crystals; 3) 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

Treatment method of 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 NH₄ ⁺、SO₄ ^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 method 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 a biochemical method, a blow-off method or a 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 of 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 multiple-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 problem is 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), the wastewater needs to be deeply treated, the salt content of the wastewater is not reduced (20000 mg/L-30000 mg/L), the wastewater cannot be directly discharged, and the wastewater needs to be further desalted.

In order to remove ammonia nitrogen from wastewater by gas stripping deamination, a large amount of alkali is needed to adjust the pH value, the alkali consumption is high, the alkali in the wastewater after deamination 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 does not change greatly, the salt content in the wastewater is not reduced (20000 mg/L-30000 mg/L), the wastewater cannot be discharged directly, further desalting treatment is needed, the operation cost of wastewater treatment 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 art₄ ⁺、SO₄ ^2-、Cl^-And Na⁺The catalyst has high treatment cost of wastewater and can only obtain mixed salt crystals, and provides a low-cost and environment-friendly NH-containing catalyst₄ ⁺、SO₄ ^2-、Cl^-And Na⁺The method for treating the catalyst production wastewater can respectively recover ammonium, sodium sulfate and sodium chloride in the catalyst production wastewater, and furthest recycle resources in the catalyst production wastewater.

In order to achieve the above object, the present invention provides a method for treating wastewater from catalyst production containing NH₄ ⁺、SO₄ ^2-、Cl^-And Na⁺The method comprises the following steps of,

1) introducing wastewater to be treated into a first MVR evaporation device for first evaporation to obtain first concentrated solution containing ammonia vapor and sodium sulfate crystals, wherein the wastewater to be treated contains the catalyst production wastewater;

2) carrying out first solid-liquid separation on the first concentrated solution 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 the sodium chloride crystals;

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

adjusting the pH value of the wastewater to be treated to be more than 9 before introducing the wastewater to be treated into a first MVR evaporation device; the first evaporation prevents sodium chloride from crystallizing out, and the second evaporation prevents sodium sulfate from crystallizing out; relative to 1 mole of SO contained in the wastewater to be treated₄ ^2-Cl contained in the wastewater to be treated^-The molar ratio is 14 mol or less.

By the technical scheme, the method aims at the content of NH₄ ⁺、SO₄ ^2-、Cl^-And Na⁺The pH value of the wastewater to be treated is adjusted to a specific range in advance, then a first MVR evaporation device is used for evaporation and separation to obtain sodium sulfate crystals and concentrated ammonia water, and then a second MVR evaporation device is used for evaporation again to obtain sodium chloride crystals and dilute ammonia water. 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.

Additional features and advantages of the invention will be set forth in the detailed description which follows.

Drawings

The accompanying drawings, which are included to provide a further understanding of the invention and are incorporated in and constitute a part of this specification, illustrate embodiments of the invention and together with the description serve to explain the principles of the invention and not to limit the invention. In the drawings:

FIG. 1 is a schematic flow diagram of a wastewater treatment process according to one embodiment of the present invention.

Description of the reference numerals

1. Second MVR evaporation plant 2, first MVR evaporation plant

31. First heat exchange device 32 and second heat exchange device

33. Third heat exchange device 34 and fourth heat exchange device

4. Vacuum degassing tank 51, first aqueous ammonia storage tank

52. Second ammonia storage tank 53, first mother liquor jar

54. Second mother liquor tank 61 and first pH value measuring device

62. Second pH value measuring device 71 and first circulating pump

72. Second circulating pump 73 and third circulating pump

74. Fourth and fifth circulating pumps 75 and 75

76. Sixth circulating pump 77 and seventh circulating pump

78. Eighth circulating pump 79 and ninth circulating pump

81. Vacuum pump 82 and circulating water pool

83. Tail gas absorption tower 91 and first solid-liquid separation device

92. Second solid-liquid separation device 101 and first compressor

102. Second compressor

Detailed Description

The following describes in detail specific embodiments of the present invention. It should be understood that the detailed description and specific examples, while indicating the present invention, are given by way of illustration and explanation only, not limitation.

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 NH₄ ⁺、SO₄ ^2-、Cl^-And Na⁺The method comprises the following steps of,

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.

Further preferably, the wastewater to be treated is a liquid phase mixed solution obtained by the catalyst production wastewater and the second solid-liquid separation.

Preferably, the pH of the wastewater to be treated is adjusted to be greater than 10.8 before it is passed into the first MVR evaporation plant 2. The upper limit of the pH of the wastewater to be treated is not limited, and may be, for example, 14 or less, preferably 13.5 or less, and more preferably 13 or less.

The method provided by the invention can be used for the treatment of the compounds containing NH₄ ⁺、SO₄ ^2-、Cl^-And Na⁺Is treated except for containing NH₄ ⁺、SO₄ ^2-、Cl^-And Na⁺In addition, the wastewater to be treated is not particularly limited. From the viewpoint of improving the treatment efficiency of wastewater, the amount of SO contained in the wastewater to be treated is 1 mole per mole₄ ^2-Cl contained in the wastewater to be treated^-13.8 moles or less, more preferably 13.75 moles or less, further preferably 13.5 moles or less, further preferably 13 moles or less, further preferably 12 moles or less, further preferably 11 moles or less, further preferably 10 moles or less, further preferably 9 moles or less, further preferably 8 moles or less, further preferably 7 moles or less; preferably 2 moles or more, more preferably 2.5 moles or more, further preferably 3 moles or more, and for example, may be 4 to 10 moles. By reacting SO₄ ^2-And Cl^-The molar ratio of (b) is controlled within the above range and the evaporation conditions described later are combined, so that sodium sulfate is precipitated in the first evaporation without precipitating sodium chloride, thereby achieving the purpose of efficiently separating sodium sulfate. In addition, as described above and below, it is also possible in the present invention to recycle the second mother liquor obtained in the second evaporation process to the first evaporation, and thereby to treat SO in the wastewater to be treated₄ ^2-And Cl^-The molar ratio of (a) to (b) is adjusted and the balance of sodium hydroxide can be maintained.

In the present invention, the first evaporation to prevent the crystallization of sodium chloride means that the concentration of sodium chloride in the mixed system is controlled not to exceed the solubility under the first evaporation conditions (including but not limited to temperature, pH, etc.), but sodium chloride carried by sodium sulfate crystals or adsorbed on the surface is not excluded. Since the water content of the crystals after the solid-liquid separation is different, the sodium chloride content in the obtained sodium sulfate crystals is usually 8 mass% or less (preferably 4 mass% or less), and in the present invention, it is considered that sodium chloride is not crystallized when the sodium chloride content in the obtained sodium sulfate crystals is 8 mass% or less.

In the present invention, the second evaporation to prevent the crystallization of sodium sulfate means that the concentration of sodium sulfate in the mixed system is controlled not to exceed the solubility under the second evaporation conditions (including but not limited to temperature, pH, etc.), and sodium sulfate entrained by sodium chloride crystals or adsorbed on the surface is not excluded. Since the water content of the crystals after the solid-liquid separation is different, the sodium sulfate content in the obtained sodium chloride crystals is usually 8 mass% or less (preferably 4 mass% or less), and in the present invention, it is considered that the sodium sulfate does not crystallize out when the sodium sulfate content in the obtained sodium chloride crystals is 8 mass% or less.

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

In the present invention, the first MVR evaporation device 2 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 evaporative crystallizer.

In the present invention, the conditions of the first evaporation may be appropriately selected as needed, and sodium sulfate may be crystallized without precipitating sodium chloride. The conditions of the first evaporation may include: the temperature is above 45 ℃ and the pressure is above-95 kPa. In order to improve the evaporation efficiency, it is preferable that the conditions of the first evaporation include: the temperature is 45-365 ℃, and the pressure is-95 kPa-18110 kPa; to further improve the evaporation efficiency, more preferably, the conditions of the first evaporation include: the temperature is 60-365 ℃, and the pressure is-87 kPa-18110 kPa; more preferably, from the viewpoint of reducing equipment cost and energy consumption, the conditions of the first evaporation include: the temperature is 75-175 ℃, and the pressure is-73 kPa-653 kPa; further preferably, the conditions of the first evaporation include: the temperature is 80-130 ℃, and the pressure is-66 kPa-117 kPa; particularly preferably, the conditions of the first evaporation include: the temperature is 95-105 ℃, and the pressure is-37 kPa to-7 kPa.

In the present invention, the operating pressure of the first evaporation is more preferably the saturated vapor pressure of the evaporated feed liquid.

In the present invention, the flow rate of the first evaporation may be appropriately selected according to the capacity of the apparatus process, and may be, for example, 0.1m³More than h (e.g. 0.1 m)³/h～500m³/h)。

By carrying out the first evaporation under the above conditions, sodium chloride is not crystallized and precipitated while sodium sulfate is crystallized, and the purity of the obtained sodium sulfate crystal can be ensured.

According to the invention, by controlling the evaporation conditions of the first MVR evaporation device 2, more than 90 mass% (preferably more than 95 mass%) of ammonia contained in the wastewater to be treated can be evaporated, and the first ammonia water can be directly reused in the production process of the catalyst, or neutralized by acid to obtain ammonium salt for reuse, or prepared with water and corresponding ammonium salt or ammonia water for use.

According to the present invention, the first evaporation does not crystallize sodium chloride in the wastewater to be treated (i.e., sodium chloride does not reach supersaturation), and preferably, the first evaporation is performed so that the concentration of sodium chloride in the first concentrated solution is X or less (preferably 0.999X or less, more preferably 0.95X to 0.999X, and further preferably 0.99X to 0.9967X). Wherein, X is the concentration of sodium chloride when the sodium sulfate and the sodium chloride in the first concentrated solution reach saturation under the condition of the first evaporation. By controlling the degree of the first evaporation within the above range, as much sodium sulfate as possible can be crystallized under the condition that sodium chloride is not precipitated. By crystallizing sodium sulfate in the first evaporation as much as possible, the wastewater treatment efficiency can be improved, and the energy waste can be reduced.

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

In the 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 (the preparation of the wastewater to be treated is required in the case that the wastewater to be treated contains a liquid phase obtained by the separation of the catalyst production wastewater and the second solid-liquid) is not particularly limited, and 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 can be completed before the wastewater to be treated is introduced into the MVR evaporation plant by appropriately selecting according to the needs.

According to a preferred embodiment of the present invention, before the wastewater to be treated is introduced into the first MVR evaporation device 2, the first ammonia-containing steam and the wastewater to be treated are subjected to a first heat exchange to obtain a first ammonia water. The first heat exchange method is not particularly limited, and may be performed by a heat exchange method that is conventional in the art. The number of the first heat exchange may be 1 or more, preferably 2 to 4, and more preferably 2 to 3. Through after the first heat exchange, the output ammonia water is cooled, and the heat is circulated in the treatment device to the maximum extent, so that the energy is reasonably utilized, and the waste is reduced.

According to a preferred embodiment of the present invention, the first heat exchange is performed by a first heat exchange device 31 and a second heat exchange device 32, specifically, first ammonia-containing steam obtained by evaporation in the first MVR evaporation device 2 passes through the second heat exchange device 32 and the first heat exchange device 31 in sequence, and the wastewater to be treated passes through the first heat exchange device 31 and the second heat exchange device 32 in sequence, and the first heat exchange is performed between the first ammonia-containing steam and the wastewater to be treated, so that the temperature of the wastewater to be treated is raised for evaporation, and the temperature of the first ammonia-containing steam is lowered to obtain first ammonia water, which can be stored in the first ammonia water storage tank 51.

In the present invention, the first heat exchange device 31 and the second heat exchange device 32 are not particularly limited, and various heat exchangers conventionally used in the art may be used to achieve the first heat exchange between the first ammonia-containing steam and the wastewater to be treated. 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.

According to a preferred embodiment of the present invention, the first ammonia-containing vapor condensate is subjected to a first heat exchange with the wastewater to be treated in a first heat exchange means 31, and the first ammonia-containing vapor is subjected to a first heat exchange with the wastewater to be treated in a second heat exchange means 32.

According to the present invention, in order to fully utilize the heat energy of the first ammonia-containing vapor condensate, it is preferable that the temperature of the wastewater to be treated is 40 to 364 ℃, more preferably 55 to 364 ℃, even more preferably 65 to 174 ℃, and even more preferably 79 to 129 ℃ after the first heat exchange is performed by the first heat exchange device 31.

According to the present invention, in order to fully utilize the heat energy of the first ammonia-containing steam, it is preferable that the temperature of the wastewater to be treated is 50 ℃ to 370 ℃, more preferably 65 ℃ to 370 ℃, even more preferably 75 ℃ to 184 ℃, and even more preferably 85 ℃ to 139 ℃ after the first heat exchange is performed by the second heat exchange device 32.

In the present invention, the method of adjusting the pH 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 may be, for example, a hydroxide such as sodium hydroxide or potassium hydroxide for the purpose of adjusting the pH. The alkaline substance is preferably NaOH in order not to introduce new impurities in the wastewater to be treated, increasing the purity of the crystals obtained.

The manner of adding the alkaline substance may be any manner known in the art, but it is preferable to mix the alkaline substance with the wastewater to be treated in the form of an aqueous solution, and for example, an aqueous solution containing the alkaline 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. 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 process is performed in the first MVR evaporation plant 2, and the first pH adjustment is performed by introducing and mixing the aqueous solution containing the basic substance in the pipe for feeding the catalyst production wastewater to the first heat exchange device 31 before feeding the catalyst production wastewater to the first heat exchange device 31 for the first heat exchange; and then mixing the catalyst production wastewater with at least part of a liquid phase obtained by second solid-liquid separation to obtain wastewater to be treated, sending the wastewater to be treated into a second heat exchange device 32 for first heat exchange, and introducing the aqueous solution containing the alkaline substance into a pipeline for sending the wastewater to be treated into the second heat exchange device 32 and mixing to perform second pH value adjustment. The pH of the wastewater to be treated is adjusted twice, so that the pH is greater than 9, preferably greater than 10.8, before the wastewater is passed into the first MVR evaporator 2. Preferably, the first pH adjustment is carried out so that the pH value of the adjusted wastewater is more than 7 (preferably 7-9), and the second pH adjustment is carried out so that the pH value of the wastewater to be treated is more than 9, preferably more than 10.8.

In order to detect the pH values after the first pH adjustment and the second pH adjustment, it is preferable that a first pH measuring device 61 is provided on a pipe for feeding the catalyst production wastewater to the first heat exchange device 31 to measure the pH value after the first pH adjustment, and a second pH measuring device 62 is provided on a pipe for feeding the wastewater to be treated to the second heat exchange device 32 to measure the pH value after the second pH adjustment.

In the present invention, in order to increase the solid content in the first MVR evaporation device 2 and reduce the ammonia content in the liquid, it is preferable that a part of the liquid evaporated by the first MVR evaporation device 2 (i.e. the liquid located inside the first MVR evaporation device, hereinafter also referred to as a first circulation liquid) is heated and then returned to the first MVR evaporation device 2 for evaporation. The above-mentioned process of returning part of the liquid evaporated by the first MVR evaporation device 2 to the first MVR evaporation device 2 is preferably that the first circulation liquid is mixed with the wastewater to be treated after the first pH adjustment and before the second pH adjustment and then fed into the first MVR evaporation device 2, for example, the first circulation liquid may be returned to the wastewater conveying pipeline between the first heat exchange device 31 and the second heat exchange device 32 by the second circulation pump 72 to be mixed with the wastewater to be treated, and then after the second pH adjustment, the heat exchange is performed in the second heat exchange device 32, and finally the wastewater is fed into the first MVR evaporation device 2. As a ratio of returning a part of the liquid evaporated by the first MVR evaporation device 2 to the MVR evaporation device 2, there is no particular limitation, and for example, the first return ratio of the first evaporation may be 10 to 200, preferably 50 to 100. 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 2 minus the amount of reflux.

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 intensification-evaporation-cooling's process goes on in succession, need input start-up steam when MVR vaporization process starts, only need pass 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 compression by the first compressor 101, the temperature of the first ammonia-containing vapor is raised by 5 to 20 ℃.

In the invention, the first concentrated solution containing sodium sulfate crystals is subjected to first solid-liquid separation to obtain sodium sulfate crystals and a first mother liquor (namely, a liquid phase obtained by the first solid-liquid separation). The method of the first solid-liquid separation is not particularly limited, and may be selected from one or more of centrifugation, filtration, and sedimentation.

According to the present invention, the solid-liquid separation of the first concentrated solution may be performed by using a first solid-liquid separation device (for example, a centrifuge, a belt filter, a plate filter, or the like) 91. After the 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 impurities such as chlorine ions, free ammonia, and hydroxide ions are adsorbed on the obtained sodium sulfate crystals, and in order to remove the adsorbed impurities, reduce the odor of solid salts, reduce corrosiveness, and improve the purity of the crystals, it is preferable that the sodium sulfate crystals are first washed with water, the catalyst production wastewater, or a sodium sulfate solution and dried.

The manner of the first solid-liquid separation and the first washing is not particularly limited, and may be carried out by using, for example, 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 first wash comprises panning and/or 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 sulfate crystals of higher purity. In the elutriation process, the waste water produced by the catalyst is generally not recycled when used as an elutriation liquid, and the washing liquid recovered by the first washing can be recycled in a counter-current manner when used as the elutriation liquid. Before the elutriation, it is preferable to perform a preliminary solid-liquid separation by sedimentation to obtain a slurry containing sodium sulfate crystals (the liquid content may be 35% by mass or less, and this step is preferably performed in an apparatus known in the art such as a sedimentation tank or a sedimentation tank). In the elutriation process, the liquid used for elutriation is 1 to 20 parts by weight relative to 1 part by weight of the slurry containing sodium sulfate crystals. In addition, the rinsing is preferably performed using an aqueous sodium sulfate solution. Preferably, the concentration of the sodium sulfate aqueous solution is the concentration of sodium sulfate in the saturated aqueous solution obtained by simultaneously obtaining sodium chloride and sodium sulfate at the temperature corresponding to the sodium sulfate crystal to be washed. In order to further enhance the elutriation effect and obtain sodium sulfate crystals with higher purity, the elutriation may preferably be performed using a liquid obtained by rinsing, and preferably using water or a sodium sulfate solution. The liquid resulting from the washing is preferably returned to the second heat exchange means 32, for example by a fifth circulation pump 75, before it is subjected to a second pH adjustment before it is evaporated by the first MVR evaporation means.

According to a preferred embodiment of the present invention, after the first concentrated solution containing sodium sulfate obtained by evaporation in the first MVR evaporation apparatus 2 is subjected to preliminary solid-liquid separation by settling, the catalyst production wastewater is subjected to first elutriation in an elutriation tank, then the liquid obtained in the subsequent sodium sulfate crystal washing is subjected to second elutriation in another elutriation tank, finally the slurry subjected to the two elutriations is sent to a solid-liquid separation apparatus for solid-liquid separation, the crystals obtained by the solid-liquid separation are subjected to elution again with an aqueous sodium sulfate solution, and the liquid obtained by the elution is returned to the second elutriation. Through the washing process, the purity of the obtained sodium sulfate crystal is improved, washing liquid is not excessively introduced, and the efficiency of wastewater treatment is improved.

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 evaporative crystallizer.

In the present invention, the evaporation conditions of the second evaporation may be appropriately selected as needed, so that sodium chloride is crystallized and sodium sulfate is not precipitated. The conditions of the second evaporation may include: the temperature is 30-85 ℃, and the pressure is-98 kPa-58 kPa. In order to improve evaporation efficiency, preferably, the conditions of the second evaporation include: the temperature is 35-60 ℃, and the pressure is-97.5 kPa to-87 kPa; preferably, the conditions of the second evaporation include: the temperature is 40 ℃ to 60 ℃, and the pressure is-97 kPa to-87 kPa; preferably, the conditions of the second evaporation include: the temperature is 45-60 ℃, and the pressure is-95 kPa to-87 kP; preferably, the conditions of the second evaporation include: the temperature is 45-55 ℃, and the pressure is-95 kPa to-90 kPa.

In the present invention, the operating pressure of the second evaporation is more 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 wastewater to be treated, and may be, for example, 0.1m³More than h (e.g. 0.1 m)³/h～500m³/h)。

By carrying out the second evaporation under the above conditions, the sodium sulfate is not crystallized while the crystallization of sodium chloride is ensured, so that the purity of the obtained sodium chloride crystal can be ensured.

According to the present invention, the second evaporation does not crystallize sodium sulfate in the second concentrated solution (i.e., sodium sulfate does not reach supersaturation), and preferably, the second evaporation makes the concentration of sodium sulfate in the second concentrated solution to be Y or less (preferably 0.9Y to 0.99Y, and more preferably 0.95Y to 0.98Y), where Y is the concentration of sodium sulfate at which both sodium sulfate and sodium chloride in the second concentrated solution reach saturation under the conditions of the second evaporation. By controlling the degree of the second evaporation within the above range, as much sodium chloride as possible can be crystallized out under the condition that sodium sulfate is not precipitated out. By crystallizing sodium chloride in the second evaporation as much as possible, the wastewater treatment efficiency can be improved, and the energy waste can be reduced.

In the present invention, the degree of progress of the second evaporation is performed by monitoring the concentration of the second evaporation-yielded liquid, and specifically, by controlling the concentration of the second evaporation-yielded liquid within the above range, the second evaporation is performed so as not to crystallize out sodium sulfate in the second concentrated solution. The concentration of the liquid resulting from the second evaporation is monitored by measuring the density, which may be carried out using a densitometer.

According to the invention, second ammonia-containing steam obtained by evaporation in the second MVR evaporation device 1 and the catalyst production wastewater are subjected to second heat exchange to obtain second ammonia water; preferably, the second ammonia-containing steam obtained by evaporation in the second MVR evaporation device 1 sequentially performs second heat exchange with the second washing liquid obtained after solid-liquid separation and the catalyst production wastewater to obtain second ammonia. The second heat exchange method is not particularly limited, and may be performed by a heat exchange method that is conventional in the art. The number of heat exchanges may be one or more, preferably 2 to 4, more preferably 2 to 3, and particularly preferably 2. Through the heat exchange, the output ammonia water is cooled, and the heat is circulated in the treatment device to the maximum extent, so that the energy is reasonably utilized, and the waste is reduced. The washing liquid after the second solid-liquid separation is subjected to the second heat exchange and then mixed with the liquid phase obtained by the first solid-liquid separation, wherein the temperature of the mixed liquid is 35 to 92 ℃, more preferably 40 to 67 ℃, even more preferably 45 to 67 ℃, and even more preferably 50 to 67 ℃.

According to a preferred embodiment of the present invention, the second heat exchange is performed by the third heat exchange device 33 and the fourth heat exchange device 34, specifically, the second ammonia-containing steam is sequentially passed through the fourth heat exchange device 34 and the third heat exchange device 33, the second heat exchange is performed between the washing liquid after the second solid-liquid separation and the second circulation liquid described later in the fourth heat exchange device 34 and the second ammonia-containing steam, and the second heat exchange is performed between a part of the catalyst production wastewater and the second ammonia-containing steam condensate obtained by condensing the second ammonia-containing steam by the fourth heat exchange device 34 in the third heat exchange device 33, so as to further reduce the temperature of the second ammonia-containing steam condensate, improve the heat energy utilization rate, and obtain the second ammonia, which can be stored in the second ammonia storage tank 52.

In the present invention, in order to increase the solid content in the second MVR evaporation device 1 and reduce the ammonia content in the liquid, it is preferable to return part of the liquid evaporated by the second MVR evaporation device 1 (i.e. the liquid located inside the second MVR evaporation device 1, also referred to as the second circulation liquid) to the second MVR evaporation device 1. In the above process of refluxing a part of the liquid evaporated by the second MVR evaporation device 1 to the second MVR evaporation device 1, preferably, the first mother liquid is mixed with the second circulation liquid and the second eluent, and then is subjected to the second heat exchange with the second ammonia-containing vapor. For example, the second circulating liquid and the second washing liquid can be mixed in the pipeline by a seventh circulating pump 77 and a ninth circulating pump 79 respectively, sent to the fourth heat exchange device 34 for the second heat exchange, then mixed with the first mother liquid and sent to the second MVR evaporation device 1. As a ratio of refluxing a part of the liquid evaporated by the second MVR evaporating device 1 to the second MVR evaporating device 1 is not particularly limited, for example, the second reflux ratio of the second evaporation may be 0.1 to 50, preferably 7 to 25. Here, the second reflux ratio means: 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.

According to the present invention, preferably, the method further comprises compressing the second ammonia-containing vapor before the second heat exchange. The compression of the second ammonia-containing vapor may be performed by a second compressor 102. Through right the second contains ammonia steam and compresses, for input energy among the MVR vaporization system, guarantee that waste water intensification-evaporation-cooling's process goes on in succession, need input when MVR vaporization process starts and start steam, only need pass through second compressor 102 energy supply after reaching continuous running state, no longer need input other energy. The second compressor 102 may be any one of various 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 ℃.

In the present invention, in order to prevent the first evaporation from crystallizing and precipitating sodium chloride and the second evaporation from crystallizing and precipitating sodium sulfate, it is preferable that the conditions of the two-time evaporation satisfy: the temperature of the first evaporation is at least 5 ℃ higher, preferably 20 ℃ higher, more preferably 35 ℃ to 70 ℃ higher, and particularly preferably 40 ℃ to 60 ℃ higher than the temperature of the second evaporation. And respectively crystallizing and separating out sodium sulfate and sodium chloride by controlling the first evaporation and the second evaporation to be carried out at different temperatures, so that the purity of the obtained sodium sulfate and sodium chloride crystals is improved.

According to a preferred embodiment of the present invention, the second evaporation process is carried out in a second MVR evaporation device 1. And introducing the first mother solution into the second MVR evaporation device 1 through a sixth circulating pump 76 for second evaporation to obtain second ammonia-containing steam and a second concentrated solution containing sodium chloride crystals.

In the invention, the second concentrated solution containing sodium chloride crystals is subjected to second solid-liquid separation to obtain sodium chloride crystals and a second mother liquor (namely, a liquid phase obtained by the second solid-liquid separation). The method of the second solid-liquid separation is not particularly limited, and may be selected from one or more of centrifugation, filtration, and sedimentation, for example.

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) is returned to the first MVR evaporation device 2 for the first evaporation again, and specifically, the second mother liquor can be returned to the first MVR evaporation device 2 by the eighth circulation pump 78 and mixed with the catalyst production wastewater after the first pH adjustment and before the second pH adjustment to obtain the wastewater to be treated, and then sent to the first MVR evaporation device 2 for the first evaporation. 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, it is preferable that the sodium chloride crystals are 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 sodium chloride aqueous solution is preferably the concentration of sodium chloride in the aqueous solution at which sodium chloride and sodium sulfate 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. Preferably, the second wash comprises panning and/or 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 catalyst production wastewater is generally not recycled when being used as an elutriation liquid, and the washing liquid recovered by the second washing can be recycled in a counter-current manner when being used as the elutriation liquid. Before the elutriation, a slurry containing sodium chloride crystals is preferably obtained by preliminary solid-liquid separation by sedimentation (the liquid content may be 35% by mass or less, and this step is preferably performed in an apparatus known in the art, such as a sedimentation tank or a sedimentation tank). 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, the elutriation may be preferably performed using a liquid obtained by rinsing. For the liquid generated by washing, it is preferable that the catalyst production wastewater elutriation liquid is returned to the second MVR evaporation device 1 before being returned to the first MVR evaporation device 2 for the second pH adjustment, and other washing liquid is returned to the second MVR evaporation device 1 for the second evaporation again, for example, by the ninth circulation pump 79.

According to a preferred embodiment of the present invention, after the primary solid-liquid separation by settling, the second concentrated solution containing sodium chloride crystals obtained by crystallization is subjected to a first elutriation in an elutriation tank using the catalyst production wastewater, then the liquid obtained by the subsequent sodium chloride crystal washing is subjected to a second elutriation in another elutriation tank, finally the slurry subjected to the two elutriations is sent to a solid-liquid separation device for solid-liquid separation, the crystals obtained by the solid-liquid separation are washed again 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 second 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 is not excessively introduced, and the efficiency of wastewater treatment is improved.

According to a preferred embodiment of the present invention, the tail gas remaining after the first ammonia-containing steam is condensed by the first heat exchange and the tail gas remaining after the second ammonia-containing steam is condensed by the second heat exchange are discharged after ammonia removal. The first ammonia-containing steam is subjected to the first heat exchange to condense the remaining tail gas, i.e., the tail gas discharged from the second heat exchange device 32, and the second ammonia-containing steam is subjected to the second heat exchange to condense the remaining tail gas, i.e., the tail gas discharged from the fourth heat exchange device 34. The ammonia in the tail gas is removed, so that the pollutant content in the discharged tail gas can be further reduced, and 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 NH₄ ⁺、SO₄ ^2-、Cl^-And Na⁺The wastewater is obtained. In addition, the method is particularly suitable for treating high-salinity wastewater. The wastewater of the present invention may be specifically wastewater from a process for producing a molecular sieve, alumina or a refinery catalyst, or wastewater from a process for producing a molecular sieve, alumina or a refinery catalyst, which is subjected to 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 wastewater₄ ⁺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 1000mg/L or more, more preferably 2000mg/L or more, further preferably 4000mg/L or more, further preferably 8000mg/L or more, further preferably 16000mg/L or more, further preferably 32000mg/L or more, further preferably 40000mg/L or more, further preferably 50000mg/L or more, further preferably 60000mg/L or more.

As SO in wastewater from the production of said catalyst₄ ^2-May be 1000mg/L or more, preferably 2000mg/L or more, more preferably 4000mg/L or more, further preferably 8000mg/L or more, further preferably 16000mg/L or more, further preferably 32000mg/L or more, further preferably 40000mg/L or more, further preferably 50000mg/L or more, further preferably 60000mg/L or more, further preferably 70000mg/L or more.

As Cl in the catalyst production wastewater^-May be 970mg/L or more, more preferably 2000mg/L or more, further preferably 4000mg/L or more, further preferably 8000mg/L or more, further preferably 16000mg/L or more, further preferably 32000mg/L or more, more preferably 40000mg/L or more, still more preferably 50000mg/L or more, and still more preferably 60000mg/L or more.

NH contained in the catalyst production wastewater₄ ⁺、SO₄ ^2-、Cl^-And Na⁺The upper limit of (3) is not particularly limited. SO in the wastewater from the viewpoint of easy access to wastewater₄ ^2-、Cl^-And Na⁺Respectively 200g/L or less, preferably 150g/L or less, and more preferably 120g/L or less; NH in wastewater₄ ⁺Is 50g/L or less, preferably 40g/L or less.

From the viewpoint of improving the efficiency of the first evaporation and reducing the energy consumption of the treatment process, the amount of SO contained in the wastewater is relatively small₄ ^2-Cl in catalyst production wastewater^-The lower the content, the better, for example, relative to 1 mole of SO contained in the catalyst production wastewater₄ ^2-Cl contained in the catalyst production wastewater^-Is 30 mol or less, preferably 20 mol or less, more preferably 15 mol or less, and still more preferably 10 mol or less. From the viewpoint of practicality, the amount of SO contained in the wastewater from the catalyst production is 1 mole₄ ^2-Cl contained in the catalyst production wastewater^-Preferably 0.1 mol or more, more preferably 0.5 mol or more, further preferably 1 mol or more, for example, 1 to 8 mol. By adding SO contained in the catalyst production wastewater₄ ^2-And Cl^-The molar ratio of (a) to (b) is limited to the above range, most of water can be evaporated in the first evaporation, the amount of circulating liquid in a treatment system is reduced, energy is saved, and the treatment process is more economical.

In the present invention, the inorganic salt ions contained in the catalyst production wastewater are other than NH₄ ⁺、SO₄ ^2-、Cl^-And Na⁺In addition, it may contain Mg²⁺、Ca²⁺、K⁺、Fe³⁺Inorganic salt ions such as rare earth element ions, Mg²⁺、Ca²⁺、K⁺、Fe³⁺Rare earth elementThe content of each inorganic salt ion such as a cellulose ion is preferably 100mg/L or less, more preferably 50mg/L or less, further 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 sodium 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 1600mg/L or more, preferably 4000mg/L or more, more preferably 8000mg/L or more, further preferably 16000mg/L or more, further preferably 32000mg/L or more, further preferably 40000mg/L or more, further preferably 50000mg/L or more, further preferably 60000mg/L or more, further preferably 100000mg/L or more, further preferably 150000mg/L or more, further preferably 200000mg/L or more.

In the invention, the pH value of the catalyst production wastewater is preferably 4-8, and more preferably 6.2-6.8.

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

In the invention, in order to reduce the concentration of impurity ions in the wastewater, ensure the continuous and stable operation of the 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; as the chemical precipitation, pH adjustment, carbonate precipitation, magnesium salt precipitation, and the like may be mentioned; 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. For 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 catalyst production 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 impurity removal). 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 the 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 treating the catalyst production wastewater 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 generated after wastewater generated in the molecular sieve production process is subjected to chemical precipitation, filtration, weak acid cation exchange and ozone biological activated carbon adsorption oxidation for impurity removal, and is concentrated by an ED membrane and a reverse osmosis method.

The conditions for the above chemical precipitation are preferably: sodium carbonate is used as a treating agent, 1.2-1.4 mol of sodium carbonate is added relative to 1 mol 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 0.5-4 h.

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 mm-1.7 mm, the grain diameter of the quartz sand is 0.5 mm-1.3 mm, and the filtering speed is 10 m/h-30 m/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 10-15 h.

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 1.5-3.0 m, the HCl concentration of the regeneration liquid is as follows: 4.5-5 mass%; the amount of the regenerant (calculated by 100%) was 50kg/m³～60kg/m³Wet resin; the flow rate of the regeneration liquid HCl is 4.5 m/h-5.5 m/h, and the regeneration contact time is 35 min-45 min; the forward washing flow rate is 18 m/h-22 m/h, and the forward washing time is 20 min-30 min; the running flow rate is 15 m/h-30 m/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 50min to 70min, and the empty bed filtration rate is 0.5m/h to 0.7 m/h.

The conditions for the concentration of the ED membrane are preferably: the current 145A to 155A and the voltage 45V to 65V. The ED membrane may be, for example, an ED membrane manufactured by astone corporation of japan.

The conditions for the reverse osmosis are preferably: the operation pressure is 5.4MPa to 5.6MPa, the water inlet temperature is 25 ℃ to 35 ℃, and the pH value is 6.5 to 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 direct operation, and if the ion content of the catalyst production wastewater meets the conditions of the invention, the first evaporation can be carried out firstly and then 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, the first evaporation can be controlled to ensure that the concentration of sodium chloride in the first concentrated solution is close to the precipitation concentration, then the first concentrated solution is subjected to the second evaporation to obtain a second concentrated solution, the solid-liquid separation is carried out to obtain sodium chloride crystals and a second mother solution, the second mother solution is mixed with the catalyst production wastewater to adjust the ion content of the wastewater to be treated to be within the range required by the invention, and then the first evaporation is carried out to obtain sodium sulfate crystals. Of course, the ion content of the wastewater to be treated can be adjusted by using sodium sulfate or sodium chloride in the initial stage as long as the wastewater to be treated satisfies the SO content in the wastewater to be treated in the present invention₄ ^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 FIG. 1, catalyst production waste water (containing NaCl 80g/L, Na)₂SO₄ 82g/L、NH₄Cl 50g/L、(NH₄)₂SO₄52.1g/L, pH 6.8) at a feed rate of 5m³The reaction mixture was fed to a vacuum degassing tank 4 at a rate of/h to carry out vacuum degassing, and an aqueous sodium hydroxide solution having a concentration of 45.16 mass% was introduced into a line leading to the first heat exchanger 31 (titanium alloy plate heat exchanger)Adjusting the pH value for the first time, monitoring the adjusted pH value (the measured value is 7.5) by a first pH value measuring device 61 (a pH meter), sending part of wastewater after the first pH value adjustment into a first heat exchange device 31 through a first circulating pump 71, carrying out first heat exchange with the recycled first ammonia-containing steam condensate to heat the wastewater to 98 ℃, sending the other part of wastewater after the first pH value adjustment into a third heat exchange device 33 through the first circulating pump 71, carrying out first heat exchange with the second ammonia-containing steam condensate to heat the wastewater to 48 ℃, merging the two parts of wastewater, and sending the merged wastewater with the returned second mother liquor (the sending speed is 8.75 m)³H) mixing to obtain wastewater to be treated (measuring SO contained in the wastewater₄ ^2-And Cl^-In a molar ratio of 1: 6.261), introducing a sodium hydroxide aqueous solution with the concentration of 45.16 mass% into a pipeline for conveying the wastewater to be treated into the second heat exchange device 32 to adjust the pH value for the second time, monitoring the adjusted pH value through a second pH value measuring device 62 (a pH meter) (the measured value is 10.8), conveying the wastewater to be treated into the second heat exchange device 32 (a titanium alloy plate heat exchanger) to perform first heat exchange with the recovered first ammonia-containing steam to heat the wastewater to be treated to 107 ℃, conveying the wastewater to be treated after the first heat exchange into the first MVR evaporation device 2 (a falling film and forced circulation two-stage MVR evaporation crystallizer) to evaporate, and obtaining first ammonia-containing steam and a first concentrated solution containing sodium sulfate crystals. Wherein the evaporation conditions of the first MVR evaporation device 2 include: the temperature is 100 ℃, the pressure is-22.82 kPa, and the evaporation capacity is 3.92m³H is used as the reference value. The first ammonia-containing steam obtained by evaporation is compressed by the first compressor 101 (the temperature rises by 12 ℃) and then sequentially passes through the second heat exchange device 32 and the first heat exchange device 31 to exchange heat with the wastewater to be treated, and is cooled to obtain first ammonia water which is stored in the first ammonia water storage tank 51. In addition, in order to increase the solid content in the first MVR evaporation device 2, part of the liquid evaporated in the first MVR evaporation device 2 is circulated as a first circulation liquid to the second heat exchange device 32 by the second circulation pump 72, and then enters the first MVR evaporation device 2 again for first evaporation (the first reflux ratio is 77.8). The density meter arranged on the first MVR evaporation device 2 is used for monitoring the degree of the first evaporation and controlling the sodium chloride in the first concentrated solutionThe concentration of (2) was 0.9935X (306.2 g/L).

The first concentrated solution obtained by evaporation in the first MVR evaporation plant 2 is sent to a first solid-liquid separation device 91 (centrifugal machine) for first solid-liquid separation, and 10.58m is obtained per hour³Contains NaCl 306.2g/L, Na₂SO₄ 54.0g/L、NaOH 1.4g/L、NH₃0.27g/L of first mother liquor is temporarily stored in a first mother liquor tank 53, sodium sulfate solid obtained by solid-liquid separation (809.7 kg of sodium sulfate crystal filter cake containing 15 mass% of water is obtained per hour, wherein the content of sodium chloride is below 6.9 mass%) is leached by 54g/L of sodium sulfate solution which is equal to the dry basis mass of the sodium sulfate crystal filter cake, 688.25kg of sodium sulfate (the purity is 99.7 weight%) is obtained per hour after drying, and a washing solution is circulated to a pipeline before entering a second heat exchange device 32 through a fifth circulating pump 75 to be mixed with wastewater, and then enters the first MVR evaporation device 2 again for first evaporation.

The second evaporation process is carried out in a second MVR evaporation plant 1 (falling film + forced circulation two-stage MVR evaporator crystallizer). And (3) sending the first mother liquor in the first mother liquor tank 53 to the second MVR evaporation device 1 through a sixth circulating pump 76 for second evaporation to obtain a second concentrated solution containing sodium chloride crystals. Wherein the evaporation conditions of the second MVR evaporation device 1 include: the temperature is 50 ℃, the pressure is-92.67 kPa, and the evaporation capacity is 1.95m³H is used as the reference value. In order to increase the solid content in the second MVR evaporation device 1, part of the first mother liquor after evaporation in the second MVR evaporation device 1 is circulated as a second circulation liquid to the fourth heat exchange device 34 through the seventh circulation pump 77 to exchange heat with the second ammonia-containing vapor, and then enters the second MVR evaporation device 1 again to perform second evaporation (the second reflux ratio is 17.8). The second ammonia-containing vapor obtained by evaporation is compressed by the second compressor 102 (the temperature is raised by 12 ℃) and then passes through the fourth heat exchange device 34 and the third heat exchange device 33 in sequence for heat exchange and cooling to obtain second ammonia, and the second ammonia is stored in the second ammonia storage tank 52. The washing liquid after the second solid-liquid separation and part of the first mother liquid after the evaporation in the second MVR evaporation device 1 are respectively pumped out by a ninth circulating pump 79 and a seventh circulating pump 77, are mixed in a pipeline and then carry out heat exchange with the second ammonia-containing steam passing through the fourth heat exchange device 34, and the first circulation is carried outPart of the wastewater to be treated transferred by the ring pump 71 is heat-exchanged with the condensate of the second ammonia-containing vapor passing through the third heat exchanging means 33. The degree of the second evaporation was monitored by a densitometer provided in the second MVR evaporation apparatus 1, and the concentration of sodium sulfate in the second concentrated solution was controlled to 0.9702Y (65.3 g/L). After the first mother liquor is evaporated in the second MVR evaporation device 1, a second concentrated solution containing sodium chloride crystals is obtained.

The second concentrated solution containing sodium chloride crystals is sent to a second solid-liquid separation device 92 (centrifugal machine) for solid-liquid separation, and 8.75m is obtained every hour³Contains 293.8g/L, Na NaCl₂SO₄ 65.3g/L、NaOH 1.7g/L、NH₃0.013g/L of the second mother liquor, and the second mother liquor is temporarily stored in the second mother liquor tank 54. The second mother liquor is circulated to a wastewater introduction pipeline before the second pH value adjustment through an eighth circulating pump 78 to be mixed with wastewater to obtain wastewater to be treated, sodium chloride solid obtained by solid-liquid separation (785.29 kg of sodium chloride crystal filter cake with the water content of 14 mass% is obtained per hour, wherein the content of sodium sulfate is below 7.1 mass%) is subjected to leaching by using 293.8g/L sodium chloride solution with the same mass as the dry basis mass of sodium chloride, and is dried in a dryer to obtain 675.35kg of sodium chloride (the purity is 99.4 weight%) per hour, and the second washing liquid obtained by washing is circulated to a second MVR evaporation device 1 through a ninth circulating pump 79.

In addition, the tail gas discharged by the vacuum degassing tank 4, the second heat exchange device 32 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 a 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.

In this example, 3.92m of ammonia water having a concentration of 3.53 mass% was obtained per hour in the first ammonia water tank 51³1.95m of 0.144 mass% ammonia water was obtained per hour in the second ammonia water tank 52³The ammonia water can be reused in the molecular sieveThe production process of (1).

In addition, the starting phase of the MVR evaporation was started by steam at a temperature of 143.3 ℃.

Example 2

The treatment of the waste water was carried out according to the method of example 1, except that: for NaCl-containing 60g/L, Na₂SO₄ 130g/L、NH₄Cl 15g/L、(NH₄)₂SO₄33.0g/L, pH of 6.8 was treated to obtain SO in the wastewater to be treated₄ ^2-And Cl^-In a molar ratio of 1: 4.462. the temperature of the wastewater after heat exchange by the first heat exchange device 31 was 67 deg.C, and the temperature of the wastewater to be treated after heat exchange by the second heat exchange device 32 was 102 deg.C. The evaporation conditions of the first MVR evaporation device 2 include: the temperature is 95 ℃, the pressure is-36.36 kPa, and the evaporation capacity is 4.48m³H is used as the reference value. The evaporation conditions of the second MVR evaporation device 1 include: the temperature is 55 deg.C, the pressure is-90.15 kPa, and the evaporation capacity is 1.05m³/h。

The first solid-liquid separation device 91 obtained 978.40kg of sodium sulfate crystal cake containing 15 mass% of water per hour, and finally obtained 831.63kg of sodium sulfate (purity 99.3 wt%) per hour; 7.25 m/hr³The concentration of NaCl is 307.2g/L, Na₂SO₄ 54.5g/L、NaOH 1.8g/L、NH₃0.18g/L of the first mother liquor.

The second solid-liquid separation device 92 obtained 444.79kg of sodium chloride crystal cake with a water content of 15 mass% per hour, and finally 378.07kg of sodium chloride (purity 99.5 wt%) per hour; yield 6.30m per hour³The concentration of NaCl is 295.5g/L, Na₂SO₄ 63.1g/L、NaOH 2.1g/L、NH₃0.01g/L of a second mother liquor.

In this example, 4.48m of ammonia water having a concentration of 1.4 mass% was obtained per hour in the first ammonia water tank 51³1.05m of aqueous ammonia having a concentration of 0.11 mass% was obtained per hour in the second aqueous ammonia tank 52³The ammonia water can be reused in the production process of the molecular sieve.

Example 3

The treatment of the waste water was carried out according to the method of example 1, except thatThe method comprises the following steps: for NaCl-containing 160g/L, Na₂SO₄ 55g/L、NH₄Cl 32g/L、(NH₄)₂SO₄11.2g/L, pH of 6.2 was treated to obtain SO in the wastewater to be treated₄ ^2-And Cl^-In a molar ratio of 1: 9.249. the temperature of the wastewater after heat exchange by the first heat exchange means 31 was 73 deg.c and the temperature of the wastewater to be treated after heat exchange by the second heat exchange means 32 was 112 deg.c. The evaporation conditions of the first MVR evaporation device 2 include: the temperature is 105 ℃, the pressure is-7.02 kPa, and the evaporation capacity is 2.63m³H is used as the reference value. The evaporation conditions of the second MVR evaporation device 1 include: the temperature was 45 ℃, the pressure was-94.69 kPa, and the evaporation capacity was 2.86m³/h。

The first solid-liquid separation device 91 obtained 385.26kg of a sodium sulfate crystal cake containing 14 mass% of water per hour, and finally obtained 331.32kg of sodium sulfate (purity: 99.4 wt%) per hour; yield 11.98m per hour³The concentration of NaCl is 306.4g/L, Na₂SO₄ 52.5g/L、NaOH 2.6g/L、NH₃0.11g/L of the first mother liquor.

The second solid-liquid separation device 92 obtained 1151.57kg of sodium chloride crystal cake with a water content of 15 mass% per hour, and finally 978.83kg of sodium chloride (purity 99.4 wt%) per hour; obtained at 9.18m per hour³The concentration of NaCl 291.2g/L, Na₂SO₄ 67.9g/L、NaOH 3.4g/L、NH₃0.0084g/L of second mother liquor.

In this example, 2.63m of aqueous ammonia having a concentration of 2.3 mass% was obtained per hour in the first aqueous ammonia tank 51³2.86m of 0.043 mass% ammonia water per hour is obtained in the second ammonia water tank 52³The ammonia water can be reused in the production process of the molecular sieve.

The preferred embodiments of the present invention have been described in detail, however, the present invention is not limited to the specific details of the above embodiments, and various simple modifications may be made to the technical solution of the present invention within the technical idea of the present invention, and these simple modifications are within the protective scope of the present invention.

It should be noted that the various features described in the above embodiments may be combined in any suitable manner without departing from the scope of the invention. The invention is not described in detail in order to avoid unnecessary repetition.

In addition, any combination of the various embodiments of the present invention is also possible, and the same should be considered as the disclosure of the present invention as long as it does not depart from the spirit of the present invention.

Claims

1. Method for treating catalyst production wastewater containing NH₄ ⁺、SO₄ ^2-、Cl^-And Na⁺Characterized in that the method comprises the following steps,

1) introducing wastewater to be treated into a first MVR evaporation device to carry out first evaporation to obtain first concentrated solution containing ammonia vapor and sodium sulfate crystals, wherein the wastewater to be treated contains the catalyst production wastewater and a liquid phase obtained by second solid-liquid separation;

adjusting the pH value of the wastewater to be treated to be more than 9 before introducing the wastewater to be treated into a first MVR evaporation device;

the first evaporation prevents sodium chloride from crystallizing out, and the second evaporation prevents sodium sulfate from crystallizing out;

the conditions of the first evaporation include: the temperature is 75-175 ℃, and the pressure is-73 kPa-653 kPa; the conditions of the second evaporation include: the temperature is 35-60 ℃, and the pressure is-97.5 kPa to-87 kPa;

relative to 1 mole of SO contained in the wastewater to be treated₄ ^2-Cl contained in the wastewater to be treated^-14 mol or less;

NH in the catalyst production wastewater₄ ⁺Is more than 8mg/L, SO₄ ^2-Over 1000mg/L, Cl^-Over 970mg/L of Na⁺Is more than 510 mg/L.

2. The method according to claim 1, wherein the SO contained in the wastewater to be treated is 1 mole relative to the SO contained in the wastewater to be treated₄ ^2-Cl contained in the wastewater to be treated^-Is 13.8 mol or less.

3. The method of claim 1, wherein the pH of the wastewater to be treated is adjusted to greater than 10.8 prior to passing the wastewater to be treated to the first MVR evaporation unit.

4. The method as claimed in claim 1, wherein the pH adjustment of the wastewater to be treated is carried out with NaOH.

5. The method of claim 1, wherein the first evaporation is conducted such that the concentration of sodium chloride in the first concentrated solution is no greater than X, where X is the concentration of sodium chloride at which both sodium sulfate and sodium chloride in the first concentrated solution are saturated under the conditions of the first evaporation.

6. The process of claim 5, wherein the second evaporation is conducted such that the concentration of sodium sulfate in the second concentrated solution is Y or less, wherein 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.

7. A process as claimed in claim 5, wherein the first evaporation provides a concentration of sodium chloride in the first concentrate of from 0.95X to 0.999X.

8. The process of claim 6, wherein the second evaporation provides a sodium sulfate concentration in the second concentrated solution of 0.9Y to 0.99Y.

9. The method of any one of claims 1-8, wherein the conditions of the first evaporation comprise: the temperature is 80-130 ℃, and the pressure is-66 kPa-117 kPa.

10. The method of claim 9, wherein the conditions of the first evaporation comprise: the temperature is 95-105 ℃, and the pressure is-37 kPa to-7 kPa.

11. The method of any one of claims 1-8, wherein the conditions of the second evaporation comprise: the temperature is 40-60 ℃, and the pressure is-97 kPa to-87 kPa.

12. The method of claim 11, wherein the conditions of the second evaporation comprise: the temperature is 45-60 ℃, and the pressure is-95 kPa to-87 kPa.

13. The method of claim 12, wherein the conditions of the second evaporation comprise: the temperature is 45-55 ℃, and the pressure is-95 kPa to-90 kPa.

14. The method of claim 9, wherein the temperature of the first evaporation is more than 20 ℃ higher than the temperature of the second evaporation.

15. The method of claim 14, wherein the temperature of the first evaporation is 35 ℃ to 70 ℃ higher than the temperature of the second evaporation.

16. The method of claim 15, wherein the temperature of the first evaporation is 40 ℃ to 60 ℃ higher than the temperature of the second evaporation.

17. The method of claim 1, wherein the first ammonia-containing vapor is subjected to a first heat exchange with the wastewater to be treated and a first ammonia solution is obtained before the wastewater to be treated is passed to a first MVR evaporation plant.

18. A method according to claim 17, wherein the pH of the wastewater to be treated is adjusted to be greater than 7 prior to the first heat exchange.

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

20. The method according to any one of claims 1 to 8, further comprising subjecting the first concentrated solution containing sodium sulfate crystals to a first solid-liquid separation to obtain sodium sulfate crystals.

21. The method of claim 20, further comprising washing the resulting sodium sulfate crystals.

22. The method according to any one of claims 1 to 8, further comprising subjecting the second concentrated solution containing sodium chloride crystals to a second solid-liquid separation to obtain sodium chloride crystals.

23. The method of claim 22, further comprising washing the resulting sodium chloride crystals.

24. The process of any one of claims 1 to 8, wherein the catalyst production wastewater is wastewater from a molecular sieve, alumina or refinery catalyst production process.

25. The method of claim 24, further comprising removing impurities and concentrating the catalyst process wastewater.