WO2020238942A1 - 牛磺酸中间体牛磺酸钠的制备方法及牛磺酸的制备方法 - Google Patents

牛磺酸中间体牛磺酸钠的制备方法及牛磺酸的制备方法 Download PDF

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WO2020238942A1
WO2020238942A1 PCT/CN2020/092547 CN2020092547W WO2020238942A1 WO 2020238942 A1 WO2020238942 A1 WO 2020238942A1 CN 2020092547 W CN2020092547 W CN 2020092547W WO 2020238942 A1 WO2020238942 A1 WO 2020238942A1
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ammonia
taurine
sodium
separator
stage
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PCT/CN2020/092547
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English (en)
French (fr)
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陈志荣
姚祥华
彭俊华
潘映霞
何孝祥
徐淞华
吴晓东
方向
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浙江新和成股份有限公司
浙江大学
上虞新和成生物化工有限公司
浙江新和成药业有限公司
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Priority to US17/424,454 priority Critical patent/US20220081394A1/en
Publication of WO2020238942A1 publication Critical patent/WO2020238942A1/zh

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    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C303/00—Preparation of esters or amides of sulfuric acids; Preparation of sulfonic acids or of their esters, halides, anhydrides or amides
    • C07C303/32—Preparation of esters or amides of sulfuric acids; Preparation of sulfonic acids or of their esters, halides, anhydrides or amides of salts of sulfonic acids
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C309/00—Sulfonic acids; Halides, esters, or anhydrides thereof
    • C07C309/01—Sulfonic acids
    • C07C309/02—Sulfonic acids having sulfo groups bound to acyclic carbon atoms
    • C07C309/03—Sulfonic acids having sulfo groups bound to acyclic carbon atoms of an acyclic saturated carbon skeleton
    • C07C309/13—Sulfonic acids having sulfo groups bound to acyclic carbon atoms of an acyclic saturated carbon skeleton containing nitrogen atoms, not being part of nitro or nitroso groups, bound to the carbon skeleton
    • C07C309/14—Sulfonic acids having sulfo groups bound to acyclic carbon atoms of an acyclic saturated carbon skeleton containing nitrogen atoms, not being part of nitro or nitroso groups, bound to the carbon skeleton containing amino groups bound to the carbon skeleton
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D3/00—Distillation or related exchange processes in which liquids are contacted with gaseous media, e.g. stripping
    • B01D3/06—Flash distillation
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D61/00—Processes of separation using semi-permeable membranes, e.g. dialysis, osmosis or ultrafiltration; Apparatus, accessories or auxiliary operations specially adapted therefor
    • B01D61/42—Electrodialysis; Electro-osmosis ; Electro-ultrafiltration; Membrane capacitive deionization
    • B01D61/44—Ion-selective electrodialysis
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D61/00—Processes of separation using semi-permeable membranes, e.g. dialysis, osmosis or ultrafiltration; Apparatus, accessories or auxiliary operations specially adapted therefor
    • B01D61/42—Electrodialysis; Electro-osmosis ; Electro-ultrafiltration; Membrane capacitive deionization
    • B01D61/44—Ion-selective electrodialysis
    • B01D61/445—Ion-selective electrodialysis with bipolar membranes; Water splitting
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J19/00—Chemical, physical or physico-chemical processes in general; Their relevant apparatus
    • B01J19/24—Stationary reactors without moving elements inside
    • B01J19/2455—Stationary reactors without moving elements inside provoking a loop type movement of the reactants
    • B01J19/2465—Stationary reactors without moving elements inside provoking a loop type movement of the reactants externally, i.e. the mixture leaving the vessel and subsequently re-entering it
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J19/00—Chemical, physical or physico-chemical processes in general; Their relevant apparatus
    • B01J19/24—Stationary reactors without moving elements inside
    • B01J19/2475—Membrane reactors
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J3/00—Processes of utilising sub-atmospheric or super-atmospheric pressure to effect chemical or physical change of matter; Apparatus therefor
    • B01J3/008—Processes carried out under supercritical conditions
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C303/00—Preparation of esters or amides of sulfuric acids; Preparation of sulfonic acids or of their esters, halides, anhydrides or amides
    • C07C303/02—Preparation of esters or amides of sulfuric acids; Preparation of sulfonic acids or of their esters, halides, anhydrides or amides of sulfonic acids or halides thereof
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C303/00—Preparation of esters or amides of sulfuric acids; Preparation of sulfonic acids or of their esters, halides, anhydrides or amides
    • C07C303/02—Preparation of esters or amides of sulfuric acids; Preparation of sulfonic acids or of their esters, halides, anhydrides or amides of sulfonic acids or halides thereof
    • C07C303/22—Preparation of esters or amides of sulfuric acids; Preparation of sulfonic acids or of their esters, halides, anhydrides or amides of sulfonic acids or halides thereof from sulfonic acids, by reactions not involving the formation of sulfo or halosulfonyl groups; from sulfonic halides by reactions not involving the formation of halosulfonyl groups
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D2311/00—Details relating to membrane separation process operations and control
    • B01D2311/04—Specific process operations in the feed stream; Feed pretreatment
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D2311/00—Details relating to membrane separation process operations and control
    • B01D2311/26—Further operations combined with membrane separation processes
    • B01D2311/263—Chemical reaction

Definitions

  • the invention relates to the technical field of preparation of taurine, in particular to an ammonolysis reaction system, a preparation method of taurine intermediate taurine sodium, and a preparation method of taurine.
  • the preparation methods of taurine mainly include biological extraction method, fermentation method and chemical synthesis method, among which the research of chemical synthesis method is the fastest.
  • chemical synthesis methods of taurine There are currently more than 20 chemical synthesis methods of taurine according to different raw materials and processes.
  • Ethanolamine method Using ethanolamine as a raw material, taurine is synthesized in two steps. According to the synthesis route, it can be divided into esterification method, chlorination method, and ethyleneimine method. Among them, the raw materials of the esterification method are easy to obtain, and the yield is higher than other methods. It is adopted by most domestic and foreign manufacturers.
  • the reaction uses ethanolamine, sulfuric acid, and sodium sulfite as raw materials. First, sulfuric acid and ethanolamine are esterified to synthesize intermediate 2-aminoethyl Sulfate ester is sulfonated with sodium sulfite or ammonium sulfite to synthesize taurine.
  • the reaction equation is as follows:
  • esterification reaction is a reversible reaction, and the reaction is not complete, which restricts the conversion rate and reaction yield of ethanolamine, and the reaction system generates sodium sulfate, which easily causes separation difficulties, affects the product yield and quality, and causes great environmental protection pressure.
  • Ethylene oxide method Using ethylene oxide as a raw material, it is firstly added with sodium sulfite by ring-opening, and then reacted with ammonia under heating and pressure conditions to synthesize sodium taurate, and acidified to obtain taurine.
  • the reaction process is as follows:
  • the ethylene oxide method includes addition, ammonolysis, and acidification steps, and its yield is higher than that of the ethanolamine method, and is currently widely used.
  • Patent US1932907 mentions the ammonolysis reaction of isethionate and amines, wherein the molar ratio of ammonia to isethionate is 6.8:1, and the reaction temperature is 240°C to 250°C for 2 hours. The yield of sodium taurate is only 80%.
  • Patent DD219023A3 mentions the composition of the ammonolysis product of sodium isethionate.
  • the mother liquor contains sodium sulfate and ethane.
  • the untreated mother liquor is circulated into the system with a variety of complex components such as alcohol, polyethylene glycol, and trace metal elements, as the number of cycles increases, a large number of impurities in the system accumulate, which is not conducive to the reaction.
  • the ammonolysis reaction is usually carried out in the form of excess ammonia. After the ammonolysis is completed, a deamination treatment is required. Usually, the ammonolysis liquid has a higher temperature and pressure, which is formed during the deamination process. The ammonia-containing gas phase still has a certain amount of heat. Some people have done some research on the recovery and utilization of this part of the heat.
  • the patent CN101528658 discloses the treatment method of the ammonia solution, and the ammonia solution is flashed through the first stage and the second stage. Flashing falling film evaporation and multi-effect falling film evaporation concentration are used for processing.
  • Flash steam is used as the heat source medium of the next stage evaporator for heating, but the patent does not mention how to deal with the recovered ammonia.
  • the common practice is to condense high-content ammonia back to the ammonolysis step. After condensing low-content ammonia, use equipment such as ammonia distillation towers for refining, and then recycle after reaching a certain concentration.
  • the treated ammonia is recycled back to the ammonolysis step, in order to achieve the high-temperature and high-pressure reaction conditions of the ammonolysis, it needs to be heated under pressure and also consumes a large amount of energy. It does not involve how to recycle the ammonia in a low-energy way.
  • One aspect of the present invention provides a preparation method of taurine intermediate taurine sodium, the preparation method comprising the following steps:
  • the sodium isethionate and the ammonia source are placed in an ammonolysis reactor for an ammonolysis reaction to obtain a mixture containing taurine intermediate taurate, wherein the ammonia in the ammonia source is The molar ratio of the sodium isethionate is 25:1 or more.
  • the molar ratio of ammonia in the ammonia source to the sodium isethionate is 25:1-100:1, more preferably 30:1-50:1.
  • ammonia separation device Separate the unreacted ammonia in the mixture by an ammonia separation device to obtain ammonia-containing gaseous substances and taurine intermediates respectively, wherein the ammonia separation device is connected to the ammonolysis reactor;
  • the ammonia-containing gaseous substance is compressed by a compression device to obtain an ammonia-containing supercritical fluid, and the supercritical fluid is circulated to the ammonolysis reactor, wherein the compression device is separated from the ammonia The device and the ammonolysis reactor are connected.
  • the ammonia separation device includes an ammonia separator.
  • the ammonia separation device includes two ammonia separators, a first-stage ammonia separator and a second-stage ammonia separator, respectively,
  • the first-stage ammonia separator is connected to the ammonolysis reactor, and the first-stage ammonia separator is used to separate the ammonia that is not involved in the reaction in the mixture after the ammonolysis reaction to obtain the first ammonia-containing The gaseous substance and the first remaining mixture;
  • the second-stage ammonia separator is connected to the first-stage ammonia separator, and the second-stage ammonia separator is used to further separate the first remaining mixture with ammonia to obtain a second ammonia-containing gaseous substance and The second residue, and the second ammonia-containing gaseous substance is recycled to the first-stage ammonia separator.
  • the compression device includes a first compression device and a second compression device
  • the first compression device is respectively connected with the first-stage ammonia separator and the ammonolysis reactor, and is used to compress the ammonia-containing gaseous substance in the first-stage ammonia separator to obtain the Supercritical fluid, and circulating the supercritical fluid to the ammonolysis reactor;
  • the second compression device is respectively connected with the first-stage ammonia separator and the second-stage ammonia separator, and is used for recycling the second ammonia-containing gaseous substance to the first-stage ammonia separator.
  • the ammonia separation device includes n ammonia separators arranged in sequence, and n is an integer greater than 2 and less than 20,
  • the first-stage ammonia separator of the n ammonia separators arranged in sequence is connected to the ammonolysis reactor, and the first-stage ammonia separator is used to remove the impurities in the mixture obtained after the ammonolysis reaction.
  • the ammonia participating in the reaction is separated to obtain the first ammonia-containing gaseous substance and the first remaining mixture;
  • the second-stage ammonia separator among the n ammonia separators arranged in sequence is connected to the first-stage ammonia separator, and the second-stage ammonia separator is used to perform ammonia again on the first remaining mixture. Separating to obtain a second ammonia-containing gaseous substance and a second residue, and recycling the second ammonia-containing gaseous substance to the first-stage ammonia separator;
  • the i-th ammonia separator of the n ammonia separators arranged in sequence is connected to the i-1th ammonia separator, i is an integer and 2 ⁇ i ⁇ n, and the i-th ammonia separator is used for
  • the i-1th residue mixture obtained by the i-1th-stage ammonia separator is further subjected to ammonia separation to obtain the i-th ammonia-containing gaseous substance and the i-th residue, and the i-th ammonia-containing gaseous substance is recycled to The i-1 level ammonia separator.
  • the ammonolysis reaction system further includes n compression devices,
  • the first compression device in the n compression devices is respectively connected to the first-stage ammonia separator and the ammonolysis reactor, and is used to remove the ammonia-containing gas in the first-stage ammonia separator Compressing the substance to obtain the supercritical fluid, and circulating the supercritical fluid to the ammonolysis reactor;
  • the second compression device in the n compression devices is respectively connected to the first-stage ammonia separator and the second-stage ammonia separator, and is used to circulate the second ammonia-containing gaseous substance to the first In the stage ammonia separator;
  • the i-th compression device of the n compression devices is respectively connected to the i-1-th stage ammonia separator and the i-th stage ammonia separator, and is used to circulate the i-th ammonia-containing gaseous substance to the i-th ammonia separator. -1 level ammonia separator.
  • the ammonia reaction system includes n ammonia separators and n compression devices arranged in sequence, and n is 3 or 4.
  • the ammonia source is at least one of an ammonia water mixture and liquid ammonia.
  • it further includes the step of adding an ammonia source to the ammonia separation device.
  • Another aspect of the present invention provides a method for preparing taurine, which includes the following steps:
  • the sodium isethionate and the ammonia source are placed in an ammonolysis reactor to perform an ammonolysis reaction to obtain a taurine intermediate-containing mixture, wherein the ammonia in the ammonia source and the hydroxyethyl
  • the molar ratio of sodium sulfonate is 25:1 or more.
  • the taurine intermediate taurine sodium is acidified to obtain taurine.
  • the molar ratio of ammonia in the ammonia source to the sodium isethionate is 25:1-100:1, more preferably 30:1-50:1.
  • the ammonia-containing gaseous substance is compressed by a compression device to obtain an ammonia-containing supercritical fluid, and the supercritical fluid is circulated to the ammonolysis reactor.
  • the taurine intermediate is subjected to the acidification treatment through a bipolar membrane to obtain the taurine and sodium hydroxide.
  • the sodium isethionate is obtained by reacting ethylene oxide with sodium bisulfite, and the sodium bisulfite is obtained through sulfur dioxide and at least partly from the bipolar membrane to the taurine.
  • the sodium hydroxide obtained by the acidification treatment of the acid intermediate is obtained by reacting.
  • the preparation method of the taurine intermediate taurine sodium and taurine has the following advantages: by increasing the ratio of the reactant ammonia source in the ammonolysis reaction, the ammonolysis reaction is fully carried out, thereby greatly improving the reaction efficiency. Yield.
  • the ammonia-containing gaseous substance can be obtained by separating the ammonia not participating in the reaction by an ammonia separation device.
  • the compression device compresses the ammonia-containing gaseous substance to obtain a supercritical fluid, and circulates the supercritical fluid to In the ammonolysis reactor, in this process, the full cycle of ammonia can be used with less energy consumption. In other words, the unreacted ammonia is recovered to participate in the ammonolysis reaction again, and finally the concentration of ammonia in the ammonolysis reaction is increased, and the cost is greatly reduced.
  • the supercritical fluid has a higher temperature and pressure. When the supercritical fluid is circulated to the ammonolysis reactor, the supercritical fluid can be directly The energy is coupled to the ammonolysis reactor to form the high temperature and high pressure conditions required in the ammonolysis reaction process, which saves energy.
  • Figure 1 is a schematic structural diagram of an ammonia separation device used in a preparation method of taurine intermediate taurine sodium in an embodiment of the present invention.
  • FIG. 2 is a schematic structural diagram of an ammonia separation device used in a preparation method of taurine intermediate taurine sodium in another embodiment of the present invention.
  • Fig. 3 is a schematic structural diagram of an ammonia separation device used in a preparation method of taurine intermediate taurine sodium in another embodiment of the present invention.
  • Figure 4 is a flow chart of the preparation method of taurine of the present invention.
  • Fig. 5 is a working schematic diagram of the acidification treatment in the preparation method of taurine of the present invention.
  • the invention provides a preparation method of taurine intermediate taurine sodium.
  • the preparation method includes the following steps:
  • S1 provide sodium isethionate and ammonia source
  • the ammonia in the ammonia source is used as an aminating agent.
  • the ammonia source is at least one of ammonia water mixture and liquid ammonia.
  • the mass fraction of ammonia in the ammonia water mixture is 20%-30%.
  • the reason why the molar ratio of ammonia in the ammonia source to the sodium isethionate is set to 25:1 or more is because it is considered that the higher the ratio of ammonia in the system, the more the positive reaction is promoted.
  • the yield of sulfonic acid intermediate taurate is higher.
  • the molar ratio of ammonia in the ammonia source to the sodium isethionate is 25:1-100:1.
  • the yield of the taurine intermediate sodium taurate is 85% or more.
  • the yield of the taurine intermediate sodium taurate is 95% or more.
  • the ammonolysis reactor serves as a vessel for the ammonolysis reaction.
  • the reaction temperature of the above-mentioned ammonolysis reaction may be 250° C. to 290° C.
  • the reaction pressure may be 10 MPa to 20 MPa
  • the reaction time may be 0.5 hour to 3.0 hours.
  • a raw material with a higher ammonia content can be directly provided, or the ammonia that is not involved in the reaction in the ammonolysis reaction can be recovered by the following step S3.
  • a raw material for the ammonolysis reaction As a raw material for the ammonolysis reaction.
  • S3 Separate the unreacted ammonia from the obtained mixture by an ammonia separation device to obtain ammonia-containing gaseous substance and taurine intermediate taurine sodium.
  • ammonia-containing gaseous substance can be directly passed into the ammonolysis reactor, or the following step S4 can be continued:
  • ammonia separation device and compression device will be described in further detail below.
  • the ammonia separation device may include a single ammonia separator, two ammonia separators, or multiple ammonia separators. They will be explained separately below.
  • the ammonia separation device may be a single-stage ammonia separator, that is, it includes a single ammonia separator 2.
  • the ammonia separator 2 is connected to the ammonolysis reactor 1.
  • the ammonia separator 2 is used to separate the ammonia that has not participated in the reaction after the ammonolysis reaction to obtain an ammonia-containing gaseous substance.
  • the compression device 3 is respectively connected to the ammonolysis reactor 1 and the ammonia separator 2. That is, the compression device 3 is located between the ammonolysis reactor 1 and the ammonia separator 2.
  • the compression device 3 is used to compress the ammonia-containing gaseous substance in the ammonolysis reactor 1 to obtain a supercritical fluid, and circulate the supercritical fluid to the ammonolysis reactor 1.
  • the ammonolysis reactor 1 can be a high-temperature and high-pressure reactor, which serves as a reaction site for preparing taurine intermediate sodium taurate.
  • the ammonolysis reactor 1 may be an autoclave, a tubular reactor or a synthesis tower, and preferably a tubular reactor.
  • the ammonia separator 2 may be a device that separates ammonia through evaporation or flash evaporation. Specifically, when the ammonia separator 2 is a flash evaporator, the flash evaporator may be pressurized during the flash evaporation process to achieve a better ammonia separation effect. When the ammonia-containing gaseous substance is discharged from the ammonia separator 2, it has a certain temperature and pressure. In other words, part of the energy in the mixture obtained after the ammonolysis reaction will be transferred to the ammonia-containing gaseous substance with a certain temperature and pressure, so as to more fully utilize the waste heat energy.
  • the compression device 3 may be a compressor, which is used to compress the ammonia-containing gaseous substance to obtain an ammonia-containing supercritical fluid.
  • the ammonia-containing gaseous material is exported to the compression device 3 through the ammonia separator 2, and the volume of the ammonia-containing gaseous material is reduced, and the internal energy is increased to obtain a supercritical fluid.
  • the supercritical fluid includes at least supercritical ammonia; the supercritical fluid also includes gaseous water and possibly supercritical water. Compared with the ammonia-containing gaseous substance, the supercritical fluid has a higher temperature and a higher pressure.
  • part of the work done by the compressor is converted into the gas molecules in the ammonia-containing gaseous state to overcome the intermolecular potential energy and into a supercritical fluid with a small particle distance, and the other part is converted into molecules
  • the kinetic energy is expressed as the supercritical fluid has a higher temperature and a higher pressure.
  • the supercritical fluid When the supercritical fluid is circulated to the ammonolysis reactor 1, it is preferably directly mixed with the sodium isethionate raw material in advance to obtain the mixture, and then passed into the ammonolysis reactor for reaction, which can achieve heating to promote the pretreatment of the raw materials.
  • the heat effect also increases the temperature and pressure in the ammonolysis reactor 1, providing high-pressure and high-heat reaction conditions for the ammonolysis reaction, which greatly saves energy.
  • the ammonia in the supercritical fluid can be used as a reaction raw material to increase the concentration of ammonia in the ammonolysis reaction, promote the full progress of the reaction, increase the reaction yield, reduce by-products, and save costs.
  • the ammonia separation device may be a two-stage ammonia separator, that is, it includes two ammonia separators: a first-stage ammonia separator 21 and a second-stage ammonia separator 22.
  • the first-stage ammonia separator 21 is connected to the ammonolysis reactor 1.
  • the second-stage ammonia separator 22 is connected to the first-stage ammonia separator 21.
  • the first-stage ammonia separator 21 is used to separate the ammonia that has not participated in the reaction from the mixture after the ammonolysis reaction to obtain the first ammonia-containing gaseous substance and the first remaining mixture.
  • the second-stage ammonia separator 22 is used to separate the first remaining mixture with ammonia to obtain a second ammonia-containing gaseous substance and a second residue, and to circulate the second ammonia-containing gaseous substance to the second One-stage ammonia separator 21.
  • the first ammonia-containing gaseous substance and the second ammonia-containing gaseous substance are mixed and transported to the first compression device 31; that is, the ammonia-containing gaseous substance in the first-stage ammonia separator 21 is the The sum of the first ammonia-containing gaseous substance and the second ammonia-containing gaseous substance.
  • This step-by-step reflux method is to gradually increase the pressure so that the load of the first compression device 31 during compression is not too large, and it is easier to compress the ammonia-containing gaseous substance into the supercritical State fluid.
  • the first compression device 31 is connected to the ammonolysis reactor 1 and the first-stage ammonia separator 21 respectively. That is, the first compression device 31 is located between the ammonolysis reactor 1 and the first-stage ammonia separator 21.
  • the first compression device 31 is the same as the compression device 3.
  • the first compression device 31 is used for compressing the ammonia-containing gas in the first-stage ammonia separator 21 to obtain the supercritical fluid, and circulate the supercritical fluid to the ammonia Solution reactor 1.
  • the second-stage ammonia separator 22 further performs ammonia separation on the first remaining mixture. This process is to further improve the recovery rate of ammonia and to maximize the utilization of waste heat energy.
  • an air pump or a second compression device 32 may be provided, preferably the second compression device 32.
  • the second compression device 32 can be set with different operating temperatures and pressures so as to give the second ammonia-containing gaseous substance a certain temperature and pressure, so as to facilitate the second ammonia-containing gaseous substance to enter the first-stage ammonia separator 21 After that, it passes through the first compression device 31 to become a supercritical fluid.
  • the ammonia separation device may be a three-stage ammonia separator, that is, it includes a first-stage ammonia separator 21, a second-stage ammonia separator 22 and a third separator 23.
  • the first-stage ammonia separator 21 is connected to the ammonolysis reactor 1.
  • the second-stage ammonia separator 22 is connected to the first-stage ammonia separator 21.
  • the third-stage ammonia separator 23 is connected to the second-stage ammonia separator 22.
  • the first-stage ammonia separator 21 is used to separate the ammonia that has not participated in the reaction from the mixture after the ammonolysis reaction to obtain the first ammonia-containing gaseous substance and the first remaining mixture.
  • the second-stage ammonia separator 22 is used to separate the first remaining mixture with ammonia to obtain a second ammonia-containing gaseous substance and a second residue, and to circulate the second ammonia-containing gaseous substance to the second One-stage ammonia separator 21.
  • the third-stage ammonia separator 23 is used to separate the second remaining mixture with ammonia to obtain a third ammonia-containing gaseous substance and a third residue, and to circulate the third ammonia-containing gaseous substance to the second stage
  • the ammonia separator 22 is continuously recycled to the first-stage ammonia separator 21.
  • the ammonia-containing gaseous substance present in the first-stage ammonia separator 21 includes the first ammonia-containing gaseous substance, the second ammonia-containing gaseous substance, and the third ammonia-containing gaseous substance, That is, the sum of the ammonia-containing gaseous substances separated by the various levels of ammonia separators.
  • the first compression device 31 is connected to the ammonolysis reactor 1 and the first-stage ammonia separator 21 respectively. That is, the first compression device 31 is located between the ammonolysis reactor 1 and the first-stage ammonia separator 21.
  • the first compression device 31 is the same as the compression device 3.
  • the first compression device 31 is used for compressing the ammonia-containing gas in the first-stage ammonia separator 21 to obtain the supercritical fluid, and circulate the supercritical fluid to the ammonia Solution reactor 1.
  • the second-stage ammonia separator 22 further separates ammonia from the first remaining mixture
  • the third-stage ammonia separator 23 further separates ammonia from the second remaining mixture.
  • the process is to further increase the recovery rate of ammonia and to maximize the utilization of waste heat energy. It can be understood that in order to enable the second ammonia-containing gaseous substance to smoothly enter the first-stage ammonia separator 21 and the third ammonia-containing gaseous substance to smoothly enter the second-stage ammonia separator 22, an air pump may be provided Or a compression device, such as the second compression device 32 and the third compression device 33.
  • the second compression device 32 and the third compression device 33 can be set with different operating temperatures and pressures, so as to give the second ammonia-containing gaseous substance and the third ammonia-containing gaseous substance a certain temperature and pressure, which is more conducive to the first After the three ammonia-containing gaseous substances enter the second-stage ammonia separator 22 and the first-stage ammonia separator 21 in sequence, and after the second ammonia-containing gaseous substances enter the first-stage ammonia separator 21, they all pass through the first The compression device 31 becomes a supercritical fluid.
  • the ammonia separation device is not limited to a two-stage ammonia separator and a three-stage ammonia separator, and may be a multi-stage ammonia separator.
  • the multi-stage ammonia separator can realize the step-by-step separation of ammonia.
  • the separated ammonia can be directly recycled to the ammonolysis reactor through the compression device to participate in the reaction, or reflux step by step, and finally pass through the first step.
  • a compression device 31 compresses into a supercritical fluid. At this time, the energy in the reacted mixture is also gradually recovered and accumulated. It can be expressed as follows:
  • the ammonia separation device includes n ammonia separators arranged in sequence, and n is an integer greater than 2 and less than 20, wherein,
  • the first-stage ammonia separator of the n ammonia separators arranged in sequence is connected to the ammonolysis reactor, and the first-stage ammonia separator is used to remove the impurities in the mixture obtained after the ammonolysis reaction.
  • the ammonia participating in the reaction is separated to obtain the first ammonia-containing gaseous substance and the first remaining mixture;
  • the second-stage ammonia separator among the n ammonia separators arranged in sequence is connected to the first-stage ammonia separator, and the second-stage ammonia separator is used to perform ammonia again on the first remaining mixture. Separating to obtain a second ammonia-containing gaseous substance and a second residue, and recycling the second ammonia-containing gaseous substance to the first-stage ammonia separator;
  • the i-th ammonia separator of the n ammonia separators arranged in sequence is connected to the i-1th ammonia separator, i is an integer and 2 ⁇ i ⁇ n, and the i-th ammonia separator is used for
  • the i-1th residue mixture obtained by the i-1th-stage ammonia separator is further subjected to ammonia separation to obtain the i-th ammonia-containing gaseous substance and the i-th residue, and the i-th ammonia-containing gaseous substance is recycled to The i-1 level ammonia separator.
  • the ammonolysis reaction system further includes n compression devices.
  • the first compression device in the n compression devices is respectively connected to the first-stage ammonia separator and the ammonolysis reactor, and is used to remove the ammonia-containing gas in the first-stage ammonia separator Compressing the substance to obtain the supercritical fluid, and circulating the supercritical fluid to the ammonolysis reactor;
  • the second compression device in the n compression devices is respectively connected to the first-stage ammonia separator and the second-stage ammonia separator, and is used to circulate the second ammonia-containing gaseous substance to the first In the stage ammonia separator;
  • the i-th compression device of the n compression devices is respectively connected to the i-1-th stage ammonia separator and the i-th stage ammonia separator, and is used to circulate the i-th ammonia-containing gaseous substance to the i-th ammonia separator. -1 level ammonia separator.
  • the ammonia reaction system includes n ammonia separators and n compression devices arranged in sequence, and n is 3 or 4.
  • the method further includes the step of adding an ammonia source to the ammonia separation device.
  • the purpose of the additional ammonia source passing through the ammonia separation device is that the added ammonia source passes through a single-stage ammonia separator or The ammonia separator returns to the compression device step by step, and finally is compressed into a supercritical fluid together with the recovered and separated ammonia.
  • the supplementary ammonia source before the supplementary ammonia source is passed into the ammonia separation device, it can exchange energy with the taurine intermediate sodium taurate mixture after ammonia separation through the heat exchanger 4 to The temperature of the added ammonia source is increased, and then it is passed to the ammonia separation device.
  • the present invention further provides a method for preparing taurine.
  • the preparation method includes the following steps:
  • S10 provide sodium isethionate and ammonia source
  • step S20 and before step S30 the following steps may be further included:
  • S201 Separate the ammonia that is not involved in the reaction from the obtained mixture by an ammonia separation device to obtain ammonia-containing gaseous substances and taurine intermediate taurine sodium;
  • S202 Compress the ammonia-containing gaseous substance by a compression device to obtain a supercritical fluid, and circulate the supercritical fluid to the ammonolysis reactor.
  • step S201 and step S202 please refer to the above-mentioned step S3 and step S4 respectively, which will not be repeated here.
  • the mass fraction of sodium taurate is 2%-30%, preferably 10%-25%.
  • step S30 the obtained taurine intermediate taurine sodium can be acidified through a bipolar membrane to obtain taurine and sodium hydroxide.
  • a three-compartment bipolar membrane electrodialysis device for acidification treatment.
  • the device is provided with an anode and a cathode, and a bipolar membrane (namely BP membrane) and a cation exchange membrane (namely C membrane) are alternately arranged between the anode and the cathode.
  • the solution containing taurine intermediate taurine sodium obtained in step S30 is passed into the material liquid chamber of the bipolar membrane electrodialysis device, and the water is passed into the lye chamber which is not in contact with the material liquid chamber.
  • the sodium aqueous solution is used as a conductive medium to pass into the cathode and anode compartments.
  • the sodium ions in the taurine sodium solution in the material liquid chamber pass through the cation exchange membrane into the lye chamber and combine with the hydroxide ions ionized by the water to form sodium hydroxide, and flow out from the lye chamber, and the water is ionized
  • the H ions pass through the bipolar membrane and combine with the taurine ion in the feed liquid chamber to form taurine, and finally taurine flows out from the feed liquid chamber.
  • the effluent taurine can be further concentrated and crystallized to obtain a taurine product.
  • the obtained crystallization mother liquor can be recycled to step S20 to perform the ammonolysis reaction.
  • the sodium hydroxide obtained in step S30 and the crystallization mother liquor obtained after the taurine crystallization can be recycled to step 20 for the ammonolysis reaction.
  • the sodium isethionate can be obtained by reacting ethylene oxide with sodium bisulfite, and the sodium bisulfite can be obtained by reacting sodium hydroxide and sulfur dioxide.
  • the sodium hydroxide obtained in step S30 can be used to prepare the sodium isethionate, thereby achieving recycling.
  • the lye may be sodium hydroxide solution.
  • the mass fraction of sodium hydroxide in the sodium hydroxide solution is 3%-30%.
  • the mass fraction of sodium hydroxide in the sodium hydroxide solution is 5%-20%.
  • the pH of the obtained sodium bisulfite solution is 3.5-7.0.
  • the pH of the sodium bisulfite solution is 4.0-6.5.
  • the pH of the solution containing sodium isethionate is 10.0 or more.
  • the pH of the solution containing sodium isethionate is 11.0 or more.
  • the mass fraction of the sodium isethionate in the solution containing sodium isethionate is 10%-20%.
  • This application adopts a bipolar membrane acidification method to produce taurine, which replaces the traditional sulfuric acid or hydrochloric acid acidification process, saves acid input, avoids the generation of by-product sodium sulfate or sodium chloride, and the sodium hydroxide produced can be Recycling greatly reduces the cost of raw materials and waste disposal. Because no inorganic salt is produced, the separation and purification process is simpler, and the equipment investment and production cost are reduced. The whole process has realized a closed loop, no three wastes are discharged, and it can be industrialized.
  • the ammonia-containing gaseous substance can be obtained by separating the ammonia not participating in the reaction by an ammonia separation device, and the ammonia-containing gaseous substance can be compressed by the compression device to obtain a supercritical fluid, and then the supercritical fluid Circulate to the ammonolysis reactor, in this process, realize the full circulation of ammonia with less energy consumption.
  • the unreacted ammonia is recovered to participate in the ammonolysis reaction again, and finally the concentration of ammonia in the ammonolysis reaction is increased, and the cost is greatly reduced.
  • the supercritical fluid has a higher temperature and pressure. When the supercritical fluid is circulated to the ammonolysis reactor, the supercritical fluid can be directly The energy is coupled to the ammonolysis reactor to form the high temperature and high pressure conditions required in the ammonolysis reaction process, which saves energy.
  • the mixture of ammonia water and liquid ammonia is mixed with sodium isethionate aqueous solution and pressurized by a high-pressure pump. After preheating, it is reacted in an ammonolysis reactor and processed by an evaporator to obtain a gas phase pressurized cycle and isethionate
  • the sodium aqueous solution is directly mixed and passed into the ammonolysis reactor for reaction.
  • the additional ammonia gas enters from the evaporator after heat exchange with the evaporation liquid. After stabilization, in the control system, the molar ratio of ammonia to sodium isethionate is 30:1.
  • the specific process conditions are: a 15% sodium isethionate aqueous solution with a mass fraction of 272Kg/h is pressurized to 18MPa by a high-pressure pump, and directly mixed with pressurized circulating ammonia, and the temperature is increased to 280°C. The mixture was passed into the ammonolysis reactor and reacted at 18MPa and 280°C for a residence time of 30min to obtain the ammonolysis reaction liquid. The ammonolysis reaction liquid is sent to an evaporator whose operating pressure is 0.1 MPa and operating temperature is 88.9°C.
  • the first ammonia-containing gaseous substance from the evaporator is compressed to 300°C and 18.2MPa by a compressor, and is recycled to the ammonolysis reactor.
  • the first liquid from the evaporator exchanges heat with the added ammonia to obtain 279Kg/h of sodium taurate solution.
  • the amount of added ammonia is 8.0Kg/h.
  • the content of sodium taurate is 13.7%
  • the content of sodium ditaurate is 1.1%
  • the content of sodium tritaurate is 0.09%
  • the yield of sodium taurate is calculated to be 94.3 %.
  • the unit production consumption is 2.16 tons of standard coal per ton of sodium taurate.
  • the mixture of ammonia water and liquid ammonia is mixed with sodium isethionate aqueous solution and pressurized by a high-pressure pump. After preheating, it is reacted by an ammonolysis reactor, and processed step by step through a flash tank and a second evaporator.
  • the gas phase obtained by the second-stage evaporation is pressurized and recycled to the first-stage flash tank.
  • the gas phase obtained by the first-stage flash is circulated under pressure and directly mixed with the passed-in sodium isethionate aqueous solution for heating, and passed into the ammonolysis reactor for reaction.
  • the additional ammonia gas enters from the secondary evaporator. After stabilization, in the control system, the molar ratio of ammonia to sodium isethionate is 30:1.
  • the specific process conditions are: a 15% sodium isethionate aqueous solution with a mass fraction of 272Kg/h is pressurized to 18MPa by a high-pressure pump, and directly mixed with pressurized circulating ammonia, and the temperature is increased to 280°C. The mixture was passed into the ammonolysis reactor and reacted at 18MPa and 280°C for a residence time of 30min to obtain the ammonolysis reaction liquid. The ammonolysis reaction liquid is sent to the first-stage flash tank for flash evaporation, the first-stage flash evaporation operating pressure is 8MPa, and the operating temperature is 220°C.
  • the first ammonia-containing gaseous substance from the first-stage flash tank is compressed to 300°C and 18.2MPa by a compressor, and is recycled to the ammonolysis reactor.
  • the first liquid from the primary flash tank enters the secondary evaporator.
  • the operating pressure of the secondary evaporator is 0.1MPa and the operating temperature is 87.8°C.
  • the second ammonia-containing gaseous substance obtained by the secondary evaporator is compressed to 210°C and 8.2MPa, and is recycled to the primary flash tank for flash evaporation.
  • the second liquid obtained by the secondary evaporator enters and exchanges heat with the supplemented ammonia to obtain 279Kg/h of sodium taurate solution.
  • the amount of added ammonia is 8Kg/h.
  • the sodium taurate content is 13.8%
  • the sodium ditaurate content is 1.09%
  • the sodium tritaurate content is 0.08%.
  • the yield of sodium taurate is calculated to be 94.9 %.
  • the unit production consumption is 1.07 tons of standard coal per ton of sodium taurate.
  • the mixture of ammonia water and liquid ammonia is mixed with sodium isethionate aqueous solution and pressurized by a high-pressure pump. After preheating, it is reacted in the ammonolysis reactor, and then passed through the first-stage flash tank, the second-stage flash tank, and the third-stage evaporation.
  • the device is processed step by step.
  • the gas phase obtained by the third-stage evaporation is pressurized and recycled to the second-stage flash tank.
  • the gas phase obtained by the secondary flash is pressurized and recycled to the primary flash tank.
  • the gas phase obtained by the first-stage flash is circulated under pressure and directly mixed with the passed-in sodium isethionate aqueous solution for heating, and passed into the ammonolysis reactor for reaction.
  • the additional ammonia gas enters from a three-stage evaporator. After stabilization, in the control system, the molar ratio of ammonia to sodium isethionate is 30:1.
  • the specific process conditions are: a 15% sodium isethionate aqueous solution with a mass fraction of 272Kg/h is pressurized to 18MPa by a high-pressure pump, and directly mixed with pressurized circulating ammonia, and the temperature is increased to 280°C. The mixture was passed into the ammonolysis reactor and reacted at 18MPa and 280°C for a residence time of 30min to obtain the ammonolysis reaction solution. The ammonolysis reaction liquid is sent to the first-stage flash tank for flash evaporation, the first-stage flash evaporation operating pressure is 8MPa, and the operating temperature is 245.5°C.
  • the first ammonia-containing gaseous substance from the first-stage flash tank is compressed to 300°C and 18.2MPa by a compressor, and is recycled to the ammonolysis reactor.
  • the first liquid from the primary flash tank enters the secondary flash tank for flash evaporation.
  • the operating pressure of the secondary flash tank is 3MPa and the operating temperature is 203.4°C.
  • the second ammonia-containing gaseous substance obtained in the secondary flash tank is compressed to 290°C and 8.2MPa, and is recycled to the primary flash tank for flash evaporation.
  • the second liquid obtained from the two-stage flash tank enters the three-stage evaporator whose operating pressure is 0.1MPa and the operating temperature is 97°C.
  • the third ammonia-containing gaseous substance obtained by the three-stage evaporator is compressed to 210°C, 3.1MPa, and then circulated to the second-stage flash tank for flash evaporation.
  • the third liquid obtained by the three-stage evaporator exchanges heat with the added ammonia to obtain 279Kg/h of taurine sodium solution.
  • the amount of added ammonia is 8Kg/h.
  • the content of each component in the sodium taurate solution was detected.
  • the sodium taurate content was 13.9%, the sodium ditaurate content was 1%, and the sodium tritaurate content was 0.08%.
  • the calculated yield of sodium taurate was 95.68. %.
  • the unit production consumption is 0.65 tons of standard coal per ton of sodium taurate.
  • the molar ratio of ammonia to sodium isethionate in the control system was 40:1.
  • the specific process conditions are: a 15% sodium isethionate aqueous solution with a mass fraction of 272Kg/h is pressurized to 18MPa by a high-pressure pump, and directly mixed with pressurized circulating ammonia, and the temperature is increased to 280°C.
  • the mixture was passed into the ammonolysis reactor and reacted at 18MPa and 280°C for a residence time of 30min to obtain the ammonolysis reaction liquid.
  • the ammonolysis reaction liquid is sent to the first-stage flash tank for flash evaporation, the first-stage flash evaporation operating pressure is 8MPa, and the operating temperature is 243°C.
  • the first ammonia-containing gaseous substance from the first-stage flash tank is compressed to 300°C and 18.2MPa by a compressor, and is recycled to the ammonolysis reactor.
  • the first liquid from the primary flash tank enters the secondary flash tank for flash evaporation.
  • the operating pressure of the secondary flash tank is 3MPa and the operating temperature is 203°C.
  • the second ammonia-containing gaseous substance obtained in the secondary flash tank is compressed to 290°C and 8.2MPa, and is recycled to the primary flash tank for flash evaporation.
  • the second liquid obtained from the two-stage flash tank enters the three-stage evaporator with an operating pressure of 0.1MPa and an operating temperature of 97°C.
  • the third ammonia-containing gaseous substance obtained by the three-stage evaporator is compressed to 210°C, 3.1MPa, and then circulated to the second-stage flash tank for flash evaporation.
  • the third liquid obtained by the three-stage evaporator exchanges heat with the added ammonia to obtain 279Kg/h of taurine sodium solution.
  • the amount of added ammonia is 8Kg/h.
  • the sodium taurate content is 14.1%
  • the sodium ditaurate content is 0.71%
  • the sodium tritaurate content is 0.05%.
  • the calculated yield of sodium taurate is 97.1 %.
  • the unit production consumption is 0.74 tons of standard coal per ton of sodium taurate.
  • the molar ratio of ammonia to sodium isethionate in the control system was 50:1.
  • the specific process conditions are: a 15% sodium isethionate aqueous solution with a mass fraction of 272Kg/h is pressurized to 18MPa by a high-pressure pump, and directly mixed with pressurized circulating ammonia, and the temperature is increased to 280°C.
  • the mixture was passed into the ammonolysis reactor and reacted at 18MPa and 280°C for a residence time of 30min to obtain the ammonolysis reaction liquid.
  • the ammonolysis reaction liquid is sent to the first-stage flash tank for flash evaporation, the first-stage flash evaporation operating pressure is 8MPa, and the operating temperature is 241°C.
  • the first ammonia-containing gaseous substance from the first-stage flash tank is compressed to 300°C and 18.2MPa by a compressor, and is recycled to the ammonolysis reactor.
  • the first liquid from the primary flash tank enters the secondary flash tank for flash evaporation.
  • the operating pressure of the secondary flash tank is 3MPa and the operating temperature is 203°C.
  • the second ammonia-containing gaseous substance obtained in the secondary flash tank is compressed to 290°C and 8.2MPa, and is recycled to the primary flash tank for flash evaporation.
  • the second liquid obtained from the two-stage flash tank enters the three-stage evaporator, the operating pressure of the evaporator is 0.1MPa, and the operating temperature is 97°C.
  • the third ammonia-containing gaseous substance obtained by the three-stage evaporator is compressed to 210°C, 3.1MPa, and then circulated to the second-stage flash tank for flash evaporation.
  • the third liquid obtained by the three-stage evaporator exchanges heat with the added ammonia to obtain 279Kg/h of taurine sodium solution.
  • the amount of added ammonia is 8Kg/h.
  • the sodium taurate content is 14.3%
  • the sodium ditaurate content is 0.65%
  • the sodium tritaurate content is 0.03%.
  • the calculated yield of sodium taurate is 98.43 %.
  • the unit production consumption is 0.80 tons of standard coal per ton of sodium taurate.
  • the caustic compartment obtains 6% lye, the material compartment obtains a taurine solution, the taurine solution is further concentrated to 45% concentration, crystallized to obtain a taurine product, the content is 99.4%, the total yield is 94% (including mother liquor Cycle yield).
  • the caustic compartment obtains 6% lye, the material compartment obtains taurine solution, the taurine solution is further concentrated to 45% concentration, crystallized to obtain taurine product, the content is 99.6%, the total yield is 94.5% (including mother liquor Cycle yield).
  • the molar ratio of sodium sulfonate is 35:1 to obtain a sodium taurate solution, which is filtered, diluted to a concentration of 10%, and enters the bipolar membrane electrodialysis system for acidification.
  • the caustic compartment obtains 6% lye
  • the material compartment obtains taurine solution
  • the taurine solution is further concentrated to 45% concentration
  • crystallized to obtain taurine product the content is 99.5%
  • the total yield is 95% (including mother liquor Cycle yield).
  • the caustic compartment obtains 6% lye
  • the material compartment obtains a taurine solution
  • the taurine solution is further concentrated to 45% concentration, and crystallized to obtain a taurine product with a content of 99.6% and a total yield of 96.2% (including mother liquor) Cycle yield).
  • the molar ratio of ammonia to sodium isethionate in the control system was 8:1.
  • the molar ratio of ammonia to sodium isethionate in the control system was 120:1.
  • the method of the prior art is adopted to carry out the ammonolysis reaction of sodium isethionate and ammonia and ammonia post-treatment.
  • 272Kg/h 15% sodium isethionate aqueous solution is mixed with a mixture of liquid ammonia and ammonia water, the molar ratio of ammonia to sodium isethionate is controlled to be 30:1, and the mixture flows through a high-pressure pump for pressure To 18MPa, preheat to 280°C, pass into the ammonolysis reactor, and react at 18MPa and 280°C for a residence time of 30min to obtain the ammonolysis reaction liquid.
  • the ammonolysis reaction liquid is sent to an evaporator for processing to obtain 278Kg/h of a sodium taurate solution with a content of 13.8%.
  • the operating pressure of the evaporator is 0.1MPa and the operating temperature is 97°C.
  • ammonia-containing substances obtained by evaporation are respectively condensed by the condenser, and then enters the ammonia distillation tower for recovery treatment.
  • the recovered ammonia is recycled to the ammonolysis reactor after adding fresh ammonia to participate in the reaction again.
  • the sodium taurate content is 13.7%
  • the sodium ditaurate content is 1.1%
  • the sodium tritaurate content is 0.1%
  • the yield of the sodium taurate is calculated to be 93.9 %.
  • the unit production consumption is 3.72 tons of standard coal per ton of sodium taurate.
  • the method of the prior art is adopted to carry out the ammonolysis reaction of sodium isethionate and ammonia and ammonia post-treatment.
  • 272Kg/h 15% sodium isethionate aqueous solution is mixed with a mixture of liquid ammonia and ammonia water, the molar ratio of ammonia to sodium isethionate is controlled to be 30:1, and the mixture flows through a high-pressure pump for pressure To 18MPa, preheat to 280°C, pass into the ammonolysis reactor and react at 18MPa and 280°C for a residence time of 30min to obtain the ammonolysis reaction liquid.
  • the ammonolysis reaction liquid was sent to a two-stage flash tank, and a three-stage evaporator was treated to obtain a sodium taurate solution of 278Kg/h.
  • the first-stage flash operation pressure was 8MPa, and the operating temperature was 241°C.
  • the second-stage flash tank The operating pressure is 3MPa, the operating temperature is 203°C, the operating pressure of the three-stage evaporator is 0.1MPa, and the operating temperature is 97°C.
  • the ammonia-containing substances obtained by flashing and evaporation at various levels are respectively condensed by the condenser, and then enters the ammonia distillation tower for recovery treatment. The ammonia recovered at the top of the tower is added with fresh ammonia and then recycled to the ammonolysis reactor to participate in the reaction again.
  • the content of each component in the sodium taurate solution was detected.
  • the sodium taurate content was 13.8%
  • the sodium ditaurate content was 1.2%
  • the sodium tritaurate content was 0.1%.
  • the calculated yield of sodium taurate was 94.59 %.
  • the unit production consumption is 1.32 tons of standard coal per ton of sodium taurate.
  • Comparative example 1 uses low ammonia ratio to carry out the ammonolysis reaction, the yield is only 64.85%, and the content of sodium ditaurate and sodium tritaurate in the ammonolysis product is relatively high; Comparative example 2 uses an ammonia ratio exceeding 100, The yield is as high as 98.7%, but the subsequent ammonia recovery requires a higher cost, and the increase in the yield is not large compared to the yield of Example 8. Comparative Examples 3 and 4 did not use the compression device of the present invention, nor did the recovered ammonia become supercritical fluid for cyclic operation. Instead, the ammonia-containing gas phase obtained by the ammonia separator was directly processed through the ammonia distillation tower to recover the recovered ammonia.

Abstract

本发明提供一种牛磺酸中间体牛磺酸钠的制备方法，包括以下步骤：提供羟乙基磺酸钠和氨源；将所述羟乙基磺酸钠和所述氨源置于氨解反应器中进行氨解反应，得到含牛磺酸中间体牛磺酸钠的混合物，其中，所述氨源中的氨与所述羟乙基磺酸钠的摩尔比为25:1以上。本发明还提供一种牛磺酸的制备方法。

Description

牛磺酸中间体牛磺酸钠的制备方法及牛磺酸的制备方法
相关申请
本申请要求2019年5月30日申请的，申请号为201910463650.2，名称为“牛磺酸中间体牛磺酸钠的制备方法及牛磺酸的制备方法”的中国专利申请的优先权，在此将其全文引入作为参考。
技术领域
本发明涉及牛磺酸的制备技术领域，特别是涉及氨解反应系统、牛磺酸中间体牛磺酸钠的制备方法及牛磺酸的制备方法。
背景技术
牛磺酸(2-氨基乙磺酸)，又名牛胆酸、牛胆素，呈白色结晶或粉末，无臭、无毒、微酸味。它是一种非蛋白质类氨基酸，是人体必需的重要氨基酸之一，具有独特的药理及营养保健作用。牛磺酸可广泛应用于医药、食品添加剂、荧光增白剂、有机合成等领域，也可用作生化试剂、湿润剂、缓冲剂等。西方发达国家已普遍将牛磺酸应用于医药及食品添加剂中。
牛磺酸的制备方法主要有生物提取法、发酵法和化学合成方法，其中化学合成方法的研究最为迅速。牛磺酸的化学合成方法根据原料和工艺不同，目前有20多种。但由于受到原料来源、生产成本、产品收率及合成工艺条件和设备要求等限制，真正能用于工业化生产的有两种方法：
(1)乙醇胺法：以乙醇胺为原料，两步合成牛磺酸，按合成路线又可分为酯化法、氯化法、乙撑亚胺法。其中酯化法原料易得，收率较其他方法高，为国内外多数厂家采用，所述反应以乙醇胺、硫酸、亚硫酸钠为原料，首先硫酸与乙醇胺进行酯化反应合成中间体2-氨基乙基硫酸酯，再与亚硫酸钠或亚硫酸铵进行磺化反应合成牛磺酸。反应方程式如下：
Figure PCTCN2020092547-appb-000001
NH 2CH 2CH 2OSO 3H+Na 2SO 3→NH 2CH 2CH 2SO 3H+Na 2SO 4，
或
NH 2CH 2CH 2OSO 3H+(NH 4) 2SO 3→NH 2CH 2CH 2SO 3H+(NH 4) 2SO 4。
但其中酯化反应为可逆反应，反应不完全，制约着乙醇胺的转化率及反应收率，且反应体系有硫酸钠生成，易造成分离困难，影响产品收率和质量，环保压力大。
(2)环氧乙烷法：以环氧乙烷为原料，先与亚硫酸钠开环加成，然后在加热加压条件下与氨反应合成牛磺酸钠，酸化得到牛磺酸。其反应过程如下：
①
Figure PCTCN2020092547-appb-000002
②HOCH 2CH 2SO 3Na+NH 3→H 2NCH 2CH 2SO 3Na+H 2O
③H 2NCH 2CH 2SO 3Na+H 2SO 4→H 2NCH 2CH 2SO 3H+Na 2SO 4
副反应：
Figure PCTCN2020092547-appb-000003
2HOCH 2CH 2SO 3Na+NH 3→HN(CH 2CH 2SO 3Na) 2+H 2O
3HOCH 2CH 2SO 3Na+NH 3→N(CH 2CH 2SO 3Na) 3+H 2O。
环氧乙烷法包括加成、氨解、酸化步骤，其收率较乙醇胺法的高，目前应用较广。
环氧乙烷法的氨解和酸化步骤是环氧乙烷法制备牛磺酸工艺的关键影响步骤。在专利US1932907中提到羟乙基磺酸盐与胺类物质的氨解反应，其中氨与羟乙基磺酸盐的摩尔比在6.8:1，反应温度240℃～250℃下反应2h时，得到牛磺酸钠的产率仅为80％。专利DD219023A3中提及羟乙基磺酸钠氨解产物的组成，当氨与羟乙基磺酸钠的摩尔比在(10～20):1时，并加入碱金属或碱金属氢氧化物作为催化剂，在200～290℃下反应5～45分钟，得到含有71％牛磺酸钠和29％二牛磺酸钠及三牛磺酸钠的氨解产物，但收率最高仅有64％。可知，在羟乙基磺酸钠氨解牛磺酸钠时，易生成副产物二牛磺酸盐、三牛磺酸盐。针对氨解反应，羟乙基磺酸钠与氨的反应为热效应不明显的可逆反应，采用的氨虽为过量状态，但氨与羟乙基磺酸钠的摩尔比低，且氨在液相中具有一定的溶解度，在反应时液相中溶解的氨量远低于设定的氨/羟乙基磺酸钠的值，如此导致大量副反应进行，易生成副产物二牛磺酸盐和三牛磺酸盐，从而造成牛磺酸钠的低收率。为了提高氨解收率，有人采取了一些研究，如专利CN105732440、CN108314633均是将氨解反应液经酸中和后分离得到的母液全部或者大部分循环至氨解，加入的母液越多，氨解反应收率越高。上述文献均提到将母液循环至氨解继续反应，收率上有了大幅的提高，但母液中除了含有副产物二牛磺酸盐和三牛磺酸盐外，还含有硫酸钠、乙二醇、聚乙二醇、微量金属元素等多种复杂成分，未经处理的母液循环至体系中时，随着循环次数增加，体系中的杂质大量聚集，不利于反应进行，如果直接排放则为高浓度污染物，对环境的影响非常大，且母液循环至氨解时，需要补加氨量，为达到氨解的高温高压条件，需要对母液及补加的氨重新进行加热加压，需要的热量大幅增加，不利于工业化生产。
牛磺酸制备过程中，氨解反应通常以氨过量的形式进行，在氨解完成后需要进行脱氨处理，而通常氨解液拥有较高的温度和压力，在脱氨处理过程中形成的含氨气相仍具有一定的热量，针对这部分热量的回收利用问题，有人作了一部分研究，如专利CN101528658中公开了氨解液的处理方式，将氨解液分别经一级闪蒸、二级闪蒸降膜蒸发、多效降膜蒸发浓缩进行处理，利用闪蒸汽作为下一级蒸发器的热源介质进行加热，但专利中并未提及回收的氨后续如何处理的问题。对于脱除的氨循环利用问题，普遍的做法是将高含量的氨冷凝循环回氨解步骤，低含量的氨冷凝后使用诸如蒸氨塔等设备进行精制，达到一定浓度后再循环利用，但经过处理的氨循环回氨解步骤时，为了达到氨解的高温高压反应条件，需加压加热，同样需要消耗大量的能源，未涉及如何将氨以低能耗的方式循环利用。
此外，针对牛磺酸钠的酸化过程，也有人采用硫酸、盐酸等试剂进行处理，如专利US9061976、CN101486669、CN101508657均是采用硫酸或者亚硫酸进行酸化。采用硫酸酸化易产生硫酸钠等大量无机盐分，造成分离困难、设备堵塞、生产成本高等问题。
发明内容
本发明一方面提供一种牛磺酸中间体牛磺酸钠的制备方法，所述制备方法包括以下步骤：
提供羟乙基磺酸钠和氨源；
将所述羟乙基磺酸钠和所述氨源置于氨解反应器中进行氨解反应，得到含牛磺酸中间体牛磺酸钠的混合物，其中，所述氨源中的氨与所述羟乙基磺酸钠的摩尔比为25:1以上。
在一些实施例中，所述氨源中的氨与所述羟乙基磺酸钠的摩尔比为25:1-100:1，进一步优选30:1-50:1。
在一些实施例中，在所述氨解反应的步骤之后还包括以下步骤：
通过氨分离装置将所述混合物中未参与反应的氨分离出来而分别得到含氨气态物以及牛磺酸中间体，其中所述氨分离装置与所述氨解反应器相连接；
通过压缩装置对所述含氨气态物进行压缩，得到含氨的超临界态流体，并将所述超临界态流体循环至所述氨解反应器，其中所述压缩装置分别与所述氨分离装置、所述氨解反应器相连接。
在一些实施例中，所述氨分离装置包括一个氨分离器。
在一些实施例中，所述氨分离装置包括两个氨分离器，分别为第一级氨分离器以及第二级氨分离器，
所述第一级氨分离器与所述氨解反应器相连接，所述第一级氨分离器用于将所述氨解反应后的混合物中未参与反应的氨分离出来，得到第一含氨气态物以及第一剩余混合物；
所述第二级氨分离器与所述第一级氨分离器相连接，所述第二级氨分离器用于对所述第一剩余混合物再进行氨气分离，得到第二含氨气态物以及第二剩余物，并将所述第二含氨气态物循环至所述第一级氨分离器中。
在一些实施例中，所述压缩装置包括第一压缩装置和第二压缩装置，
所述第一压缩装置分别与所述第一级氨分离器、所述氨解反应器相连接，用于将所述第一级氨分离器中的所述含氨气态物进行压缩得到所述超临界态流体，并将所述超临界态流体循环至所述氨解反应器；
所述第二压缩装置分别与所述第一级氨分离器和所述第二级氨分离器连接，用于将所述第二含氨气态物循环至所述第一级氨分离器中。
在一些实施例中，所述氨分离装置包括n个依次排列的氨分离器，n为大于2，小于20的整数，
所述依次排列的n个氨分离器中的第一级氨分离器与所述氨解反应器相连接，所述第一级氨分离器用于将所述氨解反应后所得到的混合物中未参与反应的氨分离出来，得到第一含氨气态物以及第一剩余混合物；
所述依次排列的n个氨分离器中的第二级氨分离器与所述第一级氨分离器相连接，所述第二级氨分离器用于对所述第一剩余混合物再进行氨气分离，得到第二含氨气态物以及第二剩余物，并将所述第二含氨气态物循环至所述第一级氨分离器中；
所述依次排列的n个氨分离器中的第i级氨分离器与第i-1级氨分离器相连接，i为整数且2<i≤n，所述第i级氨分离器用于对通过所述第i-1级氨分离器得到的第i-1剩余混合物再进行氨气分离，得到第i含氨气态物以及第i剩余物，并将所述第i含氨气态物循环至第i-1级氨分离器。
在一些实施例中，所述氨解反应系统还包括n个压缩装置，
所述n个压缩装置中的第一压缩装置分别与所述第一级氨分离器、所述氨解反应器相连接，用于将所述第一级氨分离器中的所述含氨气态物进行压缩得到所述超临界态流体，并将所述超临界态流体循环至所述氨解反应器；
所述n个压缩装置中的第二压缩装置分别与所述第一级氨分离器和所述第二级氨分离器连接，用于将所述第二含氨气态物循环至所述第一级氨分离器中；
所述n个压缩装置中的第i压缩装置分别与所述第i-1级氨分离器和所述第i级氨分离器连接，用于将所述第i含氨气态物循环至第i-1级氨分离器。
在一些实施例中，所述氨反应系统中包括n个依次排列的氨分离器，n个压缩装置，n为3或4。
在一些实施例中，所述氨源为氨水混合物、液氨中的至少一种。
在一些实施例中，还包括向所述氨分离装置补加氨源的步骤。
本发明另一方面提供一种牛磺酸的制备方法，所述制备方法包括以下步骤：
提供羟乙基磺酸钠和氨源；
将所述羟乙基磺酸钠和所述氨源置于氨解反应器中进行氨解反应，得到含牛磺酸中间体的混合物，其中所述氨源中的氨与所述羟乙基磺酸钠的摩尔比为25:1以上。
将所述牛磺酸中间体牛磺酸钠进行酸化处理，得到牛磺酸。
在一些实施例中，所述氨源中的氨与所述羟乙基磺酸钠的摩尔比为25:1-100:1，进一步优选30:1-50:1。
在一些实施例中，在所述氨解反应的步骤之后，对所述牛磺酸中间体进行酸化处理的步骤之前还包括以下步骤：
通过氨分离装置将所述混合物中未参与反应的氨分离出来而分别得到含氨气态物以及牛磺酸中间体；
通过压缩装置对所述含氨气态物进行压缩，得到含氨的超临界态流体，并将所述超临界态流体循环至所述氨解反应器。
在一些实施例中，通过双极膜对所述牛磺酸中间体进行所述酸化处理，得到所述牛磺酸和氢氧化钠。
在一些实施例中，所述羟乙基磺酸钠通过环氧乙烷与亚硫酸氢钠反应得到，所述亚硫酸氢钠通过 二氧化硫与至少部分来自所述通过双极膜对所述牛磺酸中间体进行的所述酸化处理得到的所述氢氧化钠进行反应得到。
所述牛磺酸中间体牛磺酸钠及牛磺酸的制备方法具有以下优点：通过将氨解反应中反应物氨源的比例提高，从而使得氨解反应充分进行，进而大大提高了反应的收率。
进一步的，可通过氨分离装置将未参与反应的氨分离出来而得到含氨气态物，所述压缩装置将含氨气态物进行压缩得到超临界态流体，并将所述超临界态流体循环至氨解反应器，在此过程中，以较小的能耗实现氨的全循环套用。也就是说，将未反应的氨回收再次参与所述氨解反应，最终提高了氨解反应中氨的浓度，大大降低了成本。另外，在将含氨气态物转换成所述超临界态流体后，所述超临界态流体具有较高的温度及压强，当所述超临界态流体循环至氨解反应器时，可直接将能量耦合至氨解反应器中，从而形成氨解反应过程中所需的高温高压条件，节约了能源。
附图说明
为了更好地描述和说明这里公开的那些发明的实施例和/或示例，可以参考一幅或多幅附图。用于描述附图的附加细节或示例不应当被认为是对所公开的发明、目前描述的实施例和/或示例以及目前理解的这些发明的最佳模式中的热河一者的范围的限制。
图1为本发明一实施例牛磺酸中间体牛磺酸钠的制备方法采用的氨分离装置的结构示意图。
图2为本发明另一实施例牛磺酸中间体牛磺酸钠的制备方法采用的氨分离装置的结构示意图。
图3为本发明另一实施例牛磺酸中间体牛磺酸钠的制备方法采用的氨分离装置的结构示意图。
图4为本发明牛磺酸的制备方法的流程图。
图5为本发明牛磺酸的制备方法中酸化处理的工作示意图。
具体实施方式
下面将对本发明实施方式中的技术方案进行清楚、完整地描述，显然，所描述的实施方式仅仅是本发明一部分实施方式，而不是全部的实施方式。基于本发明中的实施方式，本领域普通技术人员在没有作出创造性劳动前提下所获得的所有其它实施方式，都属于本发明保护的范围。
本发明提供一种牛磺酸中间体牛磺酸钠的制备方法。所述制备方法包括以下步骤：
S1，提供羟乙基磺酸钠和氨源；
S2，将所述羟乙基磺酸钠和所述氨源置于氨解反应器中进行氨解反应，得到混合物含牛磺酸中间体牛磺酸钠的混合物，其中，所述氨源中的氨与所述羟乙基磺酸钠的摩尔比为25:1以上。
在步骤S1中，所述氨源中的氨作为胺化剂。所述氨源为氨水混合物、液氨中的至少一种。所述氨水混合物中氨的质量分数为20％～30％。之所以将所述氨源中的氨与所述羟乙基磺酸钠的摩尔比设为25:1以上，是因为考虑到氨在体系中的比例越高，越促进正反应方向进行，牛磺酸中间体牛磺酸钠的收率越高。优选的，所述氨源中的氨与所述羟乙基磺酸钠的摩尔比为25:1～100:1。举例而言，当所述氨源中的氨与所述羟乙基磺酸钠的摩尔比设为25:1时，牛磺酸中间体牛磺酸钠的收率为85％以上。当所述氨源中的氨与所述羟乙基磺酸钠的摩尔比设为100:1时，牛磺酸中间体牛磺酸钠的收率为95％以上。
在步骤S2中，所述氨解反应器作为氨解反应的容器。上述氨解反应的反应温度可为250℃～290℃，反应压力可为10MPa～20MPa，反应时间可为0.5小时～3.0小时。
进一步的，为了实现步骤S1中所述氨源的氨的比例较高，可直接提供氨的含量较高的原料，也可以通过以下步骤S3将所述氨解反应中未参与反应的氨回收以作为所述氨解反应的原料。
S3，通过氨分离装置将所得到的混合物中未参与反应的氨分离出来而分别得到含氨气态物以及牛磺酸中间体牛磺酸钠。
其中，该含氨气态物可直接通入氨解反应器中，也可继续进行以下步骤S4：
S4，通过所述压缩装置对所述含氨气态物进行压缩，得到超临界态流体，并将所述超临界态流体循环至氨解反应器。
以下对所述氨分离装置、压缩装置进行进一步的详细说明。
所述氨分离装置可包括单个氨分离器、两个氨分离器或者多个氨分离器。以下将分别说明。
请参阅图1，所述氨分离装置可为单级氨分离器，即包括单个氨分离器2。所述氨分离器2连接于所述氨解反应器1。所述氨分离器2用于将所述氨解反应后未参与反应的氨分离出来而得到含氨气态物。所述压缩装置3分别与所述氨解反应器1、所述氨分离器2相连接。即，所述压缩装置3位于所述氨解反应器1与所述氨分离器2之间。所述压缩装置3用于将所述氨解反应器1中的所述含氨气态物进行压缩得到超临界态流体，并将所述超临界态流体循环至所述氨解反应器1。
所述氨解反应器1可为高温高压反应器，其作为制备牛磺酸中间体牛磺酸钠的反应场所。具体的，所述氨解反应器1可为高压釜、管式反应器或合成塔，优选为管式反应器。
所述氨分离器2可为通过蒸发或者闪蒸的方式进行氨分离的装置。具体的，当所述氨分离器2为闪蒸器时，所述闪蒸器可在闪蒸的过程中带压，以实现更好的氨分离效果。所述含氨气态物自所述氨分离器2导出时，会自身具有一定的温度及压力。换句话说，氨解反应后所得的混合物中的部分能量会转移至所述具有一定的温度及压力的所述含氨气态物，以便更充分的利用余热能量。
所述压缩装置3可为压缩机，其用于将所述含氨气态物进行压缩得到含氨的超临界态流体。在这一过程中，所述含氨气态物经过所述氨分离器2导出至压缩装置3，所述含氨气态物的体积会减小，内能增加，而得到超临界态流体。需要说明的是，所述超临界态流体至少包括超临界态的氨；所述超临界态流体还包括气态的水以及可能存在的超临界态的水。相比所述含氨气态物而言，所述超临界态流体具有较高的温度以及较高的压强。此过程中，可理解为：压缩机所做的功，一部分转化成所述含氨气态物中气体分子克服分子间的势能而转成分子间距较小的超临界态流体，另一部分转化成分子的动能，即表现为所述超临界态流体具有较高的温度以及较高的压强。
当所述超临界态流体循环至所述氨解反应器1时，优选与羟乙基磺酸钠原料预先直接混合，得到混合物，再通入氨解反应器中反应，可达到加热促使原料预热的效果，并同时使氨解反应器1内的温度及压力升高，而为氨解反应提供了高压高热的反应条件，大大节约了能源。并且，所述超临界态流体中的氨可作为反应原料，而提升氨解反应中氨的浓度，促使反应充分进行，提高反应的产率，降低了副产物，节约了成本。
请参阅图2，所述氨分离装置可为两级氨分离器，即包括两个氨分离器：第一级氨分离器21及第二级氨分离器22。所述第一级氨分离器21连接于所述氨解反应器1。所述第二级氨分离器22连接于所述第一级氨分离器21。所述第一级氨分离器21用于将所述氨解反应后的混合物中未参与反应的氨分离出来，得到第一含氨气态物以及第一剩余混合物。所述第二级氨分离器22用于对所述第一剩余混合物再进行氨气分离，得到第二含氨气态物以及第二剩余物，并将所述第二含氨气态物循环至第一级氨分离器21。所述第一含氨气态物和第二含氨气态物混合，而运送至所述第一压缩装置31；即，所述第一级氨分离器21中的所述含氨气态物为所述第一含氨气态物和第二含氨气态物的总和。这种逐级回流的方式，是为了逐渐增大压力，使得所述第一压缩装置31在进行压缩时的负荷不致过大，而较容易的将所述含氨气态物压缩成所述超临界态流体。
所述第一压缩装置31分别与所述氨解反应器1、所述第一级氨分离器21相连接。即，所述第一压缩装置31位于所述氨解反应器1与所述第一级氨分离器21之间。所述第一压缩装置31与所述压缩装置3相同。所述第一压缩装置31用于将所述第一级氨分离器21中的所述含氨气态物进行压缩得到所述超临界态流体，并将所述超临界态流体循环至所述氨解反应器1。
该实施例中，所述第二级氨分离器22对所述第一剩余混合物再进行氨气分离，此过程是为了进一步提升氨的回收率，也最大程度的实现余热能源利用。可以理解，为了使所述第二含氨气态物顺利的进入所述第一级氨分离器21，可设置气泵或者第二压缩装置32，优先为第二压缩装置32。所述第二压缩装置32可设置不同的操作温度及压力，以便赋予所述第二含氨气态物具有一定的温度及压力，从而更利于第二含氨气态物进入第一级氨分离器21之后，再通过第一压缩装置31变成超临界态流体。
请参阅图3，所述氨分离装置可为三级氨分离器，即包括第一级氨分离器21、第二级氨分离器22及第三分离器23。其中，所述第一级氨分离器21连接于所述氨解反应器1。所述第二级氨分离器22连接于所述第一级氨分离器21。所述第三级氨分离器23连接于所述第二级氨分离器22。所述第一级氨分离器21用于将所述氨解反应后的混合物中未参与反应的氨分离出来，得到第一含氨气态物以及第一剩余混合物。所述第二级氨分离器22用于对所述第一剩余混合物再进行氨气分离，得到第二含氨气态物以及第二 剩余物，并将所述第二含氨气态物循环至第一级氨分离器21。第三级氨分离器23用于对所述第二剩余混合物再进行氨气分离，得到第三含氨气态物以及第三剩余物，并将所述第三含氨气态物循环至第二级氨分离器22，并继续循环至所述第一级氨分离器21。此时，所述第一级氨分离器21中所存在的所述含氨气态物包括所述第一含氨气态物、所述第二含氨气态物以及所述第三含氨气态物，即是各级氨分离器的所分离得到的含氨气态物之和。
所述第一压缩装置31分别与所述氨解反应器1、所述第一级氨分离器21相连接。即，所述第一压缩装置31位于所述氨解反应器1与所述第一级氨分离器21之间。所述第一压缩装置31与所述压缩装置3相同。所述第一压缩装置31用于将所述第一级氨分离器21中的所述含氨气态物进行压缩得到所述超临界态流体，并将所述超临界态流体循环至所述氨解反应器1。
该实施例中，所述第二级氨分离器22对所述第一剩余混合物再进行氨气分离，所述第三级氨分离器23对所述第二剩余混合物再进行氨气分离，此过程是为了进一步提升氨的回收率，也最大程度的实现余热能源利用。可以理解，为了使所述第二含氨气态物顺利的进入所述第一级氨分离器21、所述第三含氨气态物顺利的进入所述第二级氨分离器22，可设置气泵或者压缩装置，如第二压缩装置32、第三压缩装置33。所述第二压缩装置32及第三压缩装置33可设置不同的操作温度及压力，以便赋予所述第二含氨气态物、第三含氨气态物具有一定的温度及压力，从而更利于第三含氨气态物依次进入所述第二级氨分离器22、第一级氨分离器21之后，以及第二含氨气态物进入所述第一级氨分离器21之后，均再通过第一压缩装置31变成超临界态流体。
以此类推，所述氨分离装置不限于两级氨分离器、三级氨分离器，可为多级氨分离器。换句话说，所述多级氨分离器可实现逐级的氨气分离，分离后的氨气可分别直接通过压缩装置循环至氨解反应器中参与反应，或者逐级回流，最终均通过第一压缩装置31压缩变成超临界流体，此时，反应后的混合物中的能量也逐级回收聚集。可表示如下：
所述氨分离装置包括n个依次排列的氨分离器，n为大于2，小于20的整数，其中，
所述依次排列的n个氨分离器中的第一级氨分离器与所述氨解反应器相连接，所述第一级氨分离器用于将所述氨解反应后所得到的混合物中未参与反应的氨分离出来，得到第一含氨气态物以及第一剩余混合物；
所述依次排列的n个氨分离器中的第二级氨分离器与所述第一级氨分离器相连接，所述第二级氨分离器用于对所述第一剩余混合物再进行氨气分离，得到第二含氨气态物以及第二剩余物，并将所述第二含氨气态物循环至所述第一级氨分离器中；
所述依次排列的n个氨分离器中的第i级氨分离器与第i-1级氨分离器相连接，i为整数且2<i≤n，所述第i级氨分离器用于对通过所述第i-1级氨分离器得到的第i-1剩余混合物再进行氨气分离，得到第i含氨气态物以及第i剩余物，并将所述第i含氨气态物循环至第i-1级氨分离器。
此时，所述氨解反应系统还包括n个压缩装置。所述n个压缩装置中的第一压缩装置分别与所述第一级氨分离器、所述氨解反应器相连接，用于将所述第一级氨分离器中的所述含氨气态物进行压缩得到所述超临界态流体，并将所述超临界态流体循环至所述氨解反应器；
所述n个压缩装置中的第二压缩装置分别与所述第一级氨分离器和所述第二级氨分离器连接，用于将所述第二含氨气态物循环至所述第一级氨分离器中；
所述n个压缩装置中的第i压缩装置分别与所述第i-1级氨分离器和所述第i级氨分离器连接，用于将所述第i含氨气态物循环至第i-1级氨分离器。
优选的，所述氨反应系统中包括n个依次排列的氨分离器，n个压缩装置，n为3或4。
进一步，本方法还包括向所述氨分离装置补加氨源的步骤，所述补加氨源需要通过氨分离装置的目的在于，所述补加的氨源经由单级氨分离器或者各级氨分离器逐级回流至压缩装置，最后和回收分离的氨一同被压缩成超临界流体。优选的，所述补加的氨源在通入至所述氨分离装置之前，可先与分离氨之后的含牛磺酸中间体牛磺酸钠的混合物通过换热器4进行换能，以提高所述补加的氨源的温度，再通入至氨分离装置。
请参阅图4，本发明进一步提供一种牛磺酸的制备方法。所述制备方法包括以下步骤：
S10，提供羟乙基磺酸钠和氨源；
S20，将所述羟乙基磺酸钠和所述氨源置于氨解反应器中进行氨解反应，得到含牛磺酸中间体牛磺酸钠的混合物，其中所述氨源中的氨与所述羟乙基磺酸钠的摩尔比为25:1以上；
S30，将得到的牛磺酸中间体牛磺酸钠进行酸化处理，得到牛磺酸。
在步骤S20之后，步骤S30之前，还可进一步包括以下步骤：
S201，通过氨分离装置将所得到的混合物中未参与反应的氨分离出来而分别得到含氨气态物以及牛磺酸中间体牛磺酸钠；
S202，通过压缩装置对所述含氨气态物进行压缩，得到超临界态流体，并将所述超临界态流体循环至所述氨解反应器。
其中步骤S201、步骤S202，可分别参照上述步骤S3、步骤S4，在此不再赘述。其中由步骤S20得到的混合物经过氨分离装置的处理后，牛磺酸钠的质量分数为2％～30％，优选为10％～25％。
在步骤S30中，可通过双极膜对得到的牛磺酸中间体牛磺酸钠进行酸化处理，得到牛磺酸和氢氧化钠。
请参阅图5，具体可采用三隔室双极膜电渗析装置进行酸化处理。所述装置设有阳极、阴极，在阳极、阴极之间分别交替设置双极膜(即BP膜)、阳离子交换膜(即C膜)。步骤S30所得到的含牛磺酸中间体牛磺酸钠的溶液通入所述双极膜电渗析装置的料液室中，水通入与料液室不接触的碱液室中，氢氧化钠水溶液作为导电介质通入阴极室、阳极室。在电场作用下，料液室中牛磺酸钠溶液中的钠离子穿过阳离子交换膜进入碱液室与水电离出来的氢氧根离子结合成氢氧化钠，从碱液室流出，水电离的H离子穿过双极膜与料液室中的牛磺酸根离子结合成牛磺酸，最后牛磺酸从料液室流出。
所述流出的牛磺酸可进一步进行浓缩结晶，得到牛磺酸产品。在所述牛磺酸的浓缩结晶后，所得到的的结晶母液可循环至步骤S20再进行氨解反应。当然，考虑到碱会对氨解反应会有催化作用，可将步骤S30得到的氢氧化钠与所述牛磺酸结晶后得到的结晶母液一同循环至步骤20进行氨解反应。
可以理解，在步骤S10之前，所述羟乙基磺酸钠可通过环氧乙烷与亚硫酸氢钠反应得到，所述亚硫酸氢钠可通过氢氧化钠和二氧化硫反应得到。此时，可利用步骤S30所得的氢氧化钠以制备所述羟乙基磺酸钠，从而实现循环利用。具体的，
(1)将二氧化硫通入碱液中，得到亚硫酸氢钠溶液；
(2)提供环氧乙烷，将环氧乙烷与得到的亚硫酸氢钠溶液进行加成反应，生成含羟乙基磺酸钠的溶液。
其中，步骤(1)中，所述碱液可为氢氧化钠溶液。所述氢氧化钠溶液中氢氧化钠的质量分数为3％～30％。优选地，所述氢氧化钠溶液中氢氧化钠的质量分数为5％～20％。得到的亚硫酸氢钠溶液的pH值为3.5～7.0。优选地，亚硫酸氢钠溶液的pH值为4.0～6.5。
步骤(2)中，所述含羟乙基磺酸钠的溶液的pH值为10.0以上。优选的，所述含羟乙基磺酸钠的溶液的pH值为11.0以上。所述含羟乙基磺酸钠的溶液中羟乙基磺酸钠的质量分数为10％～20％。
本申请采用双极膜酸化生产牛磺酸的方法，取代了传统的硫酸或盐酸酸化工艺，节省了酸的投入，避免了副产物硫酸钠或氯化钠的生成，同时产生的氢氧化钠可循环使用，大幅降低了原料成本和废固处理成本。由于没有无机盐产生，使分离提纯工艺更加简单，降低了设备投入和生产成本。整个过程实现了闭路循环，无三废排放，可产业化。
本申请所述牛磺酸中间体及牛磺酸的制备方法具有以下优点：
通过将氨解反应中反应物氨源的比例提高，从而使得氨解反应充分进行，大大提高了反应的收率。
进一步的，可通过氨分离装置将未参与反应的氨分离出来而得到含氨气态物，并通过所述压缩装置将含氨气态物进行压缩得到超临界态流体，再将所述超临界态流体循环至氨解反应器，在此过程中，以较小的能耗实现氨的全循环套用。也就是说，将未反应的氨回收再次参与所述氨解反应，最终提高了氨解反应中氨的浓度，大大降低了成本。另外，在将含氨气态物转换成所述超临界态流体后，所述超临界态流体具有较高的温度及压强，当所述超临界态流体循环至氨解反应器时，可直接将能量耦合至氨解反应器中，从而形成氨解反应过程中所需的高温高压条件，节约了能源。
以下将通过实施例对本发明所述牛磺酸中间体及牛磺酸的制备方法作进一步的说明。
实施例1
将170Kg浓度约15％的羟乙基磺酸钠水溶液中加入200Kg浓度约25％的氨水、101Kg液氨(氨与羟乙基磺酸钠的摩尔比为25:1)，形成混合物，混合物加压至18MPa后，先经预热器预热至280℃，再进入氨解反应器中反应，氨解反应温度280℃，压力18MPa，反应30min后，氨解液除氨后得到牛磺酸钠溶液320Kg，牛磺酸钠含量7.2％，计算得到牛磺酸钠的收率为90.9％。
实施例2
将170Kg浓度约15％的羟乙基磺酸钠水溶液中加入200Kg浓度约25％的氨水、372Kg液氨(氨与羟乙基磺酸钠的摩尔比为70:1)，形成混合物，混合物加压至18MPa后，先经预热器预热至280℃，再进入氨解反应器中反应，氨解反应温度280℃，压力17.8MPa，反应30min后，氨解液除氨后得到牛磺酸钠溶液319Kg，牛磺酸钠含量7.82％，计算得到牛磺酸钠的收率为98.4％。
实施例3
采用如图1所示的氨解反应系统。
将氨水与液氨的混合物与羟乙基磺酸钠水溶液混合后经高压泵加压，预热后通过氨解反应器反应，经蒸发器处理，得到的气相加压循环与羟乙基磺酸钠水溶液直接混合，通入氨解反应器中反应。补加的氨气与蒸发液换热后由蒸发器进入。稳定后，控制系统中，氨与羟乙基磺酸钠的摩尔比为30:1。
具体的工艺条件为：272Kg/h质量分数为15％的羟乙基磺酸钠水溶液经高压泵加压至18MPa，与经过加压后的循环氨直接混合后，温度升高至280℃。混合物通入氨解反应器中，在18MPa、280℃下反应，停留时间30min，得到氨解反应液。将氨解反应液送至蒸发器，蒸发器操作压力为0.1MPa，操作温度为88.9℃。蒸发器出来的第一含氨气态物经压缩机压缩至300℃，18.2MPa，循环至氨解反应器。蒸发器出来的第一液态物与补加的氨换热后得到牛磺酸钠溶液279Kg/h。补加的氨量为8.0Kg/h。
检测牛磺酸钠溶液中各组分的含量，牛磺酸钠含量13.7％，二牛磺酸钠含量1.1％，三牛磺酸钠含量0.09％，计算得到牛磺酸钠的收率为94.3％。
生产单耗为每吨牛磺酸钠消耗2.16吨标煤。
实施例4
采用如图2所示的氨解反应系统。
将氨水与液氨的混合物与羟乙基磺酸钠水溶液混合后经高压泵加压，预热后通过氨解反应器反应，分别经一级闪蒸罐、二级蒸发器逐级处理。二级蒸发得到的气相加压循环至一级闪蒸罐。一级闪蒸得到的气相加压循环而与通入的羟乙基磺酸钠水溶液直接混合加热，通入氨解反应器中反应。补加的氨气由二级蒸发器进入。稳定后，控制系统中，氨与羟乙基磺酸钠的摩尔比为30:1。
具体的工艺条件为：272Kg/h质量分数为15％的羟乙基磺酸钠水溶液经高压泵加压至18MPa，与经过加压后的循环氨直接混合后，温度升高至280℃。混合物通入氨解反应器中，在18MPa、280℃下反应，停留时间30min，得到氨解反应液。将氨解反应液送至一级闪蒸罐闪蒸，一级闪蒸操作压力为8MPa，操作温度为220℃。一级闪蒸罐出来的第一含氨气态物经压缩机压缩至300℃，18.2MPa，循环至氨解反应器。一级闪蒸罐出来的第一液态物进入二级蒸发器。二级蒸发器的操作压力为0.1MPa，操作温度为87.8℃。二级蒸发器得到的第二含氨气态物经压缩至210℃，8.2MPa，循环至一级闪蒸罐中闪蒸。二级蒸发器得到的第二液态物进入与补加的氨换热后得到牛磺酸钠溶液279Kg/h。补加的氨量为8Kg/h。
检测牛磺酸钠溶液中各组分的含量，牛磺酸钠含量13.8％，二牛磺酸钠含量1.09％，三牛磺酸钠含量0.08％，计算得到牛磺酸钠的收率为94.9％。
生产单耗为每吨牛磺酸钠消耗1.07吨标煤。
实施例5
采用如图3所示的氨解反应系统。
将氨水与液氨的混合物与羟乙基磺酸钠水溶液混合后经高压泵加压，预热后通过氨解反应器反应，分别经一级闪蒸罐、二级闪蒸罐、三级蒸发器逐级处理。三级蒸发得到的气相加压循环至二级闪蒸罐。二级闪蒸得到的气相加压循环至一级闪蒸罐。一级闪蒸得到的气相加压循环而与通入的羟乙基磺酸钠水溶液直接混合加热，通入氨解反应器中反应。补加的氨气由三级蒸发器进入。稳定后，控制系统中，氨与羟乙基磺酸钠的摩尔比为30:1。
具体的工艺条件为：272Kg/h质量分数为15％的羟乙基磺酸钠水溶液经高压泵加压至18MPa，与 经过加压后的循环氨直接混合后，温度升高至280℃。混合物通入氨解反应器中，在18MPa、280℃下反应，停留时间30min，得到氨解反应液。将氨解反应液送至一级闪蒸罐闪蒸，一级闪蒸操作压力为8MPa，操作温度为245.5℃。一级闪蒸罐出来的第一含氨气态物经压缩机压缩至300℃，18.2MPa，循环至氨解反应器。一级闪蒸罐出来的第一液态物进入二级闪蒸罐闪蒸。二级闪蒸罐的操作压力为3MPa，操作温度为203.4℃。二级闪蒸罐得到的第二含氨气态物经压缩至290℃，8.2MPa，循环至一级闪蒸罐中闪蒸。二级闪蒸罐得到的第二液态物进入三级蒸发器中，蒸发器的操作压力为0.1MPa，操作温度为97℃。三级蒸发器得到的第三含氨气态物经压缩至210℃，3.1MPa后循环至二级闪蒸罐中闪蒸。三级蒸发器得到的第三液态物与补加的氨换热后得到牛磺酸钠溶液279Kg/h。补加的氨量为8Kg/h。
检测牛磺酸钠溶液中各组分的含量，牛磺酸钠含量13.9％，二牛磺酸钠含量1％，三牛磺酸钠含量0.08％，计算得到牛磺酸钠的收率为95.68％。
生产单耗为每吨牛磺酸钠消耗0.65吨标煤。
实施例6
采用与实施例5相同的操作方法，控制系统中氨与羟乙基磺酸钠的摩尔比为40:1。
具体的工艺条件为：272Kg/h质量分数为15％的羟乙基磺酸钠水溶液经高压泵加压至18MPa，与经过加压后的循环氨直接混合后，温度升高至280℃。混合物通入氨解反应器中，在18MPa、280℃下反应，停留时间30min，得到氨解反应液。将氨解反应液送至一级闪蒸罐闪蒸，一级闪蒸操作压力为8MPa，操作温度为243℃。一级闪蒸罐出来的第一含氨气态物经压缩机压缩至300℃，18.2MPa，循环至氨解反应器。一级闪蒸罐出来的第一液态物进入二级闪蒸罐闪蒸。二级闪蒸罐的操作压力为3MPa，操作温度为203℃。二级闪蒸罐得到的第二含氨气态物经压缩至290℃，8.2MPa，循环至一级闪蒸罐中闪蒸。二级闪蒸罐得到的第二液态物进入三级蒸发器中，蒸发器操作压力为0.1MPa，操作温度为97℃。三级蒸发器得到的第三含氨气态物经压缩至210℃，3.1MPa后循环至二级闪蒸罐中闪蒸。三级蒸发器得到的第三液态物与补加的氨换热后得到牛磺酸钠溶液279Kg/h。补加的氨量为8Kg/h。
检测牛磺酸钠溶液中各组分的含量，牛磺酸钠含量14.1％，二牛磺酸钠含量0.71％，三牛磺酸钠含量0.05％，计算得到牛磺酸钠的收率为97.1％。
生产单耗为每吨牛磺酸钠消耗0.74吨标煤。
实施例7
采用与实施例5相同的操作方法，控制系统中氨与羟乙基磺酸钠的摩尔比为50:1。
具体的工艺条件为：272Kg/h质量分数为15％的羟乙基磺酸钠水溶液经高压泵加压至18MPa，与经过加压后的循环氨直接混合后，温度升高至280℃。混合物通入氨解反应器中，在18MPa、280℃下反应，停留时间30min，得到氨解反应液。将氨解反应液送至一级闪蒸罐闪蒸，一级闪蒸操作压力为8MPa，操作温度为241℃。一级闪蒸罐出来的第一含氨气态物经压缩机压缩至300℃，18.2MPa，循环至氨解反应器。一级闪蒸罐出来的第一液态物进入二级闪蒸罐闪蒸。二级闪蒸罐的操作压力为3MPa，操作温度为203℃。二级闪蒸罐得到的第二含氨气态物经压缩至290℃，8.2MPa，循环至一级闪蒸罐中闪蒸。二级闪蒸罐得到的第二液态物进入三级蒸发器中，蒸发器的操作压力为0.1MPa，操作温度为97℃。三级蒸发器得到的第三含氨气态物经压缩至210℃，3.1MPa后循环至二级闪蒸罐中闪蒸。三级蒸发器得到的第三液态物与补加的氨换热后得到牛磺酸钠溶液279Kg/h。补加的氨量为8Kg/h。
检测牛磺酸钠溶液中各组分的含量，牛磺酸钠含量14.3％，二牛磺酸钠含量0.65％，三牛磺酸钠含量0.03％，计算得到牛磺酸钠的收率为98.43％。
生产单耗为每吨牛磺酸钠消耗0.80吨标煤。
实施例8
将二氧化硫通入到73.0Kg质量浓度为18％的氢氧化钠水溶液，当pH达到4.5时停止通二氧化硫。将13.5Kg环氧乙烷通入到反应液中，反应温度控制在30-40℃，当pH＝11.0时，反应结束，得到含量羟乙基磺酸钠反应液。将羟乙基磺酸钠反应液、用碱室液调pH至11.0的结晶母液在储罐中混合均匀，增压后与循环氨混合，通入氨解反应系统中反应，控制氨与羟乙基磺酸钠的摩尔比为25:1，得到牛磺酸钠溶液，过滤，稀释至10％浓度，进入双极膜电渗析系统进行酸化。碱室得到6％的碱液，物料室获得牛磺酸溶液，牛磺酸溶液进一步浓缩至45％浓度，结晶，得到牛磺酸产品，含量为99.4％，总收率为94％(含母液循环 的收率)。
实施例9
将二氧化硫通入到73.0Kg质量浓度为18％的氢氧化钠水溶液，当pH达到4.5时停止通二氧化硫。将13.5Kg环氧乙烷通入到反应液中，反应温度控制在30-40℃，当pH＝11.0时，反应结束，得到含量羟乙基磺酸钠反应液。将羟乙基磺酸钠反应液、用碱室液调pH至11.0的结晶母液在储罐中混合均匀，增压后与循环氨混合，通入氨解反应系统中反应，控制氨与羟乙基磺酸钠的摩尔比为30:1，得到牛磺酸钠溶液，过滤，稀释至10％浓度，进入双极膜电渗析系统进行酸化。碱室得到6％的碱液，物料室获得牛磺酸溶液，牛磺酸溶液进一步浓缩至45％浓度，结晶，得到牛磺酸产品，含量为99.6％，总收率为94.5％(含母液循环的收率)。
实施例10
将二氧化硫通入到73.0Kg质量浓度为18％的氢氧化钠水溶液，当pH达到4.5时停止通二氧化硫。将13.5Kg环氧乙烷通入到反应液中，反应温度控制在30-40℃，当pH＝11.0时，反应结束，得到含量羟乙基磺酸钠反应液。将羟乙基磺酸钠反应液、用碱室液调pH至11.0的结晶母液在储罐中混合均匀，增压后与循环氨混合，通入氨解反应系统中反应，控制氨与羟乙基磺酸钠的摩尔比为35:1，得到牛磺酸钠溶液，过滤，稀释至10％浓度，进入双极膜电渗析系统进行酸化。碱室得到6％的碱液，物料室获得牛磺酸溶液，牛磺酸溶液进一步浓缩至45％浓度，结晶，得到牛磺酸产品，含量为99.5％，总收率为95％(含母液循环的收率)。
实施例11
将二氧化硫通入到73.0Kg质量浓度为18％的氢氧化钠水溶液，当pH达到4.5时停止通二氧化硫。将13.5Kg环氧乙烷通入到反应液中，反应温度控制在30-40℃，当pH＝11.0时，反应结束，得到含量羟乙基磺酸钠反应液。将羟乙基磺酸钠反应液、用碱室液调pH至11.0的结晶母液在储罐中混合均匀，增压后与循环氨混合，通入氨解反应系统中反应，控制氨与羟乙基磺酸钠的摩尔比为40:1，得到牛磺酸钠溶液，过滤，稀释至10％浓度，进入双极膜电渗析系统进行酸化。碱室得到6％的碱液，物料室获得牛磺酸溶液，牛磺酸溶液进一步浓缩至45％浓度，结晶，得到牛磺酸产品，含量为99.6％，总收率为96.2％(含母液循环的收率)。
对比例1
采用与实施例1相同的操作方法，控制系统中氨与羟乙基磺酸钠的摩尔比为8:1。
具体的工艺条件为：
将170Kg浓度约25％的羟乙基磺酸钠水溶液中加入49.0Kg浓度约25％的氨水、36Kg液氨(氨与羟乙基磺酸钠的摩尔比为8:1)，形成混合物，混合物加压至18MPa后，先经预热器预热至280℃，再进入氨解反应器中反应，氨解反应温度280℃，压力18MPa，反应30min后，氨解液除氨后得到牛磺酸钠溶液205.5Kg，牛磺酸钠含量8％，计算得到牛磺酸钠的收率为64.85％。
对比例2
采用与实施例1相同的操作方法，控制系统中氨与羟乙基磺酸钠的摩尔比为120:1。
具体的工艺条件为：
将170Kg浓度约25％的羟乙基磺酸钠水溶液中加入400.0Kg浓度约25％的氨水、623.7Kg液氨(氨与羟乙基磺酸钠的摩尔比为120:1)，形成混合物，混合物加压至18MPa后，先经预热器预热至280℃，再进入氨解反应器中反应，氨解反应温度280℃，压力18MPa，反应30min后，氨解液除氨后得到牛磺酸钠溶液465Kg，牛磺酸钠含量5.38％，计算得到牛磺酸钠的收率为98.7％。
对比例3
采用现有技术的方法进行羟乙基磺酸钠与氨的氨解反应及氨后处理。
272Kg/h质量分数为15％的羟乙基磺酸钠水溶液与液氨及氨水的混合物流混合，控制氨与羟乙基磺酸钠的摩尔比为30:1，混合物流经高压泵加压至18MPa，预热至280℃，通入氨解反应器中，在18MPa、280℃下反应，停留时间30min，得到氨解反应液。将氨解反应液送至蒸发器处理，得到含量为13.8％的牛磺酸钠溶液278Kg/h，蒸发器的操作压力为0.1MPa，操作温度为97℃。蒸发得到的含氨物质分别经过冷凝器冷凝后，进入蒸氨塔回收处理，回收的氨添加新鲜的氨后循环至氨解反应器重新参与反应。
检测牛磺酸钠溶液中各组分的含量，牛磺酸钠含量13.7％，二牛磺酸钠含量1.1％，三牛磺酸钠含量0.1％，计算得到牛磺酸钠的收率为93.9％。
生产单耗为每吨牛磺酸钠消耗3.72吨标煤。
对比例4
采用现有技术的方法进行羟乙基磺酸钠与氨的氨解反应及氨后处理。
272Kg/h质量分数为15％的羟乙基磺酸钠水溶液与液氨及氨水的混合物流混合，控制氨与羟乙基磺酸钠的摩尔比为30:1，混合物流经高压泵加压至18MPa，预热至280℃，通入氨解反应器中，在18MPa、280℃下反应，停留时间30min，得到氨解反应液。将氨解反应液分别送至两级闪蒸罐，三级蒸发器处理，得到牛磺酸钠溶液278Kg/h，一级闪蒸操作压力为8MPa，操作温度为241℃，二级闪蒸罐的操作压力为3MPa，操作温度为203℃，三级蒸发器的操作压力为0.1MPa，操作温度为97℃。各级闪蒸及蒸发得到的含氨物质分别经过冷凝器冷凝后，进入蒸氨塔回收处理，塔顶回收的氨添加新鲜的氨后循环至氨解反应器重新参与反应。
检测牛磺酸钠溶液中各组分的含量，牛磺酸钠含量13.8％，二牛磺酸钠含量1.2％，三牛磺酸钠含量0.1％，计算得到牛磺酸钠的收率为94.59％。
生产单耗为每吨牛磺酸钠消耗1.32吨标煤。
对比例1采用低氨比进行氨解反应，收率仅有64.85％，且氨解产物中二牛磺酸钠和三牛磺酸钠的含量较高；对比例2采用超过100的氨比，其收率上高达98.7％，但后续氨回收需要较高的成本，且收率的增幅对比实施例8的收率不大。对比例3、4没有采用本发明的压缩装置，也没有将回收的氨变成超临界态流体而进行循环操作，而是直接将氨分离器得到的含氨气相通过蒸氨塔处理，回收的高含量氨再循环至氨解步骤。与实施例3、实施例5对比，发现在同样的氨分离器处理下，采用对比例3、4的氨回收方法需要消耗的能量比采用本发明氨解反应系统消耗的能量多出很多，这导致了成本增加。
以上所述实施例的各技术特征可以进行任意的组合，为使描述简洁，未对上述实施例中的各个技术特征所有可能的组合都进行描述，然而，只要这些技术特征的组合不存在矛盾，都应当认为是本说明书记载的范围。
以上所述实施例仅表达了本发明的几种实施方式，其描述较为具体和详细，但并不能因此而理解为对发明专利范围的限制。应当指出的是，对于本领域的普通技术人员来说，在不脱离本发明构思的前提下，还可以做出若干变形和改进，这些都属于本发明的保护范围。因此，本发明专利的保护范围应以所附权利要求为准。

Claims (18)

  1. 一种牛磺酸中间体牛磺酸钠的制备方法，其特征在于，所述制备方法包括以下步骤：
    提供羟乙基磺酸钠和氨源；
    将所述羟乙基磺酸钠和所述氨源置于氨解反应器中进行氨解反应，得到含牛磺酸中间体牛磺酸钠的混合物，其中，所述氨源中的氨与所述羟乙基磺酸钠的摩尔比为25:1以上。
  2. 如权利要求1所述的牛磺酸中间体牛磺酸钠的制备方法，其特征在于，所述氨源中的氨与所述羟乙基磺酸钠的摩尔比为25:1-100:1。
  3. 如权利要求1所述的牛磺酸中间体牛磺酸钠的制备方法，其特征在于，所述氨源中的氨与所述羟乙基磺酸钠的摩尔比为30:1-50:1。
  4. 如权利要求1所述的牛磺酸中间体牛磺酸钠的制备方法，其特征在于，在所述氨解反应的步骤之后还包括以下步骤：
    通过氨分离装置将所述混合物中未参与反应的氨分离出来而分别得到含氨气态物以及牛磺酸中间体，其中所述氨分离装置与所述氨解反应器相连接；
    通过压缩装置对所述含氨气态物进行压缩，得到含氨的超临界态流体，并将所述超临界态流体循环至所述氨解反应器，其中所述压缩装置分别与所述氨分离装置、所述氨解反应器相连接。
  5. 如权利要求4所述的牛磺酸中间体牛磺酸钠的制备方法，其特征在于，所述氨分离装置包括一个氨分离器。
  6. 如权利要求4所述的牛磺酸中间体牛磺酸钠的制备方法，其特征在于，所述氨分离装置包括两个氨分离器，分别为第一级氨分离器以及第二级氨分离器，
    所述第一级氨分离器与所述氨解反应器相连接，所述第一级氨分离器用于将所述氨解反应后的混合物中未参与反应的氨分离出来，得到第一含氨气态物以及第一剩余混合物；
    所述第二级氨分离器与所述第一级氨分离器相连接，所述第二级氨分离器用于对所述第一剩余混合物再进行氨气分离，得到第二含氨气态物以及第二剩余物，并将所述第二含氨气态物循环至所述第一级氨分离器中。
  7. 如权利要求6所述的牛磺酸中间体牛磺酸钠的制备方法，其特征在于，所述压缩装置包括第一压缩装置和第二压缩装置，
    所述第一压缩装置分别与所述第一级氨分离器、所述氨解反应器相连接，用于将所述第一级氨分离器中的所述含氨气态物进行压缩得到所述超临界态流体，并将所述超临界态流体循环至所述氨解反应器；
    所述第二压缩装置分别与所述第一级氨分离器和所述第二级氨分离器连接，用于将所述第二含氨气态物循环至所述第一级氨分离器中。
  8. 如权利要求4所述的牛磺酸中间体牛磺酸钠的制备方法，其特征在于，所述氨分离装置包括n个依次排列的氨分离器，n为大于2，小于20的整数，
    所述依次排列的n个氨分离器中的第一级氨分离器与所述氨解反应器相连接，所述第一级氨分离器用于将所述氨解反应后所得到的混合物中未参与反应的氨分离出来，得到第一含氨气态物以及第一剩余混合物；
    所述依次排列的n个氨分离器中的第二级氨分离器与所述第一级氨分离器相连接，所述第二级氨分离器用于对所述第一剩余混合物再进行氨气分离，得到第二含氨气态物以及第二剩余物，并将所述第二含氨气态物循环至所述第一级氨分离器中；
    所述依次排列的n个氨分离器中的第i级氨分离器与第i-1级氨分离器相连接，i为整数且2<i≤n，所述第i级氨分离器用于对通过所述第i-1级氨分离器得到的第i-1剩余混合物再进行氨气分离，得到第i含氨气态物以及第i剩余物，并将所述第i含氨气态物循环至第i-1级氨分离器。
  9. 如权利要求8所述的牛磺酸中间体牛磺酸钠的制备方法，其特征在于，所述氨解反应系统还包括n个压缩装置，
    所述n个压缩装置中的第一压缩装置分别与所述第一级氨分离器、所述氨解反应器相连接，用于将所述第一级氨分离器中的所述含氨气态物进行压缩得到所述超临界态流体，并将所述超临界态流体循环至所述氨解反应器；
    所述n个压缩装置中的第二压缩装置分别与所述第一级氨分离器和所述第二级氨分离器连接，用于将所述第二含氨气态物循环至所述第一级氨分离器中；
    所述n个压缩装置中的第i压缩装置分别与所述第i-1级氨分离器和所述第i级氨分离器连接，用于将所述第i含氨气态物循环至第i-1级氨分离器。
  10. 如权利要求8所述的牛磺酸中间体牛磺酸钠的制备方法，其特征在于，所述n为3或4。
  11. 如权利要求1所述的牛磺酸中间体牛磺酸钠的制备方法，其特征在于，所述氨源为氨水混合物、液氨中的至少一种。
  12. 如权利要求4所述的牛磺酸中间体牛磺酸钠的制备方法，其特征在于，还包括向所述氨分离装置补加氨源的步骤。
  13. 一种牛磺酸的制备方法，其特征在于，所述制备方法包括以下步骤：
    提供羟乙基磺酸钠和氨源；
    将所述羟乙基磺酸钠和所述氨源置于氨解反应器中进行氨解反应，得到含牛磺酸中间体牛磺酸钠的混合物，其中所述氨源中的氨与所述羟乙基磺酸钠的摩尔比为25:1以上；
    将所述牛磺酸中间体进行酸化处理，得到牛磺酸。
  14. 如权利要求13所述的牛磺酸的制备方法，其特征在于，所述氨源中的氨与所述羟乙基磺酸钠的摩尔比为25:1-100:1。
  15. 如权利要求13所述的牛磺酸的制备方法，其特征在于，所述氨源中的氨与所述羟乙基磺酸钠的摩尔比为30:1-50:1。
  16. 如权利要求14所述的牛磺酸的制备方法，其特征在于，在所述氨解反应的步骤之后，对所述牛磺酸中间体进行酸化处理的步骤之前还包括以下步骤：
    通过氨分离装置将所述混合物中未参与反应的氨分离出来而分别得到含氨气态物以及牛磺酸中间体牛磺酸钠；
    通过压缩装置对所述含氨气态物进行压缩，得到含氨的超临界态流体，并将所述超临界态流体循环至所述氨解反应器。
  17. 如权利要求13所述的牛磺酸的制备方法，其特征在于，通过双极膜对所述牛磺酸中间体进行所述酸化处理，得到所述牛磺酸和氢氧化钠。
  18. 如权利要求13所述的牛磺酸的制备方法，其特征在于，所述羟乙基磺酸钠通过环氧乙烷与亚硫酸氢钠反应得到，所述亚硫酸氢钠通过二氧化硫与至少部分来自所述通过双极膜对所述牛磺酸中间体进行的所述酸化处理得到的所述氢氧化钠进行反应得到。
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