WO2024093816A1 - 制备己二胺的方法 - Google Patents

制备己二胺的方法 Download PDF

Info

Publication number
WO2024093816A1
WO2024093816A1 PCT/CN2023/127038 CN2023127038W WO2024093816A1 WO 2024093816 A1 WO2024093816 A1 WO 2024093816A1 CN 2023127038 W CN2023127038 W CN 2023127038W WO 2024093816 A1 WO2024093816 A1 WO 2024093816A1
Authority
WO
WIPO (PCT)
Prior art keywords
amination reaction
solvent
stream containing
ammonia
cycloheximide
Prior art date
Application number
PCT/CN2023/127038
Other languages
English (en)
French (fr)
Inventor
舒展
过良
刘智信
罗淑娟
李东风
王国清
樊小哲
田峻
史倩
李�一
李瑜龙
赵梦
邵华伟
孙汝柳
张敬升
曲玉萍
王超
李琰
Original Assignee
中国石油化工股份有限公司
中石化(北京)化工研究院有限公司
Priority date (The priority date is an assumption and is not a legal conclusion. Google has not performed a legal analysis and makes no representation as to the accuracy of the date listed.)
Filing date
Publication date
Priority claimed from CN202211350196.8A external-priority patent/CN117986127A/zh
Priority claimed from CN202211351228.6A external-priority patent/CN117986136A/zh
Application filed by 中国石油化工股份有限公司, 中石化(北京)化工研究院有限公司 filed Critical 中国石油化工股份有限公司
Publication of WO2024093816A1 publication Critical patent/WO2024093816A1/zh

Links

Classifications

    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C209/00—Preparation of compounds containing amino groups bound to a carbon skeleton
    • C07C209/04—Preparation of compounds containing amino groups bound to a carbon skeleton by substitution of functional groups by amino groups
    • C07C209/14—Preparation of compounds containing amino groups bound to a carbon skeleton by substitution of functional groups by amino groups by substitution of hydroxy groups or of etherified or esterified hydroxy groups
    • C07C209/16—Preparation of compounds containing amino groups bound to a carbon skeleton by substitution of functional groups by amino groups by substitution of hydroxy groups or of etherified or esterified hydroxy groups with formation of amino groups bound to acyclic carbon atoms or to carbon atoms of rings other than six-membered aromatic rings
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C211/00—Compounds containing amino groups bound to a carbon skeleton
    • C07C211/01—Compounds containing amino groups bound to a carbon skeleton having amino groups bound to acyclic carbon atoms
    • C07C211/02—Compounds containing amino groups bound to a carbon skeleton having amino groups bound to acyclic carbon atoms of an acyclic saturated carbon skeleton
    • C07C211/09—Diamines
    • C07C211/12—1,6-Diaminohexanes
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C213/00—Preparation of compounds containing amino and hydroxy, amino and etherified hydroxy or amino and esterified hydroxy groups bound to the same carbon skeleton
    • C07C213/02—Preparation of compounds containing amino and hydroxy, amino and etherified hydroxy or amino and esterified hydroxy groups bound to the same carbon skeleton by reactions involving the formation of amino groups from compounds containing hydroxy groups or etherified or esterified hydroxy groups
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07D—HETEROCYCLIC COMPOUNDS
    • C07D295/00—Heterocyclic compounds containing polymethylene-imine rings with at least five ring members, 3-azabicyclo [3.2.2] nonane, piperazine, morpholine or thiomorpholine rings, having only hydrogen atoms directly attached to the ring carbon atoms
    • C07D295/02—Heterocyclic compounds containing polymethylene-imine rings with at least five ring members, 3-azabicyclo [3.2.2] nonane, piperazine, morpholine or thiomorpholine rings, having only hydrogen atoms directly attached to the ring carbon atoms containing only hydrogen and carbon atoms in addition to the ring hetero elements
    • C07D295/023—Preparation; Separation; Stabilisation; Use of additives

Definitions

  • the invention relates to the technical field of hexamethylenediamine production, and in particular to a method for preparing hexamethylenediamine by taking hexamethylenediol as a raw material.
  • the main methods for industrial production of hexamethylenediamine include the adiponitrile method, the caprolactam method, the adipic acid method, the hexamethylenediol amination method, and the like.
  • the adiponitrile method is the most commonly used production method in industrial production.
  • the method is divided into two approaches: high-pressure method and low-pressure method.
  • the high-pressure method uses iron and cobalt-copper as the main catalysts, and the conversion rate can reach more than 95%, but the pressure is high and the equipment requirements are extremely high;
  • the low-pressure method uses Raney nickel as the main catalyst, but there are problems with the safety and stability of the Raney nickel catalyst and the supply of adiponitrile raw materials.
  • caprolactam is aminated with ammonia under phosphate catalysis to produce 6-aminocapronitrile, and 6-aminocapronitrile is hydrogenated to produce hexamethylenediamine.
  • the key step of this process is the ammoniation of caprolactam. Since this process is only suitable for small-scale production and has high production costs, it has been gradually eliminated.
  • the adipic acid method is to react adipic acid with ammonia through ammoniation and dehydration to generate adiponitrile, and then hydrogenate adiponitrile to obtain hexamethylenediamine. This method has high production costs, long processes, and limited process technology development.
  • the hexanediol method is to prepare hexamethylenediamine by amination reaction with hexamethylenediol and ammonia as raw materials.
  • Hexamethylenediol is a colorless and transparent liquid with low toxicity and little harm to the environment and human body. It is a green and environmentally friendly way to prepare hexamethylenediamine with hexamethylenediol as raw material.
  • the common problems in the process of amination of hexamethylenediol to produce hexamethylenediamine are low hexamethylenediol conversion rate, hexamethylenediamine selectivity and hexamethylenediamine yield. According to public data, the hexamethylenediamine yield is usually less than 50%, and the economic efficiency needs to be improved.
  • CN114433122A discloses a catalyst for preparing 1,6-hexanediamine by hydrogenation of 1,6-hexanediol.
  • the molar ratio of hydrogen: ammonia: 1,6-hexanediol is 3:8:1
  • the volume space velocity of 1,6-hexanediol is 0.6h -1
  • the reaction temperature is 205°C
  • the reaction pressure is 13MPa
  • the conversion rate of hexanediol is about 80% to 92%
  • the selectivity of hexanediamine is about 42% to 52%. According to the public data, it can be calculated that the yield of hexanediamine is about 34% to 48%.
  • JP2022112396A discloses a method for producing hexamethylenediamine from 1,6-hexamethylenediol, comprising producing hexamethylenediamine by amination reaction of 1,6-hexamethylenediol under mild conditions in the presence of a solid catalyst, wherein the catalyst comprises an active component containing a metal element selected from Group 8, Group 9, Group 10 and Group 11 elements in the periodic table and a carrier containing one or more metal elements selected from rare earth elements, Group 4 and Group 5 elements.
  • the yield of hexamethylenediamine is less than 40%.
  • the prior art process of preparing hexamethylenediamine by ammoniation of hexamethylenediamine is easy to polymerize to generate high-carbon organic amines, mainly dihexamethylenetriamine and other organic amines with carbon atoms greater than 12, which are collectively referred to as carbon 12 and above amines in the present invention. If these products are not recycled, they can only be discharged as waste liquid, resulting in a low overall yield of the target product hexamethylenediamine. Therefore, a method for preparing hexamethylenediamine is still needed, in which hexamethylenediamine can be produced from hexamethylenediamine in a high yield, thereby improving the process economy.
  • aminolysis described in the present invention refers to the cracking of high-carbon amines into low-carbon amines, specifically refers to the chemical reaction of heavy components, that is, carbon 12 and above amines defined above in the present invention, under the action of a catalyst to generate low-carbon amines such as hexamethylenediamine.
  • the present invention not only improves the yield of hexamethylenediamine by ammonolysis of the heavy components in the amination reaction, but also reduces the discharge of waste liquid from the amination reaction, thereby greatly reducing the production cost of hexamethylenediamine.
  • the process of the present invention is inherently safe, does not involve highly toxic nitrile chemicals, and has a safe and environmentally friendly technical route.
  • the hexamethylenediamine product obtained by the present invention has a stable output, high purity, high yield, and few impurities.
  • the present invention provides a method for preparing hexamethylenediamine by alcohol amination, which has a simple process flow, saves equipment investment, and has high economic added value of the product.
  • any values of the ranges disclosed in this article are not limited to the precise ranges or values, and these ranges or values should be understood to include values close to these ranges or values.
  • the endpoint values of each range, the endpoint values of each range and the individual point values, and the individual point values can be combined with each other to obtain one or more new numerical ranges, which should be regarded as specifically disclosed in this article.
  • the present invention provides a method for preparing hexamethylenediamine, which comprises the following steps:
  • solvent i.e. the first solvent
  • the aminating reaction is not indispensable, but is optionally used, so the aminating reaction can be carried out in a solvent or in the absence of a solvent.
  • the aminating reaction is usually carried out at normal pressure or low pressure, and ammonia is in a gas phase state, so it is necessary to add a solvent to dissolve the reactant stream ammonia to effectively carry out the aminating reaction.
  • the pressure of the aminating reaction is higher, and ammonia is in a liquid phase state through liquefaction, so hexylene glycol and liquefied ammonia can carry out the aminating reaction in the absence of a solvent.
  • the solvent added in the present invention is to dissolve the aminating reaction raw materials and reaction products such as hexylene glycol, hexamethylenediamine, aminohexanol, cyclohexylimine, N-(6-aminohexyl) cyclohexylimine.
  • the molar ratio of ammonia to hexylene glycol is 10-90:1, preferably 10-80:1, such as 15-70:1, 15-60:1, 20-45:1.
  • the amination reaction is carried out in the presence of hydrogen, and the molar ratio of hydrogen to hexylene glycol is 0.05-25:1, preferably 0.05-20:1, such as 0.1-10:1, 0.2-5:1, 0.3-3:1.
  • the molar ratio of ammonia, hydrogen and hexylene glycol refers to the molar ratio in the mixture at the inlet of the amination reactor.
  • the temperature of the amination reaction is 100-300° C., preferably 100-250° C., such as 120-250° C., 130-210° C., 140-200° C.
  • the pressure of the amination reaction is 4-25 MPaG, preferably 6-20 MPaG, such as 6-18 MPaG, 8-16 MPaG, 10-14 MPaG.
  • the liquid volume space velocity of fresh hexanediol is 0.01-15 h -1 , preferably 0.01-10 h -1 , such as 0.03-8 h -1 , 0.05-4 h -1 , 0.1-2 h -1 .
  • the type of the amination reaction catalyst used in the amination reaction process of the present invention can be any catalyst in the art that can amination hexanediol to produce hexamethylenediamine, for example, the amination reaction catalysts in CN114433086A, CN114433087A, CN114433113A, etc.
  • the catalyst comprises a carrier and an active component and an optional auxiliary agent supported on the carrier, wherein the carrier comprises a doping element, alumina and other optional carriers,
  • the other carriers are selected from at least one of silicon oxide, molecular sieve and diatomaceous earth;
  • the pore volume of the carrier with a pore size less than 7.5nm accounts for less than 20% of the pore volume of the carrier, the pore volume of the carrier with a pore size less than 9nm accounts for less than 40% of the pore volume of the carrier, and the pore volume of the carrier with a pore size greater than 27nm accounts for less than 5% of the pore volume of the carrier;
  • the ammonia adsorption capacity of the carrier is 0.3-0.6mmol/g;
  • the L acid of the carrier accounts for more than 90% of the sum of the L acid and the B acid;
  • the active component is cobalt and/or nickel.
  • the carrier is selected from alumina doped with at least one of silicon oxide, molecular sieve and diatomaceous earth and alumina without addition.
  • the content of the alumina carrier in the carrier accounts for more than 65% by weight of the total amount of the alumina carrier and other carriers, preferably more than 75% by weight.
  • the carrier may further include a doping element, and the content of the doping element accounts for 0.05-3% by weight, more preferably 0.08-2% by weight, and further preferably 0.1-1.5% by weight of the total weight of the components other than the doping element in the carrier.
  • the components other than the doping element mainly refer to the alumina in the carrier and other optional carriers.
  • the hetero elements added to the carrier are from acid radical ions excluding chloride ions. Since the added hetero elements are introduced during the preparation process of the carrier, the added hetero elements are mainly present in the bulk phase of the carrier.
  • the acid ion can be selected from at least one of non-metallic acid ions, more preferably at least one of borate ions, fluoride ions, phosphate ions, sulfate ions and selenate ions.
  • the doping element is preferably selected from at least one of boron, fluorine, phosphorus, sulfur and selenium.
  • the pore volume of the carrier with a pore size less than 7.5 nm accounts for 5-17% of the pore volume of the carrier, more preferably 5-10%
  • the pore volume with a pore size greater than or equal to 7.5 nm and less than 9 nm accounts for 5-17% of the pore volume of the carrier
  • the pore volume with a pore size greater than or equal to 9 nm and less than or equal to 27 nm accounts for 61-89.5% of the pore volume of the carrier
  • the pore volume with a pore size greater than 27 nm accounts for 0.5-5%, more preferably 0.5-3%.
  • the inventors of the present invention have found that the catalyst having a pore structure that meets this preferred embodiment has a more excellent catalytic performance.
  • the ammonia adsorption capacity of the carrier is preferably 0.3-0.5 mmol/g.
  • the L acid of the carrier accounts for 92-100% of the sum of the L acid and the B acid, preferably 96-100%.
  • the L acid proportion is measured by pyridine probe adsorption spectroscopy.
  • the specific surface area of the carrier is 105-220 m 2 /g, and the pore volume of the carrier is 0.4-1.1 ml/g.
  • the content of the active component can be 5-42g, preferably 10-35g, relative to every 100g of the carrier calculated as the components other than the doping element.
  • the catalyst may also contain an auxiliary agent.
  • the auxiliary agent may be selected from at least one of the VIB group, the VIIB group, the IB group, the IIB group and the lanthanide elements, preferably at least one of Cr, Mo, W, Mn, Re, Cu, Ag, Au, Zn, La and Ce.
  • the content of the auxiliary agent may be 0-10 g, preferably 0.5-6 g, per 100 g of the carrier calculated as components other than the doping element.
  • the present invention does not limit the reactor used in the amination reaction, and can be any reactor capable of performing gas-liquid-solid reaction. Considering the full reaction of the reaction raw materials, improving the reaction efficiency and enhancing the reaction effect, preferably, the reactor used in the amination reaction is one of a fixed bed, an autoclave, a trickle bed, and a fluidized bed, or other reactors that can ensure stable operation of the reaction.
  • step (2) further comprises: separating a logistics containing cycloheximide and water from the amination reaction product.
  • separating the logistics containing cycloheximide and water to obtain a logistics containing cycloheximide, and returning at least part of the logistics containing cycloheximide to step (1).
  • the logistics containing cycloheximine and water can be separated to obtain the logistics containing cycloheximine, and at least part of the logistics containing cycloheximine is returned to step (1).
  • the weight ratio of the logistics containing cycloheximine returned to step (1) to the fresh hexamethylenediol is 0.1-26:1, more preferably 0.3-24:1, for example 0.5-20:1, 1-15:1, 2-10:1, 3-8:1.
  • step (2) further comprises: separating a stream containing aminohexanol from the amination reaction product.
  • step (1) at least part of the stream containing aminohexanol is returned to step (1).
  • the aminohexanol-containing stream can be returned to step (1).
  • the weight ratio of the aminohexanol-containing stream returned to step (1) to the fresh hexamethylenediol is 0.1-26:1, more preferably 0.3-24:1, such as 0.5-20:1, 1-15:1, 2-10:1, 3-8:1.
  • the aminohexanol-containing stream may also contain 0-30% by weight of hexamethylenediamine, 0-2% by weight of hexamethylenediamine, and a small amount of other components (such as heavy components).
  • the amount of cyclohexyl imine generated by the amination reaction is small and cannot meet the weight ratio of the logistics containing cyclohexyl imine to fresh hexylene glycol specified above. Therefore, in the initial stage of the reaction, the logistics containing cyclohexyl imine is all returned to the amination reaction. The amount of imine is sufficient, and at this moment, it is only necessary to ensure that the logistics containing cyclohexylimide returned satisfies the weight ratio defined above. Similarly, the aminohexanol generated by amination reaction is also true.
  • the weight ratio of the logistics returned to step (1) and fresh hexylene glycol given in the embodiment of the application is the mass ratio in the stable operation stage of the reaction.
  • step (2) further comprises: separating a stream containing cycloheximide and water from the amination reaction product, separating the stream containing cycloheximide and water to obtain a stream containing cycloheximide, and returning at least part of the stream containing cycloheximide to step (1); and separating a stream containing aminohexanol from the amination reaction product, and returning at least part of the stream containing aminohexanol to step (1).
  • the yield of hexamethylenediamine can be further improved.
  • the weight ratio of the logistics containing cycloheximide and the logistics containing aminohexanol in the return step (1) is 0.1-5: 1, such as 0.2-4.5: 1, 0.3-4: 1, 0.4-3.5: 1, 0.5-3: 1, 0.6-2.5: 1, 0.7-2: 1, 0.8-1.5: 1.
  • the weight ratio of the total weight of the logistics containing cycloheximide and the logistics containing aminohexanol in the return step (1) to the fresh hexamethylenediol is 0.1-30: 1, preferably 0.3-25: 1, such as 0.4-20: 1, 0.5-15: 1, 0.8-10: 1, 1-8: 1, 1.5-6: 1, 2-5: 1.
  • the amount of cyclohexyl imine and aminohexanol generated by aminating reaction is less, can not meet the gross weight of the logistics containing cyclohexyl imine and the logistics containing aminohexanol and the weight ratio of fresh hexylene glycol, therefore in the initial stage of reaction, the logistics containing cyclohexyl imine and the logistics containing aminohexanol all return to the aminating reaction.And in the reaction stable operation stage, the amount of cyclohexyl imine and aminohexanol generated by aminating reaction is more, as long as the logistics containing cyclohexyl imine and the logistics containing aminohexanol that guarantee to return meet the weight ratio of the above-mentioned limitation.
  • the weight ratio of the logistics and fresh hexylene glycol returned to step (1) given in the embodiment of the application is the mass ratio in the reaction stable operation stage.
  • the conditions for separation of the amination reaction product are such that the content of cyclohexylamine in the logistics containing cyclohexylamine and water is 35-95% by weight, more preferably 40-85% by weight, most preferably 50-80% by weight; the water content is 5-65% by weight, more preferably 20-50% by weight, most preferably 30-40% by weight.
  • the method for separating the amination reaction product is: subjecting the amination reaction product to a first separation to obtain a logistics containing cycloheximide and water, and then subjecting the product to a second separation to obtain a hexamethylenediamine product and a logistics containing aminohexanol; more preferably, the first separation method and the second separation method each independently include at least one of distillation, membrane separation and pressure swing adsorption.
  • the amination reaction product is separated in the following manner: the amination reaction product is subjected to a first separation to obtain a stream containing cycloheximide and water, followed by a second separation to obtain a hexamethylenediamine product and a stream containing aminohexanol and heavy components, and then a third separation to obtain a stream containing aminohexanol and heavy components; more preferably, the first separation, the second separation and the third separation method each independently includes at least one of distillation, membrane separation and pressure swing adsorption.
  • the conditions for the first separation may include: a theoretical plate number of 10-80 (e.g., 10, 20, 30, 40, 50, 60, 70, 80, and a range consisting of any of the above values), an operating pressure of 1-600 KPa (absolute pressure, for example, 1 KPa, 10 KPa, 20 KPa, 30 KPa, 40 KPa, 50 KPa, 100 KPa, 200 KPa, 300 KPa, 400 KPa, 50 0KPa, 600KPa, and the range consisting of the above arbitrary point values), tower bottom temperature 90-300°C, reflux ratio 0.1-15;
  • the second separation conditions may include: theoretical plate number 10-80 (for example, 10, 20, 30, 40, 50, 60, 70, 80, and the range consisting of the above arbitrary point values), operating pressure 1-600KPa (absolute pressure, for example, 1KPa, 10KPa, 20KPa, 30KPa, 40KPa, 50
  • the separation method of the amination reaction product is as follows: the amination reaction product enters a first distillation tower for separation, and a logistics containing cycloheximide and water is obtained at the top of the first distillation tower; the bottom logistics of the first distillation tower is sent to a second distillation tower for separation, and a hexamethylenediamine product is obtained at the top of the tower, and a logistics containing hexamethylenediol, aminohexanol and heavy components is obtained at the bottom of the tower; the bottom logistics of the second distillation tower is sent to a third distillation tower for separation, and a logistics containing hexamethylenediol and aminohexanol is obtained at the top of the third distillation tower, and a heavy component is obtained at the bottom of the tower.
  • the operating conditions of the first distillation tower include: bottom temperature of 90-300°C, top temperature of 30-65°C, reflux ratio of 0.1-10, top operating pressure of -0.1 MPaG to 0.5 MPaG, and theoretical plate number of 10-30.
  • the operating conditions of the second distillation tower include: bottom temperature of 120-320°C, top temperature of 40-85°C, reflux ratio of 0.5-30, top operating pressure of -0.1 MPaG to 0.5 MPaG, and theoretical plate number of 10-45.
  • the operating conditions of the third distillation tower include: bottom temperature of 150-360°C, top temperature of 40-85°C, reflux ratio of 1-65, top operating pressure of -0.1MPaG to 0MPaG, and theoretical plate number of 15-85.
  • the logistics containing cycloheximide and water can be separated by a dehydration method commonly used in the art.
  • the dehydration treatment method includes at least one of atmospheric distillation, vacuum distillation, pressure swing distillation, azeotropic distillation, membrane separation, extractive distillation, adsorption dehydration, biological treatment dehydration and centrifugal separation.
  • the entrainer used in the azeotropic distillation can be at least one of cyclohexane, n-hexane, trimethylpentane, p-methyl isopropyl benzene, 1,4-dioxane, phenol, cresol, butyl ether, amyl ether and isoamyl ether.
  • the operating conditions of the azeotropic distillation are specifically selected according to the different azeotropic points formed by the specific entrainer, water and cycloheximide.
  • the water content in the cycloheximide-containing stream returned to step (1) is less than 3% by weight, preferably less than 1.5% by weight, more preferably less than 1% by weight.
  • the dehydration treatment of the logistics containing cycloheximide and water can be carried out by membrane separation dehydration, and the operating conditions of the membrane separation dehydration include: the membrane inlet pressure is 0-5MPaG, such as 1-4MPaG, and the temperature is 10-350°C, such as 10-200°C, such as 15-190°C.
  • the dehydration treatment of the logistics containing cycloheximide and water can be carried out by azeotropic distillation dehydration, and the azeotropic distillation dehydration is carried out in a distillation tower.
  • the operating conditions of the distillation tower include: the weight ratio of the entrainer to the logistics containing cycloheximide is 10-100:1, the bottom temperature is 120-200°C, the top temperature is 80-150°C, the reflux ratio is 0.1-20, the top operating pressure is 0-3MPaG, and the number of plates is 10-80.
  • the dehydration treatment of the logistics containing cycloheximide and water can be carried out by distillation and membrane separation or pressure swing distillation at the same time.
  • the logistics containing cycloheximide and water is distilled in a distillation tower, the tower bottom obtains the logistics containing cycloheximide, and the overhead logistics is subjected to membrane separation or pressure swing distillation to obtain waste water.
  • the conditions for distilling the logistics containing cycloheximide and water include: the tower bottom temperature is 60-250°C, and the tower top operating pressure is -0.09MPaG to 1MPaG.
  • the conditions for membrane separation include a membrane inlet temperature of 80-350°C and a membrane inlet pressure of 0-5MPaG.
  • the pressure swing operation range of the pressure swing distillation can be 0-8MPaG, preferably 0-5MPaG; more preferably, the pressure swing distillation is carried out in a high-pressure tower and a low-pressure tower, the top pressure of the high-pressure tower is 1-8MPaG, and the bottom temperature is 150-240°C; the top pressure of the low-pressure tower is less than 1MPaG, and the bottom temperature is 50-130°C.
  • the amination reaction product is first pre-separated to recover the hydrogen and ammonia in the amination reaction product; wherein the recovered hydrogen and ammonia are first condensed to obtain circulating ammonia (liquid phase), and then compressed to obtain circulating hydrogen (gas phase), and the circulating hydrogen and circulating ammonia are returned to the amination reaction.
  • the pre-separation method includes at least one of flash evaporation, distillation and stripping.
  • the pre-separation method may also adopt other methods capable of recovering circulating hydrogen and ammonia from the amination reaction product. More preferably, the content of hydrogen in the circulating hydrogen is 5-40% by weight, and the content of ammonia is 60-95% by weight; the content of ammonia in the circulating ammonia is 40-100% by weight, and the content of water is 10-20%.
  • the content of hexamethylenediamine is 0-1% by weight, the content of cycloheximide is 0-1% by weight.
  • the pre-separation conditions make the recovery rate of hydrogen greater than 99% and the recovery rate of ammonia greater than 98%.
  • the pre-separation method can adopt multi-stage reduced pressure flash evaporation, and the multi-stage flash evaporation method can include at least 2 stages of flash evaporation, and the pressure of the flash evaporation decreases with a gradient of 1-8MPa, and the pressure of the last stage of flash evaporation is 0.5-1MPaG.
  • the flash evaporation pressures of the first to fourth stages are 7-15MPaG, 4-7MPaG, 1-4MPaG, and 0.1-1MPaG, respectively.
  • the flash evaporation pressures of the first to third stages are 7-15MPaG, 1-7MPaG, and 0.5-3MPaG, respectively.
  • step (4) further comprises: before the aminolysis reaction, dissolving the heavy component in a second solvent, and then performing the aminolysis.
  • the second solvent is a solvent in which the solubility of the amination reaction raw materials and reaction products such as hexamethylenediol, hexamethylenediamine, aminohexanol, cyclohexylimine, and dihexamethylenetriamine is greater than 0.05 g/g, such as at least one of tetrahydrofuran, 1,4-dioxane, n-hexane, cyclohexane, and tert-butyl alcohol.
  • the weight ratio of the second solvent to the heavy component is 1-20:1, preferably 2-18:1, and more preferably 3-15:1.
  • the second solvent for dissolving the heavy components can be the same type as the first solvent used in the amination reaction (if a solvent is present).
  • the solvent for dissolving the heavy components can be a freshly introduced solvent or a solvent stream separated from the amination reaction product (if a solvent is present in the amination reaction).
  • the amount of solvent in the system is small, and a fresh solvent can be used to dissolve the heavy components.
  • the amount of solvent in the system is large, and the solvent stream separated from the amination reaction product can meet the amount of solvent required for the amination reaction and the aminolysis reaction.
  • step (3) the aminolysis is carried out in the presence of hydrogen and ammonia, and the molar ratio of hydrogen:ammonia:heavy component is 0.1-20:10-150:1, preferably 0.2-15:15-120:1, more preferably 0.2-10:20-100:1, and most preferably 0.3-8:30-90:1, wherein the heavy component is calculated as dihexamethylenetriamine.
  • the conditions for the ammonolysis reaction include: reaction temperature of 120-300°C, preferably 150-270°C; reaction pressure of 9-25 MPaG, preferably 10-22 MPaG; liquid volume space velocity of the heavy component of 0.01-8h -1 , preferably 0.05-5h -1 .
  • ammonolysis catalyst used in the ammonolysis reaction can be prepared according to the following method:
  • a carrier contacting a mixture of pseudo-boehmite, silica sol and calcium nitrate with an aqueous solution containing nitric acid and phosphoric acid, and then kneading, drying and calcining in sequence;
  • the support is impregnated in an aqueous solution containing nickel sulfate, lanthanum acetate and indium nitrate, dried at 100-140° C. for 2-6 hours, and then calcined at 350-450° C. for 2-6 hours.
  • the amount of silica sol is 0.6-0.8 g
  • the amount of calcium nitrate is 0.1-0.4 g
  • the amount of aqueous solution containing nitric acid and phosphoric acid is 0.2-0.5 g per gram of pseudo-boehmite.
  • the content of nitric acid in the aqueous solution containing nitric acid and phosphoric acid is 10-25% by weight, and the content of phosphoric acid is 5-15% by weight.
  • the drying temperature is 100-140° C.
  • the drying time is 1-6 h.
  • the amount of nickel sulfate used is 0.6-0.85 g
  • the amount of lanthanum acetate used is 0.06-0.08 g
  • the amount of indium nitrate used is 0.055-0.065 g per gram of the carrier.
  • the concentration of nickel sulfate in the aqueous solution containing nickel sulfate, lanthanum acetate and indium nitrate is 20-25% by weight, the concentration of lanthanum acetate is 1.5-2.5% by weight, and the concentration of indium nitrate is 1.5-2.5% by weight.
  • the impregnation method is preferably an equal volume impregnation method, and the impregnation can be performed in multiple times.
  • the present invention also finds that by returning the logistics containing aminohexanol and the logistics containing cycloheximide to the amination reaction in a certain proportion and subjecting the heavy components in the amination reaction to aminolysis, not only the yield of hexamethylenediamine is greatly improved, but also the discharge amount of the amination reaction waste liquid is reduced, thereby greatly reducing the production cost of hexamethylenediamine.
  • the first solvent is present in the amination reaction
  • Step (1) comprises: subjecting hexanediol and ammonia to an amination reaction in a first solvent under amination reaction conditions to obtain an amination reaction product;
  • Step (2) comprises separating a stream containing cyclohexylimide, a first solvent and water, a hexamethylenediamine product, a stream containing aminohexanol and a heavy component from the amination reaction product of step (1).
  • the method for preparing hexamethylenediamine comprises the following steps:
  • step (2) separating a stream containing cycloheximide, the first solvent and water from the amination reaction product of step (1); Hexamethylenediamine products, aminohexanol-containing streams and heavy components;
  • the method for preparing hexamethylenediamine comprises the following steps:
  • step (2) (2) separating a stream containing cycloheximide, a first solvent and water, a hexamethylenediamine product, a stream containing aminohexanol and a heavy component from the amination reaction product of step (1), then separating the stream containing cycloheximide, the first solvent and water to obtain a stream containing cycloheximide and a stream containing the first solvent, and returning at least a portion of the stream containing cycloheximide and at least a portion of the stream containing the first solvent to step (1);
  • the method for preparing hexamethylenediamine comprises the following steps:
  • step (2) (2) separating a stream containing cycloheximide, a first solvent and water, a hexamethylenediamine product, a stream containing aminohexanol and a heavy component from the amination reaction product of step (1), returning at least a portion of the stream containing aminohexanol to step (1), and then separating the stream containing cycloheximide, a first solvent and water to obtain a stream containing cycloheximide and a first solvent stream, and returning at least a portion of the stream containing cycloheximide and at least a portion of the first solvent stream to step (1);
  • step (3) Under the conditions of the ammonolysis reaction, the heavy component is ammonolyzed to obtain an ammonolysis reaction product, and the ammonolysis reaction product is returned to step (2) with or without separation.
  • Introducing a solvent into the amination reaction can maintain uniform temperature distribution in the system, which is beneficial to reducing reaction hot spots, improving reaction stability, and ensuring that solid phase crystallization does not occur in any part of the system, thereby reducing the operating temperature of each device in the device and avoiding the huge investment caused by device insulation.
  • the introduction of a solvent can effectively dissolve by-products such as hexanediol, cyclohexylimide, and aminohexanol produced by the amination reaction, which helps solve the problem of pipeline blockage during the entire operation cycle, especially in the separation process.
  • the first solvent is introduced once (introduced with hexylene glycol in the initial stage of the reaction).
  • the molar ratio of the first solvent to hexylene glycol is 0.1-10:1, preferably 0.15-10:1, and more preferably 0.25-9:1.
  • the first solvent is used in the circulation and purification process. There will be loss.
  • a small amount of fresh solvent can be added during the first solvent circulation process to maintain a constant amount of solvent in the system.
  • the first solvent is a solvent in which the solubility of amination reaction raw materials and reaction products such as hexanediol, hexamethylenediamine, aminohexanol, cycloheximide is greater than 0.1g/g, such as tetrahydrofuran, 1,4-dioxane, n-hexane, cyclohexane, tert-butyl alcohol, cyclopentane sulfone and glycerol.
  • amination reaction raw materials and reaction products such as hexanediol, hexamethylenediamine, aminohexanol, cycloheximide is greater than 0.1g/g, such as tetrahydrofuran, 1,4-dioxane, n-hexane, cyclohexane, tert-butyl alcohol, cyclopentane sulfone and glycerol.
  • the molar ratio of ammonia to hexylene glycol is 10-80:1, preferably 20-45:1; the amination reaction is carried out in the presence of hydrogen, and the molar ratio of hydrogen to hexylene glycol is 0.05-20:1, preferably 0.2-5:1.
  • the molar ratio of ammonia, hydrogen and hexylene glycol refers to the molar ratio in the mixture at the inlet of the amination reactor.
  • the temperature of the amination reaction is 100-250°C, preferably 130-210°C; the pressure of the amination reaction is 6-18 MPaG, preferably 8-16 MPaG, and the liquid volume space velocity of fresh hexanediol is 0.01-10 h -1 , preferably 0.05-4 h -1 .
  • the weight ratio of the logistics containing cyclohexyl imine and the logistics containing aminohexanol in the return step (1) is 0.1-5:1, for example 0.2-4.5:1, 0.3-4:1, 0.4-3.5:1, 0.5-3:1, 0.6-2.5:1, 0.7-2:1, 0.8-1.5:1, thereby further improving the yield of hexamethylenediamine.
  • the solvent since the solvent only plays the role of dissolving reaction raw materials and reaction product, there is no restriction on the amount of the returned solvent, and it can be all returned or partially returned.
  • the weight ratio of the logistics containing cyclohexyl imine, the first solvent logistics and the logistics containing aminohexanol can be 0.1-5:0.3-40:1, preferably 0.1-2.5:0.3-10:1, more preferably 0.5-2:1-10:1.
  • the weight ratio of the total weight of the logistics containing cyclohexylimide and the logistics containing aminohexanol returned to step (1) to the fresh hexylene glycol is 0.1-30:1, preferably 0.3-25:1, for example 0.4-20:1, 0.5-15:1, 0.8-10:1, 1-8:1, 1.5-6:1, 2-5:1, so as to further improve the yield of hexylenediamine.
  • step (2) further comprises: before separating the logistics containing cycloheximide, the first solvent and water, the hexamethylenediamine product, the logistics containing aminohexanol and the heavy component in the amination reaction product, recovering hydrogen and ammonia in the amination reaction product; wherein the recovered hydrogen and ammonia are first condensed to obtain circulating ammonia (liquid phase), and then compressed to obtain circulating hydrogen (gas phase), and the circulating hydrogen and circulating ammonia are returned to the amination reaction.
  • the content of hydrogen in the circulating hydrogen is 5-40% by weight, and the content of ammonia is 60-95% by weight; the content of ammonia in the circulating ammonia is 40-100% by weight, the content of solvent is 0-60% by weight, the content of water is 0-1% by weight, the content of hexamethylenediamine is 0-1% by weight, and the content of cycloheximide is 0-1% by weight.
  • step (2) comprises: the amination reaction product from which ammonia and hydrogen are optionally removed is fed into a first distillation tower for separation, and a logistics containing cycloheximide, water and the first solvent is obtained at the top of the first distillation tower; the bottom logistics of the first distillation tower is fed into a second distillation tower for separation, and a hexamethylenediamine product is obtained at the top of the tower, and a logistics containing hexamethylenediol, aminohexanol and heavy components is obtained at the bottom of the tower; the bottom logistics of the second distillation tower is fed into a third distillation tower for separation, and a logistics containing hexamethylenediol and aminohexanol is obtained at the top of the third distillation tower, and a heavy component is obtained at the bottom of the tower.
  • the operating conditions of the first distillation tower include: bottom temperature of 90-300°C, top temperature of 30-65°C, reflux ratio of 0.1-10, top operating pressure of -0.1 MPaG to 0.5 MPaG, and number of plates of 10-30.
  • the operating conditions of the second distillation tower include: bottom temperature of 120-320°C, top temperature of 40-85°C, reflux ratio of 0.5-30, top operating pressure of -0.1 MPaG to 0.5 MPaG, and number of plates of 10-45.
  • the operating conditions of the third distillation tower include: bottom temperature of 150-360°C, top temperature of 40-85°C, reflux ratio of 1-65, top operating pressure of -0.1 MPaG to 0 MPaG, and number of plates of 15-85.
  • step (2) comprises: distilling the logistics containing cycloheximide, the first solvent and water in a distillation tower, obtaining the logistics containing cycloheximide in the tower bottom, and performing membrane separation or pressure-swing distillation on the top logistics to obtain the solvent logistics and wastewater.
  • the conditions for distilling the logistics containing cycloheximide, the first solvent and water include: the tower bottom temperature is 60-250°C, and the tower top operating pressure is -0.09MPaG to 1MPaG.
  • the conditions for membrane separation include the membrane inlet temperature of 80-350°C and the membrane inlet pressure of 0-5MPaG.
  • the pressure swing operation range of the pressure swing distillation can be 0-8MPaG, preferably 0-5MPaG; more preferably, the pressure swing distillation is carried out in a high-pressure tower and a low-pressure tower, the tower top pressure of the high-pressure tower is 1-8MPaG, and the tower bottom temperature is 150-240°C; the tower top pressure of the low-pressure tower is less than 1MPaG, and the tower bottom temperature is 50-130°C.
  • the ammonolysis reaction product is roughly separated by distillation, and the conditions of the rough separation tower for the rough separation include: the bottom temperature of the tower is 380-520°C, the top operating pressure of the tower is -0.05MPag to 1MPag, and the reflux ratio is 2-50.
  • Step (1) comprises: under an amination reaction condition and in the absence of a first solvent, allowing hexanediol and ammonia to undergo an amination reaction to obtain an amination reaction product.
  • the method for preparing hexamethylenediamine comprises the following steps:
  • step (2) (2) separating a stream containing cyclohexylimide and water, a hexamethylenediamine product, a stream containing aminohexanol, and a heavy component from the amination reaction product of step (1);
  • the method for preparing hexamethylenediamine comprises the following steps:
  • step (2) (2) separating a stream containing cycloheximide and water, a hexamethylenediamine product, a stream containing aminohexanol, and a heavy component from the amination reaction product of step (1), then separating the stream containing cycloheximide and water to obtain a stream containing cycloheximide, and returning at least a portion of the stream containing cycloheximide to step (1);
  • the method for preparing hexamethylenediamine comprises the following steps:
  • step (2) (2) separating a stream containing cycloheximide and water, a hexamethylenediamine product, a stream containing aminohexanol, and a heavy component from the amination reaction product of step (1), returning at least a portion of the stream containing aminohexanol to step (1), and then separating the stream containing cycloheximide and water to obtain a stream containing cycloheximide, and returning at least a portion of the stream containing cycloheximide to step (1);
  • step (3) Under the conditions of the ammonolysis reaction, the heavy component is ammonolyzed to obtain an ammonolysis reaction product, and the ammonolysis reaction product is returned to step (2) with or without separation.
  • the molar ratio of ammonia to hexylene glycol is 18-50:1, preferably 22-38:1, and the molar ratio of hydrogen to hexylene glycol is 0.08-6:1, preferably 0.4-3:1.
  • the molar ratio of ammonia, hydrogen and hexylene glycol refers to the molar ratio in the mixture at the inlet of the amination reactor.
  • the temperature of the amination reaction is 120-210° C., preferably 130-200° C.
  • the pressure of the amination reaction is 7-17 MPaG, preferably 9-16 MPaG.
  • the liquid volume space velocity of fresh hexanediol is 0.05-7 h ⁇ 1 , preferably 0.09-3.9 h ⁇ 1 .
  • the weight ratio of the cycloheximide-containing logistics returned to step (1) to fresh hexylene glycol is 0.1-13:1, preferably 0.3-10:1.
  • step (1) 100.77 g of nickel sulfate hexahydrate (industrial grade, purity 98%), 5.69 g of lanthanum acetate monohydrate and 5.96 g of indium nitrate pentahydrate were added to 134.78 mL of water to prepare an aqueous solution, and the solution was loaded on 73.25 g of the carrier obtained in step (1) by an equal volume impregnation method twice. After each impregnation, it was dried at 120° C. for 4 hours. After the two impregnations, it was calcined at 390° C. for 4 hours.
  • the solvent (cyclohexane), hexanediol, ammonia and hydrogen raw materials are fed into a fixed-bed ammoniation reactor filled with a catalyst.
  • the catalyst is the catalyst of Example 14 in CN114433086A.
  • the molar ratio of ammonia to hexanediol is 35:1
  • the molar ratio of hydrogen to hexanediol is 0.5:1
  • the molar ratio of solvent to hexanediol is 3.8:1
  • the amination reaction temperature is 166°C
  • the amination reaction pressure is 12 MPaG
  • the liquid volume space velocity of fresh hexanediol is 0.1 h -1 .
  • the product of the initial amination reaction (before the circulating stream returns to the fixed-bed ammoniation reactor) comprises: hexanediamine 5.76wt%, hexanediol 1.13wt%, aminohexanol 4.91wt%, cycloheximide 4.76wt%, ammonia 48.1wt%, cyclohexane 31.59wt%, water 2.56wt%, and the rest are heavy components.
  • the molar yield of hexanediamine in the initial amination reaction is 31.1%.
  • the amination reaction product is subjected to gas-liquid separation to remove hydrogen and ammonia therefrom, and the hydrogen is returned to the fixed bed amination reactor in the form of gas through a compressor, and the liquid ammonia is pumped back to the amination reactor.
  • the gas-liquid separation method comprises: the reaction product is subjected to gas-liquid separation in four-stage flash tanks of 12MPaG, 5MPaG, 2MPaG, and 0.5MPaG, wherein the gas phase at the top of the second and third stage flash tanks is condensed and refluxed at a reflux ratio of 0.1, and the fourth stage is heated before entering the flash tank, and a 1m filler is provided in the tank.
  • the gas phase obtained by flash evaporation is then condensed with condensed water at 30-45°C to obtain circulating ammonia (liquid phase) and gas phase (hydrogen), and the gas phase is compressed to obtain circulating hydrogen, wherein the circulating hydrogen is mainly composed of 9.72wt% hydrogen and the rest is ammonia;
  • the circulating ammonia is mainly composed of 94.96 wt% of ammonia, 4.12 wt% of cyclohexane, 0.73 wt% of water, 0.01 wt% of hexamethylenediamine and 0.15 wt% of cycloheximide.
  • the amination reaction product from which ammonia and hydrogen are removed enters a first distillation tower for vacuum distillation, and a logistics containing cycloheximide, water and solvent is obtained at the top of the first distillation tower (the logistics composition is: 80.11wt% of cyclohexane, 13.01wt% of cycloheximide, and 5.79wt% of water);
  • the operating conditions of the first distillation tower include: a bottom temperature of 179.9°C, a top temperature of 40.5°C, a reflux ratio of 1.5, a top operating pressure of -0.08MPa, and 25 plates.
  • the bottom stream of the first distillation tower is sent to the second distillation tower for separation, and the hexamethylenediamine product (the content of hexamethylenediamine in the product is 99.8 weight %) is obtained at the top of the tower, and the stream containing hexamethylenediol, aminohexanol and heavy components is obtained at the bottom of the tower (the stream composition is 79.73wt% of aminohexanol, 18.35wt% of hexamethylenediol, 0.2wt% of hexamethylenediamine, and the rest is the heavy component carbon dodecylamine);
  • the operating conditions of the second distillation tower include: the bottom temperature is 189.1°C, the top temperature is 83°C, the reflux ratio is 2.7, the top operating pressure is -0.09MPa, and the number of tower plates is 24.
  • the bottom material of the second distillation tower is sent to the third distillation tower for separation, and the top of the third distillation tower obtains the material containing hexanediol and aminohexanol, and the bottom of the tower obtains the heavy components;
  • the operating conditions of the third distillation tower include: the bottom temperature is 267.5°C, the top temperature is 83.7°C, the reflux ratio is 5.3, the top operating pressure is -0.09MPa, and the number of tower plates is 38.
  • step (3) rectifying the logistics containing cyclohexylimide, water and solvent obtained in step (2) to obtain a logistics containing cyclohexylimide (the logistics composition is cyclohexylimide content of 99.91wt% and hexamethylenediamine content of 0.09wt%) and a logistics containing solvent and water, wherein the conditions of rectification include a bottom temperature of 94.3° C. and a tower top operating pressure of ⁇ 0.08MPa.
  • the logistics composition is cyclohexane content of 99.67wt% and cyclohexylimide content of 0.32wt%) and waste water (cyclohexane content of less than 100ppmw), wherein the conditions of membrane separation include a membrane inlet temperature of 100° C. and a membrane inlet pressure of 0.05MPaG.
  • step (2) The heavy components obtained from the bottom of the third distillation tower in step (2) are mixed evenly with the solvent (cyclohexane) and then sent to an ammonolysis reactor filled with a catalyst for ammonolysis. Ammonia and hydrogen are introduced into the ammonolysis reactor at the same time.
  • the molar ratio of hydrogen: ammonia: heavy components calculated as dihexamethylenetriamine is 1:40:1, the weight ratio of solvent to heavy components is 5:1, the ammonolysis reaction temperature is 186°C, the ammonolysis reaction pressure is 16 MPaG, and the liquid volume space velocity of the heavy components is 0.5 h -1 .
  • the obtained ammonolysis reaction product has the following composition: 0.07 wt% hydrogen, 55.66 wt% ammonia, 7.84 wt% hexamethylenediamine, 6.01 wt% cycloheximide, 21.92 wt% cyclohexane, and the rest are heavy components.
  • the ammonolysis reaction product is all sent back to step (2).
  • the weight ratio of the total weight of the logistics containing cyclohexylimide and the logistics containing aminohexanol in the return step (1) to the fresh hexylene glycol is 2.88.
  • the weight ratio of the logistics containing cyclohexylimide and the logistics containing aminohexanol in the return step (1) is 0.72:1.
  • the system runs continuously until it reaches a stable state, and the molar yield of hexamethylenediamine reaches 93.7%. Long enough period of continuous and stable operation.
  • the temperature distribution in the reactor containing the solvent is uniform without hot spots, and the temperature value deviation of multiple temperature measurement points distributed at the same bed height is less than 1.2°C.
  • the solvent (1,4-dioxane), hexanediol, ammonia and hydrogen raw materials are fed into a fixed-bed amination reactor filled with a catalyst.
  • the catalyst is the catalyst of Example 6 in CN114433086A.
  • the molar ratio of ammonia to hexanediol is 25:1, the molar ratio of hydrogen to hexanediol is 1:1, the molar ratio of solvent to hexanediol is 1.2:1,
  • the amination reaction temperature is 190°C, the amination reaction pressure is 10 MPaG, and the liquid volume space velocity of fresh hexanediol is 0.6 h -1 .
  • the product of the initial amination reaction (before the circulating stream returns to the fixed-bed amination reactor) comprises: hexanediamine 8.39wt%, hexanediol 0.93wt%, aminohexanol 5.97wt%, cycloheximide 6.67wt%, ammonia 34.83wt%, 1,4-dioxane 39.02wt%, water 3.12wt%, and the rest are heavy components.
  • the molar yield of hexanediamine in the initial amination reaction is 36.4%.
  • the amination reaction product is subjected to gas-liquid separation to remove hydrogen and ammonia therefrom, and the hydrogen is returned to the fixed bed amination reactor in the form of gas phase through a compressor, and the liquid ammonia is pumped back to the amination reactor.
  • the gas-liquid separation method includes: the reaction product is subjected to three-stage gas-liquid separation at 10MPaG, 2MPaG, and 0.8MPaG, wherein a section of filler is provided on the top of the second-stage flash tank, and the gas phase on the tank top is condensed and refluxed at a reflux ratio of 0.3, and a 1m filler and a tank bottom heater are provided in the third-stage flash tank.
  • the gas phase obtained by flash evaporation is then condensed with condensed water at 30-45°C to obtain circulating ammonia (liquid phase) and gas phase (hydrogen), and the gas phase is compressed to obtain circulating hydrogen, the circulating hydrogen mainly consisting of 10.19wt% hydrogen and the rest ammonia; the circulating ammonia mainly consists of 93wt% ammonia, 5.79wt% 1,4-dioxane, 0.91wt% water, 0.01wt% hexamethylenediamine, and 0.21wt% cycloheximide.
  • the amination reaction product from which ammonia and hydrogen are removed enters a first distillation tower for vacuum distillation, and a logistics containing cycloheximide, water and solvent is obtained at the top of the first distillation tower (the logistics composition is: 79.93wt% of 1,4-dioxane, 14.40wt% of cycloheximide, and 5.86wt% of water);
  • the operating conditions of the first distillation tower include: a bottom temperature of 176.9°C, a top temperature of 44.2°C, a reflux ratio of 2.6, a top operating pressure of -0.08MPa, and 19 plates.
  • the bottom stream of the first distillation tower is sent to the second distillation tower for separation, and the hexamethylenediamine product (the content of hexamethylenediamine in the product is 99.9 weight %) is obtained at the top of the tower, and the bottom of the tower obtains a stream containing hexamethylenediol, aminohexanol and heavy components (the stream composition is: aminohexanol 85.36wt%, hexamethylenediol 13.33wt%, hexamethylenediamine 0.25wt%, and the rest is the heavy component carbon dodecylamine);
  • the operating conditions of the second distillation tower include: the bottom temperature is 188.3°C, the top temperature is 84.3°C, the reflux ratio is 1.9, the top operating pressure is -0.09MPa, and the number of plates is 33.
  • the bottom stream of the second distillation tower is sent to the third distillation tower for separation, and the top of the third distillation tower obtains a stream containing hexamethylenediol and aminohexanol, and the bottom of the tower obtains heavy components;
  • the operating conditions of the three distillation towers include: the bottom temperature is 270.5°C, the top temperature is 83.6°C, the reflux ratio is 5, the top operating pressure is -0.09MPa, and the number of tower plates is 39.
  • step (3) distilling the logistics containing cyclohexylimide, water and solvent obtained in step (2) to obtain a logistics containing cyclohexylimide (the logistics composition is: 99.94wt% of cyclohexylimide and 0.06wt% of hexamethylenediamine) and a logistics containing solvent and water, wherein the distillation conditions include a bottom temperature of 207.3° C.
  • the logistics composition is: 99.64wt% of 1,4-dioxane and 0.36wt% of cyclohexylimide) and waste water (99.95wt% of water and 0.05wt% of 1,4-dioxane), wherein the pressure swing distillation conditions include a high-pressure tower top pressure of 1 MPaG and a bottom temperature of 197.2° C., a low-pressure tower top pressure of 0 MPaG and a bottom temperature of 96.3° C.
  • the heavy component obtained from the bottom of the third distillation tower in step (2) is mixed evenly with the solvent (1,4-dioxane) and then sent to an ammonolysis reactor filled with a catalyst for ammonolysis, and ammonia and hydrogen are introduced into the ammonolysis reactor at the same time.
  • the molar ratio of hydrogen: ammonia: heavy component calculated as dihexamethylenetriamine is 3:50:1, the weight ratio of solvent to heavy component is 11:1, the ammonolysis reaction temperature is 203°C, the ammonolysis reaction pressure is 15 MPaG, and the liquid volume space velocity of the heavy component is 0.8h -1 .
  • the obtained ammonolysis reaction product has the following composition: 4.75wt% hydrogen, 71.74wt% ammonia, 2.60wt% hexamethylenediamine, 3.52wt% cycloheximide, 15.26wt% 1,4-dioxane, and the rest is heavy components.
  • the product of the ammonolysis reaction is depressurized to 10 MPag, and flash distillation is performed at this pressure to separate gas and liquid to obtain a stream containing hydrogen and ammonia and a residual liquid phase, and the residual liquid phase is then distilled in a coarse fractionation tower to obtain a mixed liquid containing hexamethylenediamine and cycloheximide and a heavy component.
  • the top pressure of the coarse fractionation tower is -0.07 MPag, the top temperature is 180°C, the bottom temperature is 480°C, and the reflux ratio is 20.
  • the mixed liquid containing hexamethylenediamine and cycloheximide at the top of the coarse fractionation tower is returned to step (2), and the bottom stream is discharged as waste liquid.
  • the weight ratio of the total weight of the logistics containing cyclohexylimide and the logistics containing aminohexanol in the return step (1) to the fresh hexylene glycol is 2.95.
  • the weight ratio of the logistics containing cyclohexylimide and the logistics containing aminohexanol in the return step (1) is 0.97:1.
  • the system runs continuously to a stable state, and the molar yield of hexamethylenediamine reaches 94.8%.
  • the system has no blockage and can run continuously and stably for a long period of time.
  • the temperature distribution in the reactor containing the solvent is uniform without hot spots, and the temperature value deviation of multiple temperature measurement points distributed at the same bed height is less than 1.5°C.
  • the gas-liquid separation method comprises: the product of the amination reaction is sequentially passed through three flash tanks for three-stage vacuum flash evaporation to recover hydrogen, wherein the pressures of the first flash tank to the third flash tank are set to 7MPa, 5MPa, and 2MPa, respectively, and the gas phases at the top of the three flash tanks are cooled to 45°C and further subjected to gas-liquid separation, and the obtained liquid phase is returned to the previous vacuum flash tank, and the gas phase is pressurized to the amination reaction pressure and then returned to the amination reactor.
  • the liquid phase at the bottom of the third vacuum flash tank enters the deamination distillation tower from the top of the tower, the number of tower plates is 11, and the operating pressure at the top of the tower is 2MPa.
  • no condenser is provided at the top of the deamination distillation tower; the top logistics of the deamination distillation tower are pressurized to the amination reaction pressure and then sent back to the inlet of the amination reactor, and the bottom logistics are subsequently refined.
  • the amination reaction product from which ammonia and hydrogen are removed is cut and separated with hexamethylenediamine as a key component according to different boiling points, and the first separation is carried out in sequence to obtain a cyclohexyl imine-containing logistics and a residual logistics, and then the residual logistics is subjected to a second separation to obtain a hexamethylenediamine product and a logistics containing aminohexanol, C12 amine and a heavy component, and then the logistics containing aminohexanol, C12 amine and a heavy component is subjected to a third separation to obtain a logistics containing aminohexanol, a logistics containing C12 amine and a heavy component.
  • the logistics containing aminohexanol and the logistics containing cyclohexyl imine are mixed in a weight ratio of 1:1.25 and returned to step (1) as a circulating logistics, and the heavy component is extracted.
  • the first separation is carried out in the first distillation tower, and the operating conditions of the first distillation tower include: the bottom temperature is 187°C, the top temperature is 50°C, the reflux ratio is 2.1, the theoretical number of plates is 18, and the top operating pressure is -0.07MPa;
  • the second separation is carried out in the second distillation tower, and the operating conditions of the second distillation tower include: the bottom temperature is 195°C, the top temperature is 131°C, the reflux ratio is 5, the theoretical number of plates is 19, and the top operating pressure is -0.08MPa;
  • the third separation is carried out in the third distillation tower, and the operating conditions of the third distillation tower include: the bottom temperature is 338°C, the top temperature is 87°C, the reflux ratio is 5, the theoretical number of plates is 55, and
  • step (3) The heavy component obtained from the bottom of the third distillation tower in step (2) is sent to an ammonolysis reactor filled with a catalyst for ammonolysis, and ammonia and hydrogen are introduced into the ammonolysis reactor at the same time.
  • the molar ratio of hydrogen: ammonia: heavy component calculated as dihexamethylenetriamine is 5:100:1, the ammonolysis reaction temperature is 170° C., the ammonolysis reaction pressure is 12 MPag, and the liquid space velocity of the C12 amine-containing stream is 0.5 h -1 .
  • the product of the ammonolysis is returned to step (2).
  • Table 1 Note: 11: Logistics composition at the inlet of the amination reactor; 12: Amination reaction products; 13: Circulating hydrogen and ammonia; 14: Circulating logistics; 15: hexamethylenediamine product; 16: heavy component; 17: logistics containing C12 amine; 19: light component; "/" in the table means the content is less than 500ppm, or the content is 0.
  • Example 1 The method of Example 1 was followed, except that the heavy components were not subjected to aminolysis.
  • the reaction was running stably, the molar yield of hexamethylenediamine was 86.5%.
  • Example 2 The method of Example 2 was followed, except that the heavy components were not subjected to aminolysis.
  • the reaction was running stably, the molar yield of hexamethylenediamine was 87.7%.
  • Example 3 The method of Example 3 was followed, except that the heavy components were not subjected to aminolysis.
  • the reaction was running stably, the molar yield of hexamethylenediamine was 83.91%.

Landscapes

  • Chemical & Material Sciences (AREA)
  • Organic Chemistry (AREA)
  • Chemical Kinetics & Catalysis (AREA)
  • Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)

Abstract

本发明公开了一种制备己二胺的方法，该方法包括以下步骤：（1）在氨化反应条件下，任选在第一溶剂存在下，使己二醇与氨发生氨化反应得到氨化反应产物；（2）从步骤（1）的氨化反应产物中分离得到己二胺产品和重组分；（3）在氨解反应条件下，将重组分进行氨解得到氨解反应产物，所述氨解反应产物返回至步骤（2）。本发明不仅可以提高己二胺收率，还可以提高催化剂的使用寿命。

Description

制备己二胺的方法 技术领域
本发明涉及己二胺生产技术领域，具体涉及一种以己二醇为原料制备己二胺的方法。
背景技术
己二胺是一种重要的化工原料，是制造聚酰胺尼龙66必需的一种原料单体，还用于制造其它种类聚酰胺如尼龙610、尼龙612等，是工程塑料行业重要的中间原料。随着尼龙66的市场规模正不断扩大，其原料己二胺的需求量也在不断增长。
工业上生产己二胺的方法主要有己二腈法、己内酰胺法、己二酸法、己二醇氨化法等。
己二腈法是目前工业生产中使用最为普遍的一种生产方法，该方法又分为高压法和低压法两种途径，其中高压法以铁系、钴-铜系为主要催化剂，转化率可达到95％以上，但是压力高，对设备要求极高；低压法以雷尼镍为主要催化剂，但存在雷尼镍催化剂的安全性和稳定性问题、己二腈原料的供应问题等。
在己内酰胺法中，己内酰胺在磷酸盐催化条件下与氨进行氨化生产6-氨基己腈，6-氨基己腈加氢生成己二胺，该方法的关键步骤为己内酰胺的氨化。由于该法只适用于小规模生产，且由于生产成本偏高而逐渐被淘汰。
己二酸法是己二酸与氨经过氨化、脱水生成己二腈，己二腈加氢得到己二胺。此法生产成本高、工序长，工艺技术发展受限。
己二醇法是以己二醇和氨为原料进行氨化反应制备己二胺。己二醇为无色透明液体，毒性小，对环境及人体的危害小，以己二醇为原料制备己二胺是一条绿色环保的路径。然而，在现有技术中，己二醇氨化生成己二胺的过程中普遍存在的问题是己二醇转化率、己二胺选择性及己二胺收率不高。根据公开数据测算，己二胺收率通常低于50％，经济性有待提高。
CN114433122A公开了将1,6-己二醇临氢氨化制成1,6-己二胺的催化剂，在固定床反应器中，氢气:氨:1,6-己二醇的摩尔比为3:8:1,1,6-己二醇的体积空速为0.6h-1，反应温度205℃，反应压力13MPa，己二醇的转化率大约在80％到92％之间，己二胺选择性大约为42％～52％。根据公开的数据，计算可知己二胺收率约为34％-48％。
JP2022112396A公开了一种由1,6-己二醇生产己二胺的方法，包括在固体催化剂存在下，在温和的条件下，通过1,6-己二醇的胺化反应生产己二胺，该催化剂包含含有一种选自元素周期表中第8族、第9族、第10族和第11族元素的金属元素的活性组分和含有一种或多种选自稀土元素、第4族和第5族元素的金属元素的载体。实施例显示，己二胺的收率低于40％。
另外，现有技术的己二醇氨化制己二胺过程易聚合生成高碳原子有机胺，主要为双六甲撑三胺和其它碳原子大于12的有机胺，本发明中统称为碳十二及以上胺类。这部分产物若不加以回收利用，只能作为废液排出，导致目标产品己二胺总体收率低。因此仍需要一种制备己二胺的方法，其中可由己二醇以高的收率生产己二胺，从而提高工艺经济性。
发明内容
本发明的目的是为了克服现有技术中存在的上述问题，提供一种以高收率制备己二胺的方法。
为了实现上述目的，本发明提供了一种制备己二胺的方法，该方法包括以下步骤：
(1)在氨化反应条件下，任选在第一溶剂存在下，使己二醇与氨发生氨化反应得到氨化反应产物；
(2)从步骤(1)的氨化反应产物中分离得到己二胺产品和重组分；
(3)在氨解反应条件下，将重组分进行氨解得到氨解反应产物，所述氨解反应产物返回至步骤(2)。
需要强调的是，本发明所述的氨解指的是将高碳数的胺类裂解为低碳数的胺类，具体地是指将重组分，也就是本发明上文定义的碳十二及以上胺类，在催化剂作用下，发生化学反应，生成己二胺等低碳胺类。
本发明的方法具有以下有益效果：
(1)本发明通过将氨化反应中的重组分进行氨解，不仅提高了己二胺的收率，还减少了氨化反应废液的排放量，极大地降低己二胺生产成本。
(2)本发明工艺流程本质安全性高，不涉及剧毒的腈类化学品，技术路线安全、环保。本发明获得的己二胺产品产量稳定、纯度高、收率高，杂质少。
(3)对于有己二醇或上游原料资源的企业，本发明提供了一种醇氨化制取己二胺的方法，工艺流程简单，设备投资省，产品经济附加值高。
具体实施方式
在本文中所披露的范围的端点和任何值都不限于该精确的范围或值，这些范围或值应当理解为包含接近这些范围或值的值。对于数值范围来说，各个范围的端点值之间、各个范围的端点值和单独的点值之间，以及单独的点值之间可以彼此组合而得到一个或多个新的数值范围，这些数值范围应被视为在本文中具体公开。
本发明提供了一种制备己二胺的方法，该方法包括以下步骤：
(1)在氨化反应条件下，任选在第一溶剂存在下，使己二醇与氨发生氨化反应得到氨化反应产物；
(2)从步骤(1)的氨化反应产物中分离得到己二胺产品和重组分；
(3)在氨解反应条件下，将重组分进行氨解得到氨解反应产物，所述氨解反应产物返回至步骤(2)。
根据本发明，在己二醇和氨的氨化反应中溶剂(即第一溶剂)不是必不可少的，而是任选使用的，因此氨化反应既可以在溶剂中进行，也可以在没有溶剂的情况下进行。在现有技术中，氨化反应通常在常压或低压下进行，氨处于气相状态，因此需要添加溶剂来溶解反应物流氨气以有效进行氨化反应。在本发明中，氨化反应的压力较高，氨经液化而处于液相状态，因此己二醇和液氨可以在不存在溶剂下进行氨化反应。本发明中添加的溶剂是为了溶解己二醇、己二胺、氨基己醇、环己亚胺、N-(6-氨基己基)环己亚胺等氨化反应原料和反应产物。
根据本发明，在己二醇和氨的氨化反应过程中，氨与己二醇的摩尔比为10-90:1，优选为10-80:1，例如15-70:1、15-60:1、20-45:1。氨化反应在氢气存在下进行，氢气与己二醇的摩尔比为0.05-25:1，优选为0.05-20:1，例如0.1-10:1、0.2-5:1、0.3-3:1。氨、氢气与己二醇的摩尔比是指氨化反应器入口的混合物中的摩尔比。
根据本发明，在己二醇和氨的氨化反应中，氨化反应的温度为100-300℃，优选为100-250℃，例如120-250℃、130-210℃、140-200℃。氨化反应的压力为4-25MPaG，优选为6-20MPaG，例如6-18MPaG、8-16MPaG、10-14MPaG。新鲜己二醇的液体体积空速为0.01-15h-1，优选为0.01-10h-1，例如0.03-8h-1、0.05-4h-1、0.1-2h-1。
本发明的氨化反应过程中采用的氨化反应催化剂的种类可以为本领域中能够使己二醇发生氨化生产己二胺的任何催化剂，例如，CN114433086A、CN114433087A、CN114433113A等中的氨化反应催化剂。优选地，所述催化剂包括载体和负载于所述载体上的活性组分和任选的助剂，所述载体包括掺杂元素、氧化铝和任选的其它载体， 其中，所述其它载体选自氧化硅、分子筛和硅藻土中的至少一种；所述载体中孔径小于7.5nm的孔容占所述载体孔容的百分比低于20％，孔径小于9nm的孔容占所述载体孔容的百分比低于40％，孔径大于27nm的孔容占所述载体孔容的百分比低于5％；所述载体的氨吸附量为0.3-0.6mmol/g；所述载体的L酸占L酸与B酸之和的90％以上；所述活性组分为钴和/或镍。
根据本发明所述的氨化反应催化剂，优选地，所述载体选自掺加氧化硅、分子筛和硅藻土中至少一种的氧化铝以及未掺加的氧化铝。所述载体中氧化铝载体的含量占氧化铝载体与其它载体的总量的65重量％以上，优选75重量％以上。
根据本发明所述的氨化反应催化剂，优选地，所述载体还可以包括掺杂元素，掺杂元素的含量占载体中除掺杂元素以外的成分的总重量的0.05-3重量％，更优选为0.08-2重量％，进一步优选为0.1-1.5重量％。除掺杂元素以外的成分主要指载体中的氧化铝和任选的其它载体。
根据本发明所述的氨化反应催化剂，优选地，所述载体中掺加的杂元素来自不包括氯离子的酸根离子。由于掺加的杂元素在载体的制备过程中引入，掺加的杂元素主要存在于载体的体相中。
根据本发明所述的氨化反应催化剂，优选地，酸根离子可以选自非金属酸根离子中的至少一种，进一步优选为硼酸根离子、氟离子、磷酸根离子、硫酸根离子和硒酸根离子中的至少一种。掺杂元素优选选自硼、氟、磷、硫和硒中的至少一种。
根据本发明所述的氨化反应催化剂，优选地，所述载体中孔径小于7.5nm的孔容占所述载体孔容的百分比为5-17％，更优选为5-10％，孔径大于等于7.5nm且小于9nm的孔容占所述载体孔容的百分比为5-17％，孔径大于等于9nm且小于等于27nm的孔容占所述载体孔容的百分比为61-89.5％，孔径大于27nm的孔容占所述载体孔容的百分比为0.5-5％，更优选为0.5-3％。本发明的发明人发现，孔道结构满足该优选实施方式的催化剂具有更优异的催化性能。
根据本发明所述的氨化反应催化剂，优选地，所述载体的氨吸附量优选为0.3-0.5mmol/g。
根据本发明所述的氨化反应催化剂，优选地，所述载体的L酸占L酸与B酸之和的92-100％，优选为96-100％。L酸占比通过吡啶探针吸附光谱法测得。
根据本发明所述的氨化反应催化剂，优选地，所述载体的比表面积为105-220m2/g，所述载体的孔容为0.4-1.1ml/g。
根据本发明所述的氨化反应催化剂，优选地，相对于每100g的以除掺杂元素以外的成分计的载体，活性组分的含量可以为5-42g，优选为10-35g。根据本发明，为了更好发挥出本发明的催化剂的性能、调优反应产物比例、减少不需要的副反应，催化剂还可以含有助剂。助剂可以选自VIB族、VIIB族、IB族、IIB族和镧系元素中的至少一种，优选为Cr、Mo、W、Mn、Re、Cu、Ag、Au、Zn、La和Ce中至少一种。
根据本发明所述的氨化反应催化剂，优选地，相对于每100克的以除掺杂元素以外的成分计的载体，助剂的含量可以为0-10g，优选为0.5-6g。
本发明对氨化反应采用的反应器不做限定，可以为能够进行气液固反应的任何反应器，考虑反应原料充分反应，提高反应效率、增强反应效果，优选地，氨化反应采用的反应器为固定床、高压釜、滴流床、流化床中的一种，也可以是保证反应稳定运行的其它反应器。
根据本发明，在制备己二胺的方法中，步骤(2)还包括：从氨化反应产物中分离得到含环己亚胺和水的物流。任选地，将含环己亚胺和水的物流进行分离得到含环己亚胺的物流，将至少部分的含环己亚胺的物流返回至步骤(1)。
为了进一步提高己二胺的收率，可将含环己亚胺和水的物流进行分离得到含环己亚胺的物流，将至少部分的含环己亚胺的物流返回至步骤(1)。根据本发明，在综合考虑系统能耗、己二胺的收率和催化剂的使用寿命的情况下，优选地，返回步骤(1)中的含环己亚胺的物流与新鲜己二醇的重量比为0.1-26:1，更优选为0.3-24:1，例如0.5-20:1、1-15:1、2-10:1、3-8:1。
根据本发明，在制备己二胺的方法中，步骤(2)还包括：从氨化反应产物中分离得到含氨基己醇的物流。任选地，将至少部分的含氨基己醇的物流返回至步骤(1)。
为了进一步提高己二胺的收率，可将至少部分的含氨基己醇的物流返回至步骤(1)。根据本发明，在综合考虑系统能耗、己二胺的收率和催化剂的使用寿命的情况下，优选地，返回步骤(1)中的含氨基己醇的物流与新鲜己二醇的重量比为0.1-26:1，更优选为0.3-24:1，例如0.5-20:1、1-15:1、2-10:1、3-8:1。含氨基己醇的物流还可能含有0-30重量％的己二醇、0-2重量％的己二胺、以及少量的其它组分(例如重组分)。
需要说明的是，由于反应开始时，氨化反应生成的环己亚胺的量较少，不能满足上述限定的含环己亚胺的物流与新鲜己二醇的重量比，因此在反应初始阶段，含环己亚胺的物流全部返回至氨化反应中。而在反应稳定运行阶段，氨化反应累积生成的环 己亚胺的量足够多，此时只要保证返回的含环己亚胺的物流满足上述限定的重量比即可。同样地，氨化反应生成的氨基己醇也是如此。本申请的实施例中给出的返回至步骤(1)的物流与新鲜己二醇的重量比是反应稳定运行阶段的质量比。
根据本发明，在制备己二胺的方法中，步骤(2)还包括：从氨化反应产物中分离得到含环己亚胺和水的物流，和将含环己亚胺和水的物流进行分离得到含环己亚胺的物流，将至少部分的含环己亚胺的物流返回至步骤(1)；和从氨化反应产物中分离得到含氨基己醇的物流，和将至少部分的含氨基己醇的物流返回至步骤(1)。由此，可以进一步提高己二胺的收率。
根据本发明，在综合考虑系统能耗、己二胺的收率和催化剂的使用寿命的情况下，优选地，返回步骤(1)中的含环己亚胺的物流和含氨基己醇的物流的重量比为0.1-5:1，例如0.2-4.5:1、0.3-4:1、0.4-3.5:1、0.5-3:1、0.6-2.5:1、0.7-2:1、0.8-1.5:1。优选地，返回步骤(1)中的含环己亚胺的物流和含氨基己醇的物流的总重量与新鲜己二醇的重量比为0.1-30:1，优选为0.3-25:1，例如0.4-20:1、0.5-15:1、0.8-10:1、1-8:1、1.5-6:1、2-5:1。
需要说明的是，由于反应开始时，氨化反应生成的环己亚胺和氨基己醇的量较少，不能满足上述限定的含环己亚胺的物流和含氨基己醇的物流的总重量与新鲜己二醇的重量比，因此在反应初始阶段，含环己亚胺的物流和含氨基己醇的物流全部返回至氨化反应中。而在反应稳定运行阶段，氨化反应累积生成的环己亚胺和氨基己醇的量较多多，此时只要保证返回的含环己亚胺的物流和含氨基己醇的物流满足上述限定的重量比即可。本申请的实施例中给出的返回至步骤(1)的物流与新鲜己二醇的重量比是反应稳定运行阶段的质量比。
根据本发明，优选地，氨化反应产物分离的条件使得含环己亚胺和水的物流中环己亚胺的含量为35-95重量％，更优选为40-85重量％，最优选为50-80重量％；水含量为5-65重量％，更优选为20-50重量％，最优选为30-40重量％。
优选地，氨化反应产物分离的方式为：将氨化反应产物进行第一分离得到含环己亚胺和水的物流，然后进行第二分离得到己二胺产品和含氨基己醇的物流；更优选地，第一分离和第二分离的方式各自独立地包括精馏、膜分离和变压吸附中的至少一种。
优选地，氨化反应产物的分离方式为：将氨化反应产物进行第一分离得到含环己亚胺和水的物流，接着进行第二分离得到己二胺产品和含氨基己醇和重组分的物流，然后进行第三分离得到含氨基己醇的物流和重组分；更优选地，第一分离、第二分离 和第三分离的方式各自独立地包括精馏、膜分离和变压吸附中的至少一种。
本发明中，例如，当采用精馏时，第一分离的条件可以包括：理论塔板数10-80(例如，10、20、30、40、50、60、70、80，以及上述任意点值组成的范围)，操作压力1-600KPa(绝压，例如，1KPa、10KPa、20KPa、30KPa、40KPa、50KPa、100KPa、200KPa、300KPa、400KPa、500KPa、600KPa，以及上述任意点值组成的范围)，塔釜温度90-300℃，回流比0.1-15；第二分离的条件可以包括：理论塔板数10-80(例如，10、20、30、40、50、60、70、80，以及上述任意点值组成的范围)，操作压力1-600KPa(绝压，例如，1KPa、10KPa、20KPa、30KPa、40KPa、50KPa、100KPa、200KPa、300KPa、400KPa、500KPa、600KPa，以及上述任意点值组成的范围)，塔釜温度130-380℃，回流比0.1-40；第三分离条件可以包括：理论塔板数10-100(例如，10、20、30、40、50、60、70、80，90，100，以及上述任意点值组成的范围)，操作压力1-1100KPa(绝压，例如，1KPa、10KPa、20KPa、30KPa、40KPa、50KPa、100KPa、200KPa、300KPa、400KPa、500KPa、600KPa、700KPa、800KPa、900KPa、1000KPa、1100KPa，以及上述任意点值组成的范围)，塔釜温度150-400℃，回流比0.1-70。本发明的分离条件可以使己二胺产品纯度在99.7wt％以上。
更优选地，氨化反应产物的分离方式为：氨化反应产物进入第一精馏塔中进行分离，第一精馏塔的塔顶得到含环己亚胺和水的物流；第一精馏塔的塔釜物流送入第二精馏塔进行分离，塔顶得到己二胺产品，塔釜得到含己二醇、氨基己醇和重组分的物流；第二精馏塔的塔釜物流送入第三精馏塔进行分离，第三精馏塔的塔顶得到含己二醇和氨基己醇的物流，塔釜得到重组分。
根据本发明，优选地，第一精馏塔的操作条件包括：塔釜温度为90-300℃，塔顶温度为30-65℃，回流比为0.1-10，塔顶操作压力为-0.1MPaG至0.5MPaG，理论塔板数为10-30块。
根据本发明，优选地，第二精馏塔的操作条件包括：塔釜温度为120-320℃，塔顶温度为40-85℃，回流比为0.5-30，塔顶操作压力为-0.1MPaG至0.5MPaG，理论塔板数为10-45块。
根据本发明，优选地，第三精馏塔的操作条件包括：塔釜温度为150-360℃，塔顶温度为40-85℃，回流比为1-65，塔顶操作压力为-0.1MPaG至0MPaG，理论塔板数为15-85块。
根据本发明，含环己亚胺和水的物流进行分离的方式可以为本领域中常用的脱水 方式，但为了进一步降低含环己亚胺的物流中的水含量，优选地，脱水处理的方式包常压精馏、减压精馏、变压精馏、共沸精馏、膜分离、萃取精馏、吸附法脱水、生物处理脱水和离心分离中的至少一种。共沸精馏中采用的共沸剂可以为环己烷、正己烷、三甲基戊烷、对甲基异丙苯、1,4-二氧六环、苯酚、甲酚、丁醚、戊醚和异戊醚中的至少一种。根据具体共沸剂与水和环己亚胺形成的共沸点不同来具体选择共沸精馏的操作条件。
优选地，返回至步骤(1)的含环己亚胺的物流中水含量小于3重量％，优选小于1.5重量％，更优选小于1重量％。
根据本发明的一种优选实施方式，含环己亚胺和水的物流进行脱水处理可以采用膜分离脱水，膜分离脱水的操作条件包括：入膜压力为0-5MPaG，例如1-4MPaG，温度为10-350℃，例如10-200℃，例如15-190℃。
根据本发明的一种优选实施方式，含环己亚胺和水的物流进行脱水处理可以采用共沸精馏脱水，共沸精馏脱水在精馏塔中进行，精馏塔的操作条件包括：共沸剂与含环己亚胺的物流的重量比为10-100:1，塔釜温度为120-200℃，塔顶温度为80-150℃，回流比为0.1-20，塔顶操作压力为0-3MPaG，塔板数为10-80块。
根据本发明的一种优选实施方式，含环己亚胺和水的物流进行脱水处理可以同时采用精馏和膜分离或变压精馏。例如，将含环己亚胺和水的物流在精馏塔中进行精馏，塔釜得到含环己亚胺的物流，塔顶物流进行膜分离或者变压精馏得到废水。优选地，将含环己亚胺和水的物流精馏的条件包括：塔釜温度为60-250℃，塔顶操作压力为-0.09MPaG至1MPaG。优选地，膜分离的条件包括入膜温度80-350℃，入膜压力0-5MPaG。变压精馏的变压操作范围可以为0-8MPaG，优选为0-5MPaG；更优选地，变压精馏在高压塔和低压塔中进行，高压塔的塔顶压力为1-8MPaG，塔釜温度为150-240℃；低压塔的塔顶压力小于1MPaG，塔釜温度为50-130℃。
根据本发明，优选地，在分离氨化反应产物中的含环己亚胺和水的物流、己二胺产品、含氨基己醇的物流和重组分之前，先对氨化反应产物进行预分离，回收氨化反应产物中的氢气和氨；其中，回收的氢气和氨先经过冷凝得到循环氨(液相)，然后经过压缩得到循环氢气(气相)，循环氢气和循环氨返回至氨化反应中。优选地，预分离的方式包括闪蒸、精馏和汽提中的至少一种。预分离的方法也可以采用其它能够从氨化反应产物中回收循环氢气和氨的方法。更优选地，循环氢气中氢气的含量为5-40重量％，氨的含量为60-95重量％；循环氨中氨的含量为40-100重量％，水的含量为 0-1重量％，己二胺的含量为0-1重量％，环己亚胺的含量为0-1重量％。由于氨化反应过程中会消耗部分氨和氢气，氢气和氨回收过程中也会有少量损耗，因此还需要在氨化反应进行时补充新鲜的氨和氢气以维持氨化反应体系中己二醇(包括新鲜补充的接触和循环物流中的己二醇)、氨和氢气的摩尔比。
为保证氢气和氨的有效利用，降低工艺整体能耗，优选情况下，预分离的条件使得氢气的回收率大于99％，氨的回收率大于98％。本发明中，预分离的方法可以采用多级减压闪蒸，多级闪蒸的方式可以包括至少2级闪蒸，闪蒸的压力以1-8MPa的梯度递减，最后一级闪蒸的压力为0.5-1MPaG。例如，采用四级闪蒸，第一级到第四级的闪蒸压力分别为7-15MPaG、4-7MPaG、1-4MPaG、0.1-1MPaG。例如，采用三级闪蒸，第一级到第三级的闪蒸压力分别为7-15MPaG、1-7MPaG、0.5-3MPaG。
根据本发明，在制备己二胺的方法中，步骤(4)还包括：在氨解反应前，将重组分溶解在第二溶剂中，然后进行氨解。第二溶剂为己二醇、己二胺、氨基己醇、环己亚胺、双六甲撑三胺等氨化反应原料和反应产物在其中溶解度大于0.05g/g的溶剂，比如四氢呋喃、1,4-二氧六环、正己烷、环己烷和叔丁醇等中的至少一种。第二溶剂与重组分的重量比为1-20:1，优选为2-18:1，更优选为3-15:1。
溶解重组分的第二溶剂可与氨化反应(如果存在溶剂的话)中采用的第一溶剂种类相同。溶解重组分的溶剂可以是新鲜引入的溶剂，也可以是氨化反应产物(如果氨化反应存在溶剂的话)中分离得到的溶剂物流。通常情况下，在反应初始阶段，系统中溶剂的量较少，此时可以采用新鲜的溶剂溶解重组分。当反应运行一段时间后，系统中溶剂的量较多，氨化反应产物中分离得到的溶剂物流能够满足氨化反应和氨解反应所需要的溶剂的量，此时系统不再引入新鲜溶剂，将氨化反应产物中分离得到的溶剂物流的一部分返回氨化反应，另一部分用于溶解重组分。由于反应运行系统中循环的溶剂会不可避免地损失，因此定期向系统中补充新鲜溶剂。
根据本发明，在步骤(3)中，氨解在氢气和氨存在下进行，氢气：氨：重组分的摩尔比为0.1-20:10-150:1，优选为0.2-15:15-120:1，更优选为0.2-10:20-100:1，最优选为0.3-8:30-90:1，其中，重组分以双六甲撑三胺计。
根据本发明，优选地，氨解反应的条件包括：反应温度为120-300℃，优选为150-270℃；反应压力为9-25MPaG，优选为10-22MPaG；重组分的液体体积空速为0.01-8h-1，优选为0.05-5h-1。
本发明中，氨解反应中采用的氨解催化剂可以按照如下方法制备得到：
(1)制备载体：将拟薄水铝石、硅溶胶和硝酸钙的混合物与含硝酸和磷酸的水溶液接触，再依次进行捏合、干燥和焙烧；
(2)将载体浸渍于含硫酸镍、醋酸镧和硝酸铟的水溶液中，在100-140℃下干燥2-6h后，然后在350-450℃下焙烧2-6h。
根据本发明的氨解催化剂的制备方法，优选地，步骤(1)中，相对于每克的拟薄水铝石，硅溶胶的用量为0.6-0.8g，硝酸钙的用量为0.1-0.4g，含硝酸和磷酸的水溶液0.2-0.5g。
根据本发明的氨解催化剂的制备方法，优选地，步骤(1)中，含硝酸和磷酸的水溶液中硝酸的含量为10-25重量％，磷酸的含量为5-15重量％。
根据本发明的氨解催化剂的制备方法，优选地，步骤(1)中，干燥的温度为100-140℃，干燥的时间为1-6h。
根据本发明的氨解催化剂的制备方法，优选地，步骤(2)中，相对于每克的载体，硫酸镍的用量为0.6-0.85g，醋酸镧的用量为0.06-0.08g，硝酸铟的用量为0.055-0.065g。
根据本发明的氨解催化剂的制备方法，优选地，步骤(2)中，含硫酸镍、醋酸镧和硝酸铟的水溶液中硫酸镍的浓度为20-25重量％，醋酸镧的浓度为1.5-2.5重量％，硝酸铟的浓度为1.5-2.5重量％。浸渍法优选为等体积浸渍法，浸渍可以分多次进行。
本发明还发现，通过使含氨基己醇的物流、含环己亚胺的物流以一定比例返回氨化反应中，并将氨化反应中的重组分进行氨解，不仅大大提高了己二胺的收率，还减少了氨化反应废液的排放量，极大地降低己二胺生产成本。
氨化反应中存在第一溶剂
在本发明的的制备己二胺的方法中，
步骤(1)包括：在氨化反应条件下使己二醇和氨在第一溶剂中发生氨化反应得到氨化反应产物；
步骤(2)包括从步骤(1)的氨化反应产物中分离出含环己亚胺、第一溶剂和水的物流、己二胺产品、含氨基己醇的物流和重组分。
在本发明的一个实施方案中，制备己二胺的方法包括以下步骤：
(1)在氨化反应条件下，使己二醇与氨在第一溶剂中发生氨化反应得到氨化反应产物；
(2)从步骤(1)的氨化反应产物中分离出含环己亚胺、第一溶剂和水的物流、 己二胺产品、含氨基己醇的物流和重组分；
(3)在氨解反应条件下，将重组分进行氨解得到氨解反应产物，所述氨解反应产物返回至步骤(2)。
在本发明的一个实施方案中，制备己二胺的方法包括以下步骤：
(1)在氨化反应条件下，使己二醇与氨在第一溶剂中发生氨化反应得到氨化反应产物；
(2)从步骤(1)的氨化反应产物中分离出含环己亚胺、第一溶剂和水的物流、己二胺产品、含氨基己醇的物流和重组分，然后将含环己亚胺、第一溶剂和水的物流进行分离得到含环己亚胺的物流和第一溶剂物流，将至少部分的含环己亚胺的物流和至少部分的第一溶剂物流返回至步骤(1)；
(3)在氨解反应条件下，将重组分进行氨解得到氨解反应产物，所述氨解反应产物返回至步骤(2)。
在本发明的一个实施方案中，制备己二胺的方法包括以下步骤：
(1)在氨化反应条件下，使己二醇与氨在第一溶剂中发生氨化反应得到氨化反应产物；
(2)从步骤(1)的氨化反应产物中分离出含环己亚胺、第一溶剂和水的物流、己二胺产品、含氨基己醇的物流和重组分，将至少部分的含氨基己醇的物流返回至步骤(1)，然后将含环己亚胺、第一溶剂和水的物流进行分离得到含环己亚胺的物流和第一溶剂物流，将至少部分的含环己亚胺的物流和至少部分的第一溶剂物流返回至步骤(1)；
(3)在氨解反应条件下，将重组分进行氨解得到氨解反应产物，所述氨解反应产物经过或不经过分离后返回至步骤(2)。
在氨化反应中引入溶剂，可以维持体系温度分布均匀，有利于减少反应热点，有助于提高反应稳定性，并保证体系各处不发生固相结晶析出，降低了装置中各设备的操作温度，避免了装置保温而带来的巨大投资。另外，引入溶剂能够有效溶解己二醇、氨化反应产生的环己亚胺、氨基己醇等副产物，有助于解决整个操作周期内、特别是分离过程中出现的管路堵塞的问题。
在本发明中，第一溶剂为一次性引入(在反应初始阶段随己二醇引入)，在反应初始阶段(循环物流返回氨化反应之前)，第一溶剂与己二醇的摩尔比为0.1-10:1，优选0.15-10:1，更优选为0.25-9:1。但可以理解的是，第一溶剂在循环和净化过程中 会有损耗，为保证体系中循环溶剂的量，可以在第一溶剂循环过程中补充少量新鲜溶剂维持系统中溶剂的量恒定。根据本发明，第一溶剂为己二醇、己二胺、氨基己醇、环己亚胺等氨化反应原料和反应产物在其中溶解度大于0.1g/g的溶剂，比如四氢呋喃、1,4-二氧六环、正己烷、环己烷、叔丁醇、环丁砜和甘油等中的至少一种。
优选地，在步骤(1)中，在氨化反应过程中，氨与己二醇的摩尔比为10-80:1，优选为20-45:1；氨化反应在氢气存在下进行，氢气与己二醇的摩尔比为0.05-20:1，优选为0.2-5:1。其中，氨、氢气与己二醇的摩尔比是指氨化反应器入口的混合物中的摩尔比。
优选地，在步骤(1)中，氨化反应的温度为100-250℃，优选为130-210℃；氨化反应的压力为6-18MPaG，优选为8-16MPaG，新鲜己二醇的液体体积空速为0.01-10h-1，优选为0.05-4h-1。
优选地，返回步骤(1)中的含环己亚胺的物流和含氨基己醇的物流的重量比为0.1-5:1，例如0.2-4.5:1、0.3-4:1、0.4-3.5:1、0.5-3:1、0.6-2.5:1、0.7-2:1、0.8-1.5:1，从而进一步提高己二胺的收率。由于溶剂仅起溶解反应原料和反应产物的作用，因此对返回溶剂的量没有限制，可以全部返回，也可以部分返回。当一部分溶剂返回步骤(1)时，例如，含环己亚胺的物流、第一溶剂物流和含氨基己醇的物流的重量比可以为0.1-5:0.3-40:1，优选为0.1-2.5:0.3-10:1，更优选为0.5-2:1-10:1。
优选地，返回步骤(1)中的含环己亚胺的物流和含氨基己醇的物流的总重量与新鲜己二醇的重量比为0.1-30:1，优选为0.3-25:1，例如0.4-20:1、0.5-15:1、0.8-10:1、1-8:1、1.5-6:1、2-5:1，从而进一步提高己二胺的收率。
优选地，步骤(2)还包括：在分离氨化反应产物中含环己亚胺、第一溶剂和水的物流、己二胺产品、含氨基己醇的物流和重组分之前，回收氨化反应产物中的氢气和氨；其中，回收的氢气和氨先经过冷凝得到循环氨(液相)，然后经过压缩得到循环氢(气相)，循环氢和循环氨返回至氨化反应中。更优选地，循环氢气中氢气的含量为5-40重量％，氨的含量为60-95重量％；循环氨中氨的含量为40-100重量％，溶剂的含量为0-60重量％，水的含量为0-1重量％，己二胺的含量为0-1重量％，环己亚胺的含量为0-1重量％。由于氨化反应过程中会消耗部分氨和氢气，氢气和氨回收过程中也会有少量损耗，因此还需要在氨化反应进行时补充新鲜的氨和氢气以维持氨化反应体系中己二醇(包括新鲜补充的接触和循环物流中的己二醇)、氨和氢气的摩尔比。
优选地，步骤(2)包括：任选脱除氨、氢气的氨化反应产物进入第一精馏塔中进行分离，第一精馏塔的塔顶得到含环己亚胺、水和第一溶剂的物流；第一精馏塔的塔釜物流送入第二精馏塔进行分离，塔顶得到己二胺产品，塔釜得到含己二醇、氨基己醇和重组分的物流；第二精馏塔的塔釜物流送入第三精馏塔进行分离，第三精馏塔的塔顶得到含己二醇和氨基己醇的物流，塔釜得到重组分。
优选地，第一精馏塔的操作条件包括：塔釜温度为90-300℃，塔顶温度为30-65℃，回流比为0.1-10，塔顶操作压力为-0.1MPaG至0.5MPaG，塔板数为10-30块。
优选地，第二精馏塔的操作条件包括：塔釜温度为120-320℃，塔顶温度为40-85℃，回流比为0.5-30，塔顶操作压力为-0.1MPaG至0.5MPaG，塔板数为10-45块。
优选地，第三精馏塔的操作条件包括：塔釜温度为150-360℃，塔顶温度为40-85℃，回流比为1-65，塔顶操作压力为-0.1MPaG至0MPaG，塔板数为15-85块。
优选地，步骤(2)包括：将含环己亚胺、第一溶剂和水的物流在精馏塔中进行精馏，塔釜得到含环己亚胺的物流，塔顶物流进行膜分离或者变压精馏得到溶剂物流和废水。优选地，将含环己亚胺、第一溶剂和水的物流精馏的条件包括：塔釜温度为60-250℃，塔顶操作压力为-0.09MPaG至1MPaG。优选地，膜分离的条件包括入膜温度80-350℃，入膜压力0-5MPaG。变压精馏的变压操作范围可以为0-8MPaG，优选为0-5MPaG；更优选地，变压精馏在高压塔和低压塔中进行，高压塔的塔顶压力为1-8MPaG，塔釜温度为150-240℃；低压塔的塔顶压力小于1MPaG，塔釜温度为50-130℃。
优选地，氨解反应产物通过精馏进行粗分离，进行粗分离的粗分塔的条件包括：塔釜温度为380-520℃，塔顶操作压力为-0.05MPag至1MPag，回流比为2-50。
氨化反应中不存在第一溶剂
在本发明的制备己二胺的方法中，
步骤(1)包括：在氨化反应条件下，在不存在第一溶剂下，使己二醇和氨发生氨化反应得到氨化反应产物。
在本发明的一个实施方案中，制备己二胺的方法包括以下步骤：
(1)在氨化反应条件下，在不存在第一溶剂下，使己二醇与氨发生氨化反应得到氨化反应产物；
(2)从步骤(1)的氨化反应产物中分离出含环己亚胺和水的物流、己二胺产品、含氨基己醇的物流和重组分；
(3)在氨解反应条件下，将重组分进行氨解得到氨解反应产物，所述氨解反应产物返回至步骤(2)。
在本发明的一个实施方案中，制备己二胺的方法包括以下步骤：
(1)在氨化反应条件下，在不存在第一溶剂下，使己二醇与氨发生氨化反应得到氨化反应产物；
(2)从步骤(1)的氨化反应产物中分离出含环己亚胺和水的物流、己二胺产品、含氨基己醇的物流和重组分，然后将含环己亚胺和水的物流进行分离得到含环己亚胺的物流，将至少部分的含环己亚胺的物流返回至步骤(1)；
(3)在氨解反应条件下，将重组分进行氨解得到氨解反应产物，所述氨解反应产物返回至步骤(2)。
在本发明的一个实施方案中，制备己二胺的方法包括以下步骤：
(1)在氨化反应条件下，在不存在第一溶剂下，使己二醇与氨发生氨化反应得到氨化反应产物；
(2)从步骤(1)的氨化反应产物中分离出含环己亚胺和水的物流、己二胺产品、含氨基己醇的物流和重组分，将至少部分的含氨基己醇的物流返回至步骤(1)，然后将含环己亚胺和水的物流进行分离得到含环己亚胺的物流，将至少部分的含环己亚胺的物流返回至步骤(1)；
(3)在氨解反应条件下，将重组分进行氨解得到氨解反应产物，所述氨解反应产物经过或不经过分离后返回至步骤(2)。
优选地，在步骤(1)中，氨与己二醇的摩尔比为18-50:1，优选为22-38:1，氢气与己二醇的摩尔比为0.08-6:1，优选为0.4-3:1。其中，氨、氢气与己二醇的摩尔比是指氨化反应器入口的混合物中的摩尔比。
优选地，在步骤(1)中，氨化反应的温度为120-210℃，优选为130-200℃。氨化反应的压力为7-17MPaG，优选为9-16MPaG。新鲜己二醇的液体体积空速为0.05-7h-1，优选为0.09-3.9h-1。
优选地，返回步骤(1)中的含环己亚胺的物流与新鲜己二醇的重量比为0.1-13:1，优选为0.3-10:1。
实施例
以下将通过实施例对本发明进行详细描述。以下实施例中，
产物的组成通过气相色谱进行分析。
己二胺的收率＝己二胺的摩尔量÷新鲜进料的己二醇的摩尔量×100％。
制备例1
通过多步浸渍法制备氨解催化剂：
(1)称量拟薄水铝石(硫酸铝法生产，比表面积310m2/g，孔容1.19ml/g)94.2g、硅溶胶(JN-40)72.5g和四水合硝酸钙25.26g。将拟薄水铝石置于捏合机中，将称量的硅溶胶和四水合硝酸钙加入24.77g水中配成溶液，并加入捏合机中与拟薄水铝石中充分搅拌，再加入由16.51g水、4.71g硝酸和2.83g磷酸配成的水溶液充分搅拌，然后捏合挤出成三叶草型，将其在120℃下干燥4h，接着在马弗炉中900℃下焙烧6h，降温后制得载体。
(2)将100.77g六水合硫酸镍(工业级，纯度98％)、5.69g一水合醋酸镧和5.96g五水合硝酸铟加入水134.78mL配成水溶液，分两次用等体积浸渍法将溶液负载在步骤(1)获得的73.25g载体上，每次浸渍后需在120℃下干燥4小时，两次浸渍完毕后在390℃下焙烧4小时。
实施例1
(1)将溶剂(环己烷)、己二醇、氨、氢原料送入装填有催化剂的固定床氨化反应器，催化剂采用CN114433086A中实施例14的催化剂，氨与己二醇的摩尔比35:1，氢气与己二醇的摩尔比0.5:1，溶剂与己二醇的摩尔比3.8:1，氨化反应温度为166℃，氨化反应压力为12MPaG，新鲜己二醇的液体体积空速为0.1h-1。初始氨化反应的产物(循环物流返回固定床氨化反应器之前)组成包括：己二胺5.76wt％，己二醇1.13wt％，氨基己醇4.91wt％，环己亚胺4.76wt％，氨48.1wt％，环己烷31.59wt％，水2.56wt％，其它为重组分。初始氨化反应中己二胺的摩尔收率为31.1％。
(2)将氨化反应产物进行气液分离，脱除其中的氢气、氨，将氢气以气相形式经压缩机返回固定床氨化反应器，将液氨泵送回氨化反应器。气液分离的方法包括：反应产物经12MPaG、5MPaG、2MPaG、0.5MPaG四级闪蒸罐气液分离，其中第二、第三级闪蒸罐顶气相冷凝后以0.1回流比回流，第四级进闪蒸罐前加热，罐内设1m填料。闪蒸得到的气相再经30-45℃的冷凝水冷凝后得到循环氨(液相)和气相(氢气)，气相经压缩后得到循环氢气，循环氢气主要组成为氢气9.72wt％，其它为氨； 循环氨主要组成为氨94.96wt％，环己烷4.12wt％，水0.73wt％，己二胺0.01wt％，环己亚胺0.15wt％。
脱除氨、氢气的氨化反应产物进入第一精馏塔中进行减压精馏，第一精馏塔的塔顶得到含环己亚胺、水和溶剂的物流(物流组成为：环己烷80.11wt％，环己亚胺13.01wt％，水5.79wt％)；第一精馏塔的操作条件包括：塔釜温度为179.9℃，塔顶温度为40.5℃，回流比为1.5，塔顶操作压力为-0.08MPa，塔板数为25块。第一精馏塔的塔釜物流送入第二精馏塔进行分离，塔顶得到己二胺产品(产品中己二胺的含量为99.8重量％)，塔釜得到含己二醇、氨基己醇和重组分的物流(物流组成为氨基己醇79.73wt％，己二醇18.35wt％，己二胺0.2wt％，其它为重组分碳十二胺)；第二精馏塔的操作条件包括：塔釜温度为189.1℃，塔顶温度为83℃，回流比为2.7，塔顶操作压力为-0.09MPa，塔板数为24块。第二精馏塔的塔釜物流送入第三精馏塔进行分离，第三精馏塔的塔顶得到含己二醇和氨基己醇的物流，塔釜得到重组分；第三精馏塔的操作条件包括：塔釜温度为267.5℃，塔顶温度为83.7℃，回流比为5.3，塔顶操作压力为-0.09MPa，塔板数为38块。
(3)将步骤(2)得到的含环己亚胺、水和溶剂的物流进行精馏得到含环己亚胺的物流(物流组成为环己亚胺含量为99.91wt％，己二胺0.09wt％)和含溶剂和水的物流，精馏的条件包括塔釜温度为94.3℃，塔顶操作压力为-0.08MPa。采用膜分离对含溶剂和水的物流进行脱水得到溶剂物流(物流组成为环己烷99.67wt％，环己亚胺0.32wt％)和废水(环己烷含量小于100ppmw)，膜分离的条件包括入膜温度100℃，入膜压力0.05MPaG。
(4)将步骤(2)中第三精馏塔塔釜得到的重组分与溶剂(环己烷)混合均匀后送入装填有催化剂的氨解反应器中进行氨解，同时向氨解反应器中通入氨和氢气，氢气：氨：以双六甲撑三胺计的重组分的摩尔比为1:40:1，溶剂与重组分的重量比为5:1，氨解反应温度为186℃，氨解反应压力为16MPaG，重组分液体体积空速为0.5h-1。获得氨解反应产物组成为：氢气0.07wt％、氨55.66wt％、己二胺7.84wt％，环己亚胺6.01wt％，环己烷21.92wt％，其它为重组分。氨解反应产物全部送回步骤(2)。
返回步骤(1)中的含环己亚胺的物流和含氨基己醇的物流的总重量与新鲜己二醇的重量比为2.88。返回步骤(1)中的含环己亚胺的物流和含氨基己醇的物流的重量比为0.72:1。
系统连续运行至稳定状态，己二胺的摩尔收率达到93.7％。系统无堵塞情况，能 够长周期连续平稳运行。
此外，含溶剂的反应器中温度分布均匀无热点，同一床层高度上分布的多个温度测点的温度值偏差小于1.2℃。
实施例2
(1)将溶剂(1,4-二氧六环)、己二醇、氨、氢原料送入装填有催化剂的固定床氨化反应器，催化剂采用CN114433086A中实施例6的催化剂，氨与己二醇的摩尔比25:1，氢气与己二醇的摩尔比1:1，溶剂与己二醇的摩尔比1.2:1，氨化反应温度为190℃，氨化反应压力为10MPaG，新鲜己二醇的液体体积空速为0.6h-1。初始氨化反应的产物(循环物流返回固定床氨化反应器之前)组成包括：己二胺8.39wt％，己二醇0.93wt％，氨基己醇5.97wt％，环己亚胺6.67wt％，氨34.83wt％，1,4-二氧六环39.02wt％，水3.12wt％，其它为重组分。初始氨化反应中己二胺的摩尔收率为36.4％。
(2)将氨化反应产物进行气液分离，脱除其中的氢气、氨，将氢气以气相形式经压缩机返回固定床氨化反应器，将液氨泵送回氨化反应器。气液分离的方法包括：反应产物经10MPaG、2MPaG、0.8MPaG三级气液分离，其中第二级闪蒸罐顶设一段填料，罐顶气相冷凝后以0.3回流比回流，第三级闪蒸罐内设1m填料与罐底加热器。闪蒸得到的气相再经30-45℃的冷凝水冷凝后得到循环氨(液相)和气相(氢气)，气相经压缩后得到循环氢气，循环氢气主要组成为氢气10.19wt％，其它为氨；循环氨主要组成为氨93wt％，1,4-二氧六环5.79wt％，水0.91wt％，己二胺0.01wt％，环己亚胺0.21wt％。
脱除氨、氢气的氨化反应产物进入第一精馏塔中进行减压精馏，第一精馏塔的塔顶得到含环己亚胺、水和溶剂的物流(物流组成为：1,4-二氧六环79.93wt％，环己亚胺14.40wt％，水5.86wt％)；第一精馏塔的操作条件包括：塔釜温度为176.9℃，塔顶温度为44.2℃，回流比为2.6，塔顶操作压力为-0.08MPa，塔板数为19块。第一精馏塔的塔釜物流送入第二精馏塔进行分离，塔顶得到己二胺产品(产品中己二胺的含量为99.9重量％)，塔釜得到含己二醇、氨基己醇和重组分的物流(物流组成为：氨基己醇85.36wt％，己二醇13.33wt％，己二胺0.25wt％，其它为重组分碳十二胺)；第二精馏塔的操作条件包括：塔釜温度为188.3℃，塔顶温度为84.3℃，回流比为1.9，塔顶操作压力为-0.09MPa，塔板数为33块。第二精馏塔的塔釜物流送入第三精馏塔进行分离，第三精馏塔的塔顶得到含己二醇和氨基己醇的物流，塔釜得到重组分；第 三精馏塔的操作条件包括：塔釜温度为270.5℃，塔顶温度为83.6℃，回流比为5，塔顶操作压力为-0.09MPa，塔板数为39块。
(3)将步骤(2)得到的含环己亚胺、水和溶剂的物流进行精馏得到含环己亚胺的物流(物流组成为：环己亚胺99.94wt％，己二胺0.06wt％)和含溶剂和水的物流，精馏的条件包括塔釜温度为207.3℃，塔顶操作压力为0.5MPaG，采用变压精馏对含溶剂和水的物流进行脱水得到溶剂物流(物流组成为：1,4-二氧六环99.64wt％，环己亚胺0.36wt％)和废水(水99.95wt％，1,4-二氧六环0.05wt％)，变压精馏的条件包括高压塔塔顶压力1MPaG，塔釜温度197.2℃，低压塔塔顶压力0MPaG，塔釜温度96.3℃。
(4)将步骤(2)中第三精馏塔塔釜得到的重组分与溶剂(1,4-二氧六环)混合均匀后送入装填有催化剂的氨解反应器中进行氨解，同时向氨解反应器中通入氨和氢气，氢气：氨：以双六甲撑三胺计的重组分的摩尔比为3:50:1，溶剂与重组分的重量比为11:1，氨解反应温度为203℃，氨解反应压力为15MPaG，重组分液体体积空速为0.8h-1。获得氨解反应产物组成为：氢气4.75wt％、氨71.74wt％、己二胺2.60wt％，环己亚胺3.52wt％，1,4-二氧六环15.26wt％，其它为重组分。将氨解反应产物减压到10MPag，在该压力下经闪蒸进行气液分离，得到含氢气和氨的物流和剩余液相，剩余液相再经过粗分塔精馏得到含己二胺和环己亚胺的混合液和重组分，粗分塔的塔顶压力为-0.07MPag，塔顶温度为180℃，塔釜温度为480℃，回流比为20。其中粗分塔塔顶的含己二胺和环己亚胺的混合液返回步骤(2)，塔釜物流作为废液排出。
返回步骤(1)中的含环己亚胺的物流和含氨基己醇的物流的总重量与新鲜己二醇的重量比为2.95。返回步骤(1)中的含环己亚胺的物流和含氨基己醇的物流的重量比为0.97:1。
系统连续运行至稳定状态，己二胺的摩尔收率达到94.8％。系统无堵塞情况，能够长周期连续平稳运行。
此外，含溶剂的反应器中温度分布均匀无热点，同一床层高度上分布的多个温度测点的温度值偏差小于1.5℃。
实施例3
(1)将氨、氢气与己二醇送入装填有催化剂的固定床氨化反应器，催化剂采用按照CN114433113A中实施例2的催化剂，其中，氨与己二醇摩尔比为35:1，氢气与 己二醇摩尔比为2:1，氨化反应温度为150℃，反应压力为12MPa，己二醇液相体积空速为1h-1。
(2)将氨化反应产物进行气液分离，脱除其中的氢气、氨，将氢气以气相形式经压缩机返回固定床氨化反应器，将液氨泵送回氨化反应器。气液分离的方法包括：将氨化反应的产物依次通过3个闪蒸罐进行3级减压闪蒸回收氢气，其中，第一个闪蒸罐至第3个闪蒸罐的压力依次设置为7MPa、5MPa、2MPa，三个闪蒸罐顶气相分别冷却至45℃后进一步进行气液分离，得到的液相返回前一个减压闪蒸罐，气相升压至氨化反应压力后返回氨化反应器。第三个减压闪蒸罐底液相从塔顶进入脱氨精馏塔，塔板数为11，塔顶操作压力为2MPa，出于节能降耗考虑，脱氨精馏塔塔顶不设冷凝器；脱氨精馏塔塔顶物流加压至氨化反应压力后送回氨化反应器入口，塔釜物流进行后续精制。
脱除氨、氢气的氨化反应产物依据沸点不同，以己二胺为关键组分进行切割分离，依次进行第一分离得到含环己亚胺的物流和剩余物流，然后剩余物流进行第二分离得到己二胺产品和含氨基己醇、C12胺和重组分的物流，然后含氨基己醇、C12胺和重组分的物流进行第三分离得到含氨基己醇的物流、含C12胺的物流和重组分。含氨基己醇的物流与含环己亚胺的物流按照1:1.25的重量比混合后作为循环物流返回至步骤(1)，重组分抽出。其中，第一分离在第一精馏塔中进行，第一精馏塔的操作条件包括：塔釜温度为187℃，塔顶温度为50℃，回流比为2.1，理论塔板数为18，塔顶操作压力为-0.07MPa；第二分离在第二精馏塔中进行，第二精馏塔的操作条件包括：塔釜温度为195℃，塔顶温度为131℃，回流比为5，理论塔板数为19，塔顶操作压力为-0.08MPa；第三分离在第三精馏塔中进行，第三精馏塔的操作条件包括：塔釜温度为338℃，塔顶温度为87℃，回流比为5，理论塔板数为55，塔顶操作压力为-0.09MPa。
(3)将步骤(2)中第三精馏塔塔釜得到的重组分送入装填有催化剂的氨解反应器中进行氨解，同时向氨解反应器中通入氨和氢气，氢气：氨：以双六甲撑三胺计的重组分的摩尔比为5:100:1，氨解反应温度170℃，氨解反应压力12MPag，含C12胺的物流的液体空速为0.5h-1。氨解的产物返回至步骤(2)。
反应稳定运行时各物流的重量组成如表1所示。己二胺的摩尔收率为90.69％。
表1

注：11：氨化反应器进口处的物流组成；12：氨化反应产物；13：循环氢气和氨；14：
循环物流；15：己二胺产品；16：重组分；17：含C12胺的物流；19：轻组分；表中“/”表示含量小于500ppm，或者含量为0。
对比例1
按照实施例1的方法进行，不同的是，得到的重组分不进行氨解，反应稳定运行时，己二胺的摩尔收率为86.5％。
对比例2
按照实施例2的方法进行，不同的是，得到的重组分不进行氨解，反应稳定运行时，己二胺的摩尔收率为87.7％。
对比例3
按照实施例3的方法进行，不同的是，得到的重组分不进行氨解，反应稳定运行时，己二胺的摩尔收率为83.91％。
以上详细描述了本发明的优选实施方式，但是，本发明并不限于此。在本发明的技术构思范围内，可以对本发明的技术方案进行多种简单变型，包括各个技术特征以任何其它的合适方式进行组合，这些简单变型和组合同样应当视为本发明所公开的内容，均属于本发明的保护范围。

Claims (19)

  1. 一种制备己二胺的方法，该方法包括以下步骤：
    (1)在氨化反应条件下，任选在第一溶剂存在下，使己二醇与氨发生氨化反应得到氨化反应产物；
    (2)从步骤(1)的氨化反应产物中分离得到己二胺产品和重组分；
    (3)在氨解反应条件下，将重组分进行氨解得到氨解反应产物，所述氨解反应产物返回至步骤(2)。
  2. 根据权利要求1所述的方法，其中在氨化反应过程中，氨与己二醇的摩尔比为10-90:1，优选为15-70:1；氨化反应在氢气存在下进行，氢气与己二醇摩尔比为0.05-25:1，优选为0.1-10:1。
  3. 根据权利要求1或2所述的方法，其中氨化反应的温度为100-300℃，优选为120-250℃；氨化反应的压力为4-25MPaG，优选为6-20MPaG，新鲜己二醇的液体体积空速为0.01-15h-1，优选为0.03-8h-1。
  4. 根据权利要求1-3中任一项所述的方法，其中步骤(2)还包括：
    从氨化反应产物中分离得到含环己亚胺和水的物流，和任选地将所述含环己亚胺和水的物流进行分离得到含环己亚胺的物流，将至少部分的所述含环己亚胺的物流返回至步骤(1)；优选地，返回步骤(1)中的所述含环己亚胺的物流与新鲜己二醇的重量比为0.1-26:1，更优选为0.3-24:1；和/或
    从氨化反应产物中分离得到含氨基己醇的物流，和任选地将至少部分的所述含氨基己醇的物流返回至步骤(1)；优选地，返回步骤(1)中的所述含氨基己醇的物流与新鲜己二醇的重量比为0.1-26:1，更优选为0.3-24:1。
  5. 根据权利要求1-3中任一项所述的方法，其中步骤(2)还包括：
    从氨化反应产物中分离得到含环己亚胺和水的物流，和将所述含环己亚胺和水的物流进行分离得到含环己亚胺的物流，将至少部分的所述含环己亚胺的物流返回至步骤(1)；和
    从氨化反应产物中分离得到含氨基己醇的物流，和将至少部分的所述含氨基己醇的物流返回至步骤(1)；
    优选地，返回步骤(1)中的所述含环己亚胺的物流和所述含氨基己醇的物流的重量比为0.1-5:1；
    优选地，返回步骤(1)中的所述含环己亚胺的物流和所述含氨基己醇的物流的总重量与新鲜己二醇的重量比为0.1-30:1，优选为0.3-25:1。
  6. 根据权利要求4或5所述的方法，其中所述含环己亚胺和水的物流的分离方式包括常压精馏、减压精馏、变压精馏、共沸精馏、膜分离和离心分离中的至少一种；
    优选的，共沸精馏中采用的共沸剂为环己烷、正己烷、三甲基戊烷、对甲基异丙苯、1,4-二氧六环、苯酚、甲酚、丁醚、戊醚和异戊醚中的至少一种。
  7. 根据权利要求1-6中任一项所述的方法，其中氨化反应产物的分离方式为：将氨化反应产物进行第一分离得到含环己亚胺和水的物流，接着进行第二分离得到己二胺产品和含氨基己醇和重组分的物流，然后进行第三分离得到含氨基己醇的物流和重组分；
    优选地，所述第一分离、第二分离和第三分离的方式各自独立地包括精馏、膜分离和变压吸附中的至少一种。
  8. 根据权利要求1-7中任一项所述的方法，其中，在分离氨化反应产物之前，先对氨化反应产物进行预分离得到循环氢气和氨，其中，循环氢气和氨返回至氨化反应中；
    优选地，所述预分离的方式包括闪蒸、精馏和汽提中的至少一种。
  9. 根据权利要求1-8中任一项所述的方法，其中步骤(3)还包括：在氨解反应前，将重组分溶解在第二溶剂中，然后进行氨解；所述第二溶剂为四氢呋喃、1,4-二氧六环、正己烷、环己烷和叔丁醇中的至少一种；所述第二溶剂与重组分的重量比0.1-20:1。
  10. 根据权利要求1-9中任一项所述的方法，其中在步骤(3)中，氨解在氢气和 氨的存在下进行，氢气：氨：重组分的摩尔比为0.1-20:10-150:1，其中，重组分以双六甲撑三胺计。
  11. 根据权利要求1-10中任一项所述的方法，其中氨解反应的条件包括：反应温度为120-300℃，优选为150-270℃；反应压力为9-25MPaG，优选为10-22MPaG；重组分的液体体积空速为0.01-8h-1，优选为0.05-5h-1。
  12. 根据权利要求1-11中任一项所述的方法，其中步骤(1)包括在氨化反应条件下，使己二醇与氨在第一溶剂中发生氨化反应得到氨化反应产物；步骤(2)包括从步骤(1)的氨化反应产物中分离出含环己亚胺、第一溶剂和水的物流、己二胺产品、含氨基己醇的物流和重组分，将至少部分的所述含氨基己醇的物流返回至步骤(1)，然后将所述含环己亚胺、第一溶剂和水的物流进行分离得到含环己亚胺的物流和第一溶剂物流，将至少部分的所述含环己亚胺的物流和至少部分的第一溶剂物流返回至步骤(1)。
  13. 根据权利要求12所述的方法，其中在氨化反应过程中，氨与己二醇摩尔比为10-80:1，优选为20-45:1；氨化反应在氢气存在下进行，氢气与己二醇摩尔比为0.05-20:1，优选为0.2-5:1。
  14. 根据权利要求12或13所述的方法，其中氨化反应的温度为100-250℃，优选为130-210℃；氨化反应的压力为6-18MPaG，优选为8-16MPaG，新鲜己二醇的液体体积空速为0.01-10h-1，优选为0.05-4h-1。
  15. 根据权利要求12-14中任一项所述的方法，其中所述第一溶剂为四氢呋喃、1,4-二氧六环、正己烷、环己烷和叔丁醇中的至少一种；
    和/或，第一溶剂与己二醇摩尔比为0.1-10:1。
  16. 根据权利要求12-15中任一项所述的方法，其中步骤(2)包括：将所述含环己亚胺、第一溶剂和水的物流在精馏塔中进行精馏，塔釜得到含环己亚胺的物流，塔顶物流进行膜分离或者变压精馏得到溶剂物流和废水。
  17. 根据权利要求1-11中任一项所述的方法，其中步骤(1)包括在氨化反应条件下，在不存在第一溶剂下，使己二醇和氨发生氨化反应得到氨化反应产物；
    优选地，返回步骤(1)中的含环己亚胺的物流与新鲜己二醇的重量比为0.1-13:1，优选为0.3-10:1。
  18. 根据权利要求17所述的方法，其中步骤(1)中氨与己二醇摩尔比为18-50:1，优选为22-38:1，氢气与己二醇摩尔比为0.08-6:1，优选为0.4-3:1。
  19. 根据权利要求17或18所述的方法，其中氨化反应的温度为120-210℃，优选为130-200℃；氨化反应的压力为7-17MPaG，优选为8-15MPaG，新鲜己二醇的液体体积空速为0.05-7h-1，优选为0.09-3.9h-1。
PCT/CN2023/127038 2022-10-31 2023-10-27 制备己二胺的方法 WO2024093816A1 (zh)

Applications Claiming Priority (6)

Application Number Priority Date Filing Date Title
CN202211350196.8 2022-10-31
CN202211351228.6 2022-10-31
CN202211350196.8A CN117986127A (zh) 2022-10-31 2022-10-31 以己二醇为原料制备己二胺的方法
CN202211351228.6A CN117986136A (zh) 2022-10-31 2022-10-31 提高己二胺收率的方法和系统
CN202311368569.9 2023-10-20
CN202311368569 2023-10-20

Publications (1)

Publication Number Publication Date
WO2024093816A1 true WO2024093816A1 (zh) 2024-05-10

Family

ID=90929710

Family Applications (1)

Application Number Title Priority Date Filing Date
PCT/CN2023/127038 WO2024093816A1 (zh) 2022-10-31 2023-10-27 制备己二胺的方法

Country Status (1)

Country Link
WO (1) WO2024093816A1 (zh)

Citations (10)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US3268588A (en) * 1962-11-16 1966-08-23 Celanese Corp Process for producing hexamethylenediamine from 1-6-hexanediol
CA838703A (en) * 1970-04-07 Horlenko Theodore Production of hexamethylene diamine
CA892614A (en) * 1972-02-08 Horlenko Theodore Production of n-(6-aminohexyl)hexamethyleneimines
CN106807377A (zh) * 2015-11-27 2017-06-09 中国科学院大连化学物理研究所 一种用于合成己二胺的催化剂
CN106810454A (zh) * 2015-11-27 2017-06-09 中国科学院大连化学物理研究所 一种制备己二胺的方法
CN111495383A (zh) * 2020-04-22 2020-08-07 陕西延长石油（集团）有限责任公司 一种己二醇与氨气制备己二胺的方法及催化剂
CN114433113A (zh) * 2020-10-30 2022-05-06 中国石油化工股份有限公司 具有催化醇氨化功能的催化剂和载体及其制备方法和应用
CN114433087A (zh) * 2020-10-30 2022-05-06 中国石油化工股份有限公司 具有经醇临氢氨化制备胺功能的催化剂和载体及其制备方法和应用
CN114433086A (zh) * 2020-10-30 2022-05-06 中国石油化工股份有限公司 具有催化醇临氢氨化合成有机胺功能的催化剂和载体及其制备方法和应用
JP2022112396A (ja) * 2021-01-21 2022-08-02 旭化成株式会社 ヘキサメチレンジアミンの製造方法

Patent Citations (10)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CA838703A (en) * 1970-04-07 Horlenko Theodore Production of hexamethylene diamine
CA892614A (en) * 1972-02-08 Horlenko Theodore Production of n-(6-aminohexyl)hexamethyleneimines
US3268588A (en) * 1962-11-16 1966-08-23 Celanese Corp Process for producing hexamethylenediamine from 1-6-hexanediol
CN106807377A (zh) * 2015-11-27 2017-06-09 中国科学院大连化学物理研究所 一种用于合成己二胺的催化剂
CN106810454A (zh) * 2015-11-27 2017-06-09 中国科学院大连化学物理研究所 一种制备己二胺的方法
CN111495383A (zh) * 2020-04-22 2020-08-07 陕西延长石油（集团）有限责任公司 一种己二醇与氨气制备己二胺的方法及催化剂
CN114433113A (zh) * 2020-10-30 2022-05-06 中国石油化工股份有限公司 具有催化醇氨化功能的催化剂和载体及其制备方法和应用
CN114433087A (zh) * 2020-10-30 2022-05-06 中国石油化工股份有限公司 具有经醇临氢氨化制备胺功能的催化剂和载体及其制备方法和应用
CN114433086A (zh) * 2020-10-30 2022-05-06 中国石油化工股份有限公司 具有催化醇临氢氨化合成有机胺功能的催化剂和载体及其制备方法和应用
JP2022112396A (ja) * 2021-01-21 2022-08-02 旭化成株式会社 ヘキサメチレンジアミンの製造方法

Similar Documents

Publication Publication Date Title
RU2492160C2 (ru) Способ селективного гидрирования фенилацетилена в присутствии стирола с использованием композитного слоя
JP5702299B2 (ja) スチレンの存在下でフェニルアセチレンに選択的に水素を付加する方法
CN103819344B (zh) 一种1，2-丙二胺的合成方法
WO2024093816A1 (zh) 制备己二胺的方法
TW202423892A (zh) 製備己二胺的方法
US8574522B2 (en) Process for selective oxidative dehydrogenation of a hydrogen-containing CO mixed gas
CN102451674A (zh) 甲基叔丁基醚裂解制异丁烯催化剂及其制备方法和应用
CN102070464A (zh) 混合二硝基苯钯催化剂加氢还原生产苯二胺的方法
CN102093510B (zh) 一种热聚合合成间戊二烯石油树脂的方法
WO2024093807A1 (zh) 制备己二胺的方法
CN106316915B (zh) 一种邻乙基苯胺脱氢环化制吲哚的方法
TWI457313B (zh) Study on the selective hydrogenation of phenylethylene in the presence of styrene in the presence of
CN117986127A (zh) 以己二醇为原料制备己二胺的方法
CN109438396B (zh) 一种四氢呋喃-3-甲胺的制备方法
TW202423893A (zh) 製備己二胺的方法
CN117986128A (zh) 通过氨解制备己二胺的方法
CN112661620A (zh) 一种环戊酮的制备方法
CN114733569B (zh) 一种用于乙炔加氢制乙烯的共价有机框架负载钯催化剂的制备方法和应用
CN117945930A (zh) 氨氢回收以及己二醇氨化制己二胺的系统和方法
CN117417262B (zh) 一种二甘醇胺的制备方法
CN117986136A (zh) 提高己二胺收率的方法和系统
CN117945925A (zh) 己二醇生产己二胺的方法和系统
CN107774343B (zh) 催化剂的再生工艺方法
CN1706808B (zh) 甲苯一步直接氨基化合成甲苯胺
JP2002128716A (ja) イソプロピルアルコールの製造方法

Legal Events

Date Code Title Description
121 Ep: the epo has been informed by wipo that ep was designated in this application

Ref document number: 23884748

Country of ref document: EP

Kind code of ref document: A1