Process and apparatus for the production of bio-based pentanediamine

Abstract

This application discloses a method and apparatus for producing bio-based pentanediamine. The method includes: using pentanediamine fermentation broth and / or its treated broth as raw materials, decarbonizing the raw materials to obtain a decarbonized liquid, then subjecting the decarbonized liquid to alkaline hydrolysis to obtain an alkaline hydrolysate; extracting the alkaline hydrolysate using an extractant and an auxiliary agent, collecting the oil phase to obtain an extract; removing light components from the extract, then separating pentanediamine and the extractant by distillation to obtain high-purity pentanediamine; and passing the raffinate through a crystallization evaporation system to obtain crystalline salt and organic fertilizer. This application enables the continuous industrial production of pentanediamine. Extraction with a multi-component extractant and auxiliary agent transfers pentanediamine to the oil phase, and controls emulsification and entrainment in the oil phase, preventing salt from entering the oil phase and avoiding salt precipitation during subsequent refining processes that could cause equipment and pipeline blockage. High-purity pentanediamine is obtained by distillation of the oil phase, avoiding the enormous energy consumption of traditional water concentration methods.

Description

Technical Field

[0001] This application belongs to the field of pentanediamine preparation technology, specifically relating to a method and apparatus for producing bio-based pentanediamine. Background Technology

[0002] Pentylenediamine (1,5-diaminopentane) can be used as a monomer for polyamides and polyurethanes. It is an important bio-based compound that plays a vital role in replacing petrochemical raw materials and promoting the transformation of green chemical industry.

[0003] There are two main biological methods for producing pentanediamine: microbial fermentation and enzyme-catalyzed conversion. Both methods produce pentanediamine fermentation broths containing large amounts of salts, bacterial cells, cell debris, residual sugars, pigments, proteins, peptides, nucleic acids, and cell metabolites, making the system complex and difficult to separate. Furthermore, since the pentanediamine content in the fermentation broth is only 3-10%, conventional separation and purification methods require the evaporation of large amounts of water, resulting in extremely high energy consumption and production costs, which restricts the large-scale industrial production of pentanediamine. At the same time, residual sugars and proteins in the fermentation broth will denature and coke when heated, and inorganic salts will precipitate due to water evaporation, easily causing equipment and pipeline blockages, affecting the purification and production of pentanediamine. Summary of the Invention

[0004] The purpose of this application is to provide a method and apparatus for producing bio-based pentanediamine, in order to solve the problems of complex system, difficult separation, high energy consumption, and easy blockage of equipment and pipelines in the existing pentanediamine fermentation broth purification process.

[0005] To achieve the above objectives, the first aspect of this application provides a method for producing bio-based pentanediamine, comprising:

[0006] Using pentanediamine fermentation broth and / or pentanediamine fermentation broth treatment liquid as raw materials, the raw materials are decarbonized to obtain decarbonized liquid, and then the decarbonized liquid is subjected to alkaline hydrolysis to obtain alkaline hydrolysate.

[0007] The alkaline hydrolysate is extracted using an extractant and an auxiliary agent, and the oil phase is collected to obtain an extract. The extractant includes dodecanol and hexadecane, and the auxiliary agent is Span-40.

[0008] The extract was subjected to a light-removal treatment, and then pentanediamine and the extractant were separated by distillation to obtain high-purity pentanediamine.

[0009] In one or more embodiments, the volume ratio of dodecanol to hexadecane is (65~85):(15~35), the adjuvant is Span-40, and the volume ratio of the adjuvant to the extractant is (0.02~1):100.

[0010] In one or more embodiments, it further includes:

[0011] The extractant obtained by distillation is purified and regenerated for recycling.

[0012] In one or more embodiments, it further includes:

[0013] Collect the extracted aqueous phase to obtain the raffinate;

[0014] The raffinate is subjected to fractional evaporation and crystallization to obtain crystalline salt and non-crystalline concentrate;

[0015] The concentrated liquid is spray-dried to obtain organic fertilizer.

[0016] In one or more embodiments, the method for purifying and regenerating the extractant obtained by distillation separation includes:

[0017] The extractant obtained by distillation separation is purified and regenerated by using the vapor condensate generated during the fractional evaporation of the raffinate as a back-extraction wash.

[0018] To achieve the above objectives, a second aspect of this application provides an apparatus for producing bio-based pentanediamine using the production method of any of the above embodiments, comprising:

[0019] A decarbonization alkaline hydrolysis system includes a decarbonization tower and an alkaline hydrolysis reactor. The decarbonization tower is used to decarbonize the raw material to obtain a decarbonized liquid, and the alkaline hydrolysis reactor is used to perform an alkaline hydrolysis reaction on the decarbonized liquid to obtain an alkaline hydrolysis liquid.

[0020] An extraction system includes an extraction tower and a phase separation tank. The extraction tower is used to extract the alkaline hydrolysate, and the phase separation tank is used to collect the top effluent from the extraction tower for phase separation and to return the aqueous phase to the extraction tower to collect the oil phase to obtain the extract.

[0021] The purification system includes a light-light-removal tower and a product tower. The light-light-removal tower is used to remove light-light-removal from the extract, and the product tower is used to distill the light-light-removal-treated extract to separate pentanediamine and extractant, thereby obtaining a high-purity pentanediamine product.

[0022] In one embodiment, the decarbonization alkaline hydrolysis system further includes:

[0023] The first condenser is used to condense the gaseous discharge from the top of the decarbonization tower and return the condensate to the decarbonization tower.

[0024] The first heat exchanger is used to exchange heat between a portion of the alkaline hydrolysate and the raw material, and to return the heat-exchanged alkaline hydrolysate to the alkaline hydrolysis reactor.

[0025] The second heat exchanger is used to exchange heat between the raw material and the decarbonization liquid before the raw material enters the decarbonization tower.

[0026] In one embodiment, the extraction system further includes:

[0027] A washing tower is used to perform reverse extraction washing on the extractant obtained from the product tower separation.

[0028] A mixing tank is used to mix the extractant, which has been washed by reverse extraction, with the auxiliary agent and then introduce it into the extraction tower.

[0029] In one embodiment, the product column includes a crude pentanediamine outlet at the top of the column, an extractant outlet at the bottom of the column, and a product outlet at the side of the column.

[0030] The refining system also includes:

[0031] The second condenser is used to condense the crude pentanediamine collected from the top of the product column, and to partially return the condensate to the top of the product column and partially return it to the extraction column.

[0032] The third condenser is used to condense the gaseous output from the top of the light component removal tower, and to return part of the condensate to the light component removal tower, while the remaining condensed light components are discharged.

[0033] In one embodiment, the extraction system further includes:

[0034] The third heat exchanger is used to exchange heat between the crude pentanediamine refluxed from the product tower and the extract.

[0035] The fourth heat exchanger is used to exchange heat between the extract and the extractant obtained from the product tower.

[0036] The fifth heat exchanger is used to exchange heat between the alkaline hydrolysate and the extractant obtained from the product tower.

[0037] In one embodiment, the washing tower includes a detergent inlet located at the top of the tower, an extractant inlet located at the bottom of the tower, and a washing outlet located at the top of the tower;

[0038] The production apparatus also includes:

[0039] The crystallization evaporation system is used to perform staged evaporation and crystallization of the raffinate from the extraction tower to obtain crystalline salt and non-crystallized concentrate, and to send the condensate vapor generated by evaporation to the detergent inlet of the washing tower.

[0040] In one embodiment, the extraction system further includes a sixth heat exchanger for exchanging heat between the vapor condensate and the extractant obtained from the product tower.

[0041] In one embodiment, the crystallization evaporation system includes:

[0042] The three-stage evaporation unit includes an evaporator and a heater in each stage. The heater is used to circulate and heat the material in the evaporator. The evaporator is used to evaporate the material to obtain a concentrated liquid. The first and second stage evaporation units also include separators. The separators are used to separate the vapor phase output from the evaporator into gas and liquid phases, and the separated liquid phase output is returned to the evaporator. The concentrated liquid from the first stage evaporator is fed into the second stage evaporator. The vapor phase output from the first and second stage separators is fed as steam into the heater of the next stage evaporation unit.

[0043] The secondary crystallization unit includes a crystallization growth device, a crystallization centrifuge, and a crystallization mother liquor tank connected in sequence. The crystallization growth device is used to crystallize to obtain a crystallized liquid. The crystallization centrifuge is used to centrifuge the crystallized liquid. The crystallization mother liquor tank is used to collect the uncrystallized mother liquor. The primary and secondary crystallization growth devices are used to crystallize the concentrates from the secondary and tertiary evaporators, respectively. Part of the mother liquor collected from the primary and secondary crystallization mother liquor tanks is fed into the tertiary evaporator, and part is used for spray drying to prepare organic fertilizer.

[0044] Multiple preheaters are used to preheat the raffinate by using the vapor phase discharge of the three-stage evaporator, the steam condensate of the second-stage heater and the third-stage heater as heat sources, and to input the preheated raffinate into the first-stage evaporator and the cooled steam condensate into the washing tower.

[0045] In one embodiment, the operating pressure of the first-stage evaporator is 1~1.2 bara and the operating temperature is 100~105 ℃, the operating pressure of the second-stage evaporator is 0.4~0.6 bara and the operating temperature is 75~88 ℃, and the operating pressure of the third-stage evaporator is 0.12~0.2 bara and the operating temperature is 49~60 ℃.

[0046] In one embodiment, the operating pressure of the primary crystallizer is 1.0~1.6 bara and the operating temperature is 75~88 ℃; the operating pressure of the primary crystallizer mother liquor tank is 1.0~1.2 bara and the operating temperature is 75~88 ℃; the operating pressure of the secondary crystallizer is 1.0~1.6 bara and the operating temperature is 49~60 ℃; the operating pressure of the secondary crystallizer mother liquor tank is 1.0~1.2 bara and the operating temperature is 49~60 ℃.

[0047] In one embodiment, the decarbonization tower operates at a pressure of 0.2 to 0.6 bara and has a top temperature of 50 to 85°C.

[0048] In one embodiment, the operating pressure of the extraction column is 3.0~5.0 bara, and the top temperature is 45~75°C.

[0049] In one embodiment, the operating pressure of the light-weight removal tower is 0.1~0.4 bara, and the top temperature is 50~80°C.

[0050] In one embodiment, the product tower operates at a pressure of 0.05 to 0.2 bara and has a top temperature of 85 to 130°C.

[0051] In one embodiment, the alkaline hydrolysis reactor includes an alkaline hydrolysis jet reactor and an alkaline hydrolysis reaction tower connected in series. The alkaline hydrolysis jet reactor operates at a pressure of 3.0~5.0 bara and a reaction temperature of 45~75 ℃. The alkaline hydrolysis reaction tower operates at a pressure of 1.0~1.4 bara and a top temperature of 45~75 ℃.

[0052] In one embodiment, the decarbonization tower is internally arranged with one or more tray packing elements arranged vertically at intervals, the tray packing elements comprising:

[0053] Multiple plates are arranged sequentially at intervals, with at least a portion of the plates extending vertically to form a vertical channel between adjacent plates;

[0054] The packing material is arranged within the vertical channel;

[0055] A tray mass transfer element is arranged above the vertical channel, and the tray mass transfer element is used to guide the material into the vertical channel.

[0056] The advantages of this application, which differ from existing technologies, are:

[0057] This application enables the continuous industrial production of pentanediamine. By extracting with multi-component extractants and auxiliaries, pentanediamine can be transferred to the oil phase. The emulsion entrainment in the oil phase is controlled to prevent salt from entering the oil phase, thus avoiding salt precipitation during the subsequent refining process and preventing equipment and pipeline blockage. High-purity pentanediamine products are obtained by distilling the oil phase, avoiding the huge energy consumption of the traditional water distillation and concentration process.

[0058] This application uses the condensate from the crystallization evaporation process to backwash the extractant, which enables continuous regeneration of the extractant without additional energy consumption, achieving low-cost regeneration of the extractant, ensuring extractant performance, and thus realizing stable and continuous pentanediamine extraction.

[0059] This application utilizes the internal structural design of the decarbonization tower. After the fermentation broth is dispersed through the tray mass transfer element, it enters the vertical channel below through the opening. When the fermentation broth is vertically conducted in the packing, it further stretches the membrane and destroys the cell residue structure, thereby removing carbon dioxide to a greater extent and reducing the alkali consumption in the alkaline hydrolysis process. Furthermore, the design of the tray mass transfer element and the vertical channel can effectively avoid clogging problems.

[0060] This application significantly improves the yield and purity of pentanediamine by collecting crude pentanediamine from the top of the product tower and partially refluxing it to the extraction tower, while collecting high-purity pentanediamine product from the side stream, thus achieving the required purity for polymerization grade.

[0061] This application utilizes a crystallization evaporation system to convert raffinate into inorganic salts through staged crystallization, effectively improving the purity of the crystallized salts, precisely controlling the crystal particle size, and obtaining high-purity inorganic salt crystals. Simultaneously, the uncrystallized organic protein concentrate mother liquor is dried by spraying to prepare organic fertilizer. At the same time, the evaporation condensate is used to regenerate the extractant, and finally, all by-products are fully utilized, solving the waste residue problem and minimizing production costs. Attached Figure Description

[0062] To more clearly illustrate the technical solutions in the embodiments of this application or the prior art, the drawings used in the description of the embodiments or the prior art will be briefly introduced below. Obviously, the drawings described below are only some embodiments recorded in this application. For those skilled in the art, other drawings can be obtained based on these drawings without creative effort.

[0063] Figure 1 This is a schematic diagram of one embodiment of the pentamethylenediamine production apparatus of this application;

[0064] Figure 2 This is a schematic diagram of one embodiment of the decarbonization alkaline hydrolysis system of this application;

[0065] Figure 3 This is a schematic diagram of one embodiment of the decarbonization tower of this application;

[0066] Figure 4 This is a schematic diagram of one embodiment of the extraction system of this application;

[0067] Figure 5 This is a schematic diagram of one embodiment of the refining system of this application;

[0068] Figure 6 This is a schematic diagram of one embodiment of the crystallization evaporation system of this application;

[0069] Figure 7 This is a flowchart illustrating one embodiment of the method for producing bio-pentanediamine according to this application.

[0070] Explanation of key figure labels:

[0071] Decarbonization alkaline hydrolysis system 100; decarbonization tower 101; tower tray packing element 1011; plate 1012; vertical channel 1013; packing 1014; tower tray mass transfer element 1015; alkaline hydrolysis reactor 102; alkaline hydrolysis jet reactor 1021; alkaline hydrolysis reaction tower 1022; first condenser 103; first heat exchanger 104; second heat exchanger 105;

[0072] Extraction system 200; extraction tower 201; phase separation tank 202; washing tower 203; mixing tank 204; third heat exchanger 205; fourth heat exchanger 206; fifth heat exchanger 207; sixth heat exchanger 208;

[0073] Refining system 300; Light weight removal tower 301; Product tower 302; Second condenser 303; Reflux tank 304; Third condenser 305;

[0074] Crystallization evaporation system 400; Evaporation unit 401; Primary evaporator 4011; Heater 4012; Primary separator 4013; Secondary evaporator 4014; Secondary separator 4015; Tertiary evaporator 4016; Crystallization unit 402; Primary crystallization growth unit 4021; Crystallization centrifuge 4022; Primary crystallization mother liquor tank 4023; Secondary crystallization growth unit 4024; Secondary crystallization mother liquor tank 4025; Preheater 403. Detailed Implementation

[0075] To enable those skilled in the art to better understand the technical solutions in this disclosure, the technical solutions in the embodiments of this disclosure will be clearly and completely described below with reference to the accompanying drawings. Obviously, the described embodiments are only some embodiments of this disclosure, and not all embodiments. Based on the embodiments in this disclosure, all other embodiments obtained by those skilled in the art without creative effort should fall within the scope of protection of this disclosure.

[0076] Pentylenediamine (1,5-diaminopentane) has amino groups at both ends and can be used as a monomer for polyamides and polyurethanes. It is an important bio-based compound and plays a vital role in replacing petrochemical raw materials and promoting the green transformation of chemical industry.

[0077] The condensation polymerization of pentanediamine and adipic acid can produce polyamide 56 (nylon 56). Nylon 56's performance is comparable to petrochemical-based nylon 66, with superior temperature resistance and moisture absorption, making it the best bio-based material to replace nylon 66 and breaking the limitations of the traditional petrochemical route. The core monomer of traditional high-end nylon 66, hexanediamine, is mainly produced globally by companies like Nvidia and Ascend through a petrochemical route (adiponitrile hydrogenation). my country's self-sufficiency rate is less than 30%, heavily reliant on imports. The biosynthesis route of pentanediamine utilizes renewable resources (food, biomass), does not rely on petroleum, and avoids the highly polluting and high-barrier technologies of adiponitrile synthesis, providing a "self-reliant and controllable" alternative for my country's nylon industry and reducing dependence on the international petrochemical supply chain. Furthermore, pentanediamine, as a key intermediate in the synthesis of isocyanates (such as 1,5-pentanediisocyanate, PDI), can replace the highly toxic phosgene process in the production of polyurethane, used in high-end coatings and adhesives (formaldehyde-free, with outstanding environmental friendliness).

[0078] However, the biological production of pentanediamine is currently difficult to apply on a large scale industrially. The main reasons are as follows: the pentanediamine fermentation broth contains a large amount of salt, bacterial cells, cell debris, residual sugar, pigments, proteins, peptides, nucleic acids, cell metabolites and other substances, making the system complex and difficult to separate; since the pentanediamine content in the fermentation broth is only 3-10%, conventional separation and purification methods require the evaporation of a large amount of water, resulting in extremely high energy consumption and production costs; at the same time, residual sugars and proteins in the fermentation broth will denature and coke when heated, and inorganic salts will precipitate due to the evaporation of water, which can easily cause equipment and pipeline blockages, affecting the purification and production of pentanediamine.

[0079] To address the aforementioned issues, the applicant has developed a novel pentamethylenediamine production apparatus. This apparatus decarbonizes and alkali-hydrolyzes the fermentation broth, then transfers pentamethylenediamine to the oil phase via extraction, while preventing salt from entering the oil phase. High-purity pentamethylenediamine is then prepared by distillation along with the oil phase, effectively avoiding the problems of conventional separation and purification methods, such as the need to evaporate large amounts of water, high energy consumption, and the potential for inorganic salt precipitation and protein denaturation and coking that can cause equipment and pipeline blockages.

[0080] Specifically, please refer to Figure 1 , Figure 1 This is a schematic diagram of one embodiment of the pentamethylenediamine production apparatus of this application.

[0081] like Figure 1 As shown, the production apparatus includes a decarbonization alkaline hydrolysis system 100, an extraction system 200, and a purification system 300.

[0082] The decarbonization alkaline hydrolysis system 100 includes a decarbonization tower 101 and an alkaline hydrolysis reactor 102. The decarbonization tower 101 is used to decarbonize the raw material to obtain a decarbonized liquid, and the alkaline hydrolysis reactor 102 is used to perform an alkaline hydrolysis reaction on the decarbonized liquid to obtain an alkaline hydrolysis liquid.

[0083] The extraction system 200 includes an extraction tower 201 and a phase separation tank 202. The extraction tower 201 is used to extract the alkaline hydrolysate, and the phase separation tank 202 is used to collect the top discharge of the extraction tower 201 for phase separation and to return the aqueous phase to the extraction tower 201 to collect the oil phase to obtain the extract.

[0084] The refining system 300 includes a light-light removal tower 301 and a product tower 302. The light-light removal tower 301 is used to remove light-light substances from the extract, and the product tower 302 is used to distill the extract after light-light removal to separate pentanediamine and extractant, thereby obtaining a high-purity pentanediamine product.

[0085] Based on the above-mentioned production device, carbon dioxide is first removed from the fermentation broth through decarbonization tower 101 to avoid additional alkali consumption; then, the decarbonization broth is subjected to alkaline hydrolysis treatment through alkaline hydrolysis reactor 102 to convert pentanediamine salt into free pentanediamine, resulting in an alkaline hydrolysis broth containing free pentanediamine.

[0086] By extracting the alkaline hydrolysate, pentanediamine is transferred to the oil phase, and further distillation yields a high-purity pentanediamine product. Since the pentanediamine is distilled along with the oil phase, the enormous energy consumption of the traditional water distillation process is effectively avoided, as well as the blockage of pipelines and equipment caused by the evaporation and precipitation of inorganic salts in the oil phase.

[0087] The system architecture is described in detail below. Please refer to [link / reference]. Figure 1 and Figure 2 , Figure 2 This is a schematic diagram of one embodiment of the decarbonization alkaline hydrolysis system of this application.

[0088] like Figure 1 and Figure 2 As shown, in order to ensure that the fermentation broth is fully decarbonized, the decarbonization alkaline hydrolysis system 100 in this embodiment also includes a first condenser 103. The first condenser 103 is connected to the top gas phase discharge of the decarbonization tower 101 and is used to condense the top gas phase discharge of the decarbonization tower 101 and return the condensate to the decarbonization tower 101.

[0089] In order to make full use of the thermal energy of the alkaline hydrolysis solution and the decarbonization solution and reduce energy consumption, the decarbonization alkaline hydrolysis system 100 also includes a first heat exchanger 104 and a second heat exchanger 105.

[0090] The outlet of the alkaline hydrolysis reactor 102 is connected to both the hot end inlet of the first heat exchanger 104 and the extraction system 200. The hot end outlet of the first heat exchanger 104 is connected to the inlet of the alkaline hydrolysis reactor 102. The cold end inlet of the first heat exchanger 104 is connected to the raw material source, and the cold end outlet is connected to the cold end inlet of the second heat exchanger 105.

[0091] Furthermore, the cold end outlet of the second heat exchanger 105 is connected to the decarbonization tower 101, the bottom outlet of the decarbonization tower 101 is connected to the hot end inlet of the second heat exchanger 105, and the hot end outlet of the second heat exchanger 105 is connected to the inlet of the alkaline hydrolysis reactor 102.

[0092] Based on the above design, the first heat exchanger 104 exchanges heat between part of the alkaline hydrolysis solution and the raw material, and returns the alkaline hydrolysis solution after heat exchange to the alkaline hydrolysis reactor 102; the second heat exchanger 105 exchanges heat between the raw material and the decarbonization liquid before the raw material enters the decarbonization tower 101, thereby making full use of the heat of the alkaline hydrolysis solution and the decarbonization liquid, and ensuring the complete reaction of the alkaline hydrolysis treatment by the return of the alkaline hydrolysis solution.

[0093] Specifically, in this embodiment, the alkaline hydrolysis reactor 102 includes an alkaline hydrolysis jet reactor 1021 and an alkaline hydrolysis reaction tower 1022 connected in series. The alkaline hydrolysis jet reactor 1021 is a microreactor used to react the decarbonization liquid and liquid alkali in the early stage. It has a fast reaction rate and small reaction volume, making it suitable for the early stage high-concentration material reaction zone. The alkaline hydrolysis reaction tower 1022 is a vertical tubular reactor used in the later stage low-concentration material reaction zone to ensure that the pentanediamine conversion reaction is complete.

[0094] Because the pentanediamine fermentation broth contains a large amount of cell debris, residual sugar, protein, and other organic matter, and has a high viscosity, it is difficult to remove the large amount of carbon dioxide. To effectively remove carbon dioxide and avoid clogging of the decarbonization tower 101, please refer to [the relevant documentation / reference needed]. Figure 3 , Figure 3 This is a structural schematic diagram of one embodiment of the decarbonization tower of this application.

[0095] like Figure 3 As shown, in this embodiment, the decarbonization tower 101 is internally arranged with a plurality of tray packing elements 1011 arranged vertically at intervals. The tray packing elements 1011 include a plurality of plates 1012 arranged at intervals. The plates 1012 extend vertically to form vertical channels 1013 between adjacent plates 1012. Packing 1014 is arranged in the vertical channels 1013. A tray mass transfer element 1015 is arranged above the vertical channels 1013 to guide the material to the vertical channels 1013.

[0096] Specifically, the tray mass transfer element 1015 is provided with openings that correspond one-to-one with the vertical channels 1013. After the fermentation broth is dispersed through the tray mass transfer element 1015, it enters the vertical channel 1013 below through the openings. When the fermentation broth is vertically conducted in the packing 1014, it further stretches the membrane and destroys the cell residue structure, so that carbon dioxide is removed to a greater extent, reducing the alkali consumption in the alkaline hydrolysis process. Moreover, the design of the tray mass transfer element 1015 and the vertical channel 1013 can effectively avoid clogging problems.

[0097] In one embodiment, the operating pressure of the decarbonization tower 101 can be 0.2~0.6 bara, and the top temperature of the tower can be 50~85 ℃.

[0098] In one embodiment, the operating pressure of the alkaline hydrolysis jet reactor 1021 can be 3.0~5.0 bara, and the reaction temperature can be 45~75 ℃. The operating pressure of the alkaline hydrolysis reaction tower 1022 can be 1.0~1.4 bara, and the tower top temperature can be 45~75 ℃.

[0099] Please see Figure 1 and Figure 4 , Figure 4 This is a schematic diagram of one embodiment of the extraction system of this application.

[0100] like Figure 4 As shown, the extraction system 200 also includes a washing tower 203 and a mixing tank 204. The washing tower 203 is used to perform reverse extraction washing on the extractant obtained by distillation separation in the product tower 302. The mixing tank 204 is used to mix the extractant washed by reverse extraction with the auxiliary agent and introduce it into the extraction tower 201.

[0101] Specifically, in this embodiment, the extraction tower 201 uses an extractant and an auxiliary agent to extract the alkaline hydrolysate, so as to transfer the pentanediamine to the oil phase and prevent salt from entering the interior of the oil phase, thereby avoiding the problem of equipment and pipeline blockage caused by salt precipitation in the subsequent refining process.

[0102] The extractants include dodecyl alcohol and hexadecane, with Span-40 as the auxiliary agent. Dodecyl alcohol, as the main extractant, has hydroxyl groups that can form hydrogen bonds with the amino groups at both ends of the pentanediamine molecule, thus selectively transferring pentanediamine from the complex aqueous phase to the oil phase. Hexadecane, as a diluent and regulator, enhances the hydrophobic properties of the oil phase, effectively preventing hydrophilic impurities such as inorganic salts and water-soluble proteins from entering the oil phase. The fermentation broth has a high protein content, making it prone to emulsification during extraction, leading to difficult stratification and reduced efficiency. Span-40, as... The surfactant Span-40 adsorbs onto the interface between the oil and water phases, significantly reducing interfacial tension and making it easier for the oil phase to disperse into fine droplets. This increases the contact area between the two phases and accelerates the mass transfer rate of pentanediamine from the aqueous phase to the oil phase. Meanwhile, the proteins, bacteria, and other substances in the fermentation broth are themselves natural surfactants that can stabilize the oil-water emulsion and make it difficult to separate. The addition of Span-40 can replace or interfere with the arrangement of these impurities at the interface, disrupting the stable emulsion film structure, thereby enabling the oil and water phases to be separated quickly and clearly.

[0103] In one embodiment, the volume ratio of dodecanol to hexadecane is (65~85):(15~35), the auxiliary is Span-40, and the volume ratio of the auxiliary to the extractant is (0.02~1):100.

[0104] In this embodiment, the extractant obtained by distillation separation is regenerated by reverse extraction washing in washing tower 203, and then auxiliary agents are added to achieve continuous purification of the extractant, with stable extraction performance, thereby realizing continuous production and effectively reducing costs and energy consumption.

[0105] Furthermore, such as Figure 1 As shown, the production apparatus in this embodiment also includes a crystallization evaporation system 400. The aqueous phase (raffinate) obtained by the extraction tower 201 during the extraction process contains a large amount of inorganic salts, proteins, etc. The crystallization evaporation system 400 is used to perform fractional evaporation and crystallization on the raffinate of the extraction tower 201 to obtain crystalline salt and non-crystallized concentrate. The steam condensate generated by evaporation is sent to the washing tower 203 to use the steam condensate generated by the crystallization evaporation system 400 to wash the extractant. The extractant regeneration process does not require additional heat consumption, further reducing the system energy consumption.

[0106] Specifically, in this embodiment, the washing tower 203 includes a detergent inlet located at the top of the tower, an extractant inlet located at the bottom of the tower, and a washing outlet located at the top of the tower. The crystallization evaporation system 400 is connected to the detergent inlet, the extractant inlet is connected to the product tower 302, and the washing outlet is connected to the mixing tank 204.

[0107] Inside the washing tower 203, the extractant flows from bottom to top, while the steam condensate from the crystallization evaporation system 400 flows from top to bottom, achieving backwashing of the extractant. The washed extractant is discharged from the washing outlet, realizing the regeneration and recycling of the extractant without additional energy consumption.

[0108] In one embodiment, the operating pressure of the extraction column 201 can be 3.0~5.0 bara, and the top temperature of the column can be 45~75 °C.

[0109] Please see Figure 1 and Figure 5 , Figure 5 This is a schematic diagram of one embodiment of the refining system of this application. For example... Figure 1 and Figure 5 As shown, product column 302 includes a crude pentanediamine outlet at the top of the column, an extractant outlet at the bottom of the column, and a product outlet at the side of the column.

[0110] The refining system 300 also includes a second condenser 303 and a reflux tank 304. The second condenser 303 is used to condense the crude pentanediamine collected from the top of the product column 302 and store the condensate in the reflux tank 304. Part of the condensate in the reflux tank 304 is returned to the top of the product column 302 and part is returned to the extraction column 201.

[0111] Based on the above design, the crude pentanediamine collected from the top of product tower 302 is returned to the bottom of extraction tower 201 for recovery and purification, while the high-purity pentanediamine product is collected from the side stream, which can significantly improve the purity of the product and meet the purity requirements of polymerization grade.

[0112] Meanwhile, some impurities in the fermentation broth, especially some unknown impurities, can be enriched by distillation in product tower 302 at low concentrations. After being enriched to a certain concentration, they are difficult to completely separate from pentanediamine. By collecting crude pentanediamine and refluxing it to extraction tower 201, the difference in the distribution of these components and pentanediamine in the water / oil phase can be utilized to dissolve most of the impurities into the aqueous phase through extraction again, and the pentanediamine returns to the oil phase, thus completing the recovery of pentanediamine and significantly improving the yield and purity of pentanediamine.

[0113] Furthermore, the refining system 300 also includes a third condenser 305, which is used to condense the gaseous output from the top of the light component removal tower 301 and return part of the condensate to the light component removal tower 301, while discharging part of the condensed light component, thereby improving the light component removal effect.

[0114] In one embodiment, the operating pressure of the light component removal tower 301 can be 0.1~0.4 bara, and the top temperature can be 50~80 ℃; the operating pressure of the product tower 302 can be 0.05~0.2 bara, and the top temperature can be 85~130 ℃.

[0115] To fully utilize the heat from the crude pentanediamine refluxed to extraction tower 201 and the heat from the extractant separated in product tower 302, such as Figure 4 As shown, the extraction system 200 also includes a third heat exchanger 205, a fourth heat exchanger 206, and a fifth heat exchanger 207.

[0116] The third heat exchanger 205 is used to exchange heat between the crude pentanediamine refluxed from the product tower 302 and the extractant; the fourth heat exchanger 206 is used to exchange heat between the extract and the extractant obtained from the product tower 302; and the fifth heat exchanger 207 is used to exchange heat between the alkaline hydrolysis solution and the extractant obtained from the product tower 302.

[0117] Specifically, in this embodiment, the hot end inlet of the fourth heat exchanger 206 is connected to the extractant outlet of the product tower 302, the hot end outlet is connected to the hot end inlet of the fifth heat exchanger 207, the cold end inlet is connected to the oil phase outlet of the phase separation tank 202, and the cold end outlet is connected to the light phase removal tower 301.

[0118] The hot end outlet of the fifth heat exchanger 207 is connected to the cold end inlet of the third heat exchanger 205, the cold end inlet is connected to the bottom outlet of the alkaline reaction tower 1022, and the cold end outlet is connected to the extraction tower 201.

[0119] The cold end outlet of the third heat exchanger 205 is connected to the extractant inlet of the scrubbing tower 203, the hot end inlet is connected to the reflux tank 304, and the hot end outlet is connected to the extraction tower 201.

[0120] Based on the above design, the extractant separated in product tower 302 first exchanges heat with the extractant in the fourth heat exchanger 206. The extracted liquid after heat exchange enters the light component removal tower 301. The extractant continues to exchange heat with the alkaline hydrolysate in the fifth heat exchanger 207. The alkaline hydrolysate after heat exchange enters the extraction tower 201. The extractant continues to exchange heat with the refluxed crude pentanediamine in the third heat exchanger 205. The crude pentanediamine after heat exchange enters the extraction tower 201. The extractant then enters the washing tower 203 for washing and regeneration, significantly improving heat utilization and reducing energy consumption.

[0121] Furthermore, in order to make full use of the heat of the steam condensate used for washing and regenerating the extractant, the extraction system 200 in this embodiment also includes a sixth heat exchanger 208, which is used to exchange heat between the steam condensate and the extractant obtained from the product tower 302.

[0122] Specifically, the sixth heat exchanger 208 is arranged between the fourth heat exchanger 206 and the fifth heat exchanger 207. Its cold end inlet is connected to the hot end outlet of the fourth heat exchanger 206, its cold end outlet is connected to the hot end inlet of the fifth heat exchanger 207, its hot end inlet is connected to the crystallization evaporation system 400, and its hot end outlet is connected to the detergent inlet of the scrubbing tower 203, thereby realizing heat exchange between the steam condensate and the extractant.

[0123] Please see Figure 6 , Figure 6 This is a schematic diagram of one embodiment of the crystallization evaporation system of this application. Figure 6 As shown, the crystallization evaporation system 400 includes a three-stage evaporation unit 401, a two-stage crystallization unit 402, and three preheaters 403.

[0124] Each stage of the three-stage evaporation unit 401 includes an evaporator and a heater 4012. The heater 4012 is used to circulate and heat the material in the evaporator. The evaporator is used to evaporate the material to obtain a concentrated liquid. The first and second stage evaporation units 401 also include separators. The separators are used to separate the vapor phase output from the evaporator into gas and liquid phases, and return the separated liquid phase output to the evaporator. The concentrated liquid from the first stage evaporator 4011 is input into the second stage evaporator 4014. The vapor phase output from the first stage separator 4013 and the second stage separator 4015 is used as steam and input into the heater 4012 of the next stage evaporation unit 401.

[0125] Each stage of the secondary crystallization unit 402 includes a crystallization growth device, a crystallization centrifuge 4022, and a crystallization mother liquor tank connected in sequence. The crystallization growth device is used to crystallize to obtain crystallization liquid. The crystallization centrifuge 4022 is used to centrifuge the crystallization liquid to collect inorganic salt crystals. The crystallization mother liquor tank is used to collect the uncrystallized mother liquor. The primary crystallization growth device 4021 and the secondary crystallization growth device 4024 are used to crystallize the concentrates of the secondary evaporator 4014 and the tertiary evaporator 4016, respectively. The mother liquor collected by the primary crystallization mother liquor tank 4023 and the secondary crystallization mother liquor tank 4025 is partly input into the tertiary evaporator 4016 and partly used for spray drying to prepare organic fertilizer.

[0126] The three preheaters 403 are respectively used to preheat the raffinate by using the vapor phase discharge of the three-stage evaporator 4016, the steam condensate of the two-stage heater 4012 and the three-stage heater 4012 as heat sources, and input the preheated raffinate into the first-stage evaporator 4011, and input the cooled steam condensate into the washing tower 203.

[0127] In one embodiment, the operating pressure of the first-stage evaporator 4011 is 1~1.2 bara and the operating temperature is 100~105 °C; the operating pressure of the second-stage evaporator 4014 is 0.4~0.6 bara and the operating temperature is 75~88 °C; and the operating pressure of the third-stage evaporator 4016 is 0.12~0.2 bara and the operating temperature is 49~60 °C.

[0128] In one embodiment, the operating pressure of the primary crystallizer 4021 is 1.0~1.6 bara and the operating temperature is 75~88 ℃; the operating pressure of the primary crystallizer mother liquor tank 4023 is 1.0~1.2 bara and the operating temperature is 75~88 ℃; the operating pressure of the secondary crystallizer 4024 is 1.0~1.6 bara and the operating temperature is 49~60 ℃; and the operating pressure of the secondary crystallizer mother liquor tank 4025 is 1.0~1.2 bara and the operating temperature is 49~60 ℃.

[0129] Based on the above design, the purity of crystalline salt can be effectively improved by graded crystallization, the crystal particle size can be precisely controlled, and high-purity inorganic salt crystals can be obtained. At the same time, the uncrystallized organic protein concentrated mother liquor can be dried by spraying to prepare organic fertilizer. Finally, all by-products are fully utilized, solving the waste residue problem and minimizing production costs.

[0130] The bio-based pentanediamine production apparatus based on the above embodiments can realize the continuous industrial production of pentanediamine. By extracting with extractants and additives, pentanediamine can be transferred to the oil phase, and the emulsion entrainment in the oil phase can be controlled to prevent salt from entering the oil phase, thus avoiding the blockage of equipment and pipelines caused by salt precipitation in the subsequent refining process. High-purity pentanediamine products are obtained by distilling the oil phase, avoiding the huge energy consumption of the traditional water distillation and concentration process.

[0131] By backwashing the extractant with the condensate from the crystallization evaporation process, continuous regeneration of the extractant can be achieved without additional energy consumption, thus enabling low-cost regeneration of the extractant, ensuring its performance, and achieving stable and continuous pentanediamine extraction.

[0132] Through the internal structural design of the decarbonization tower 101, the fermentation broth is dispersed through the tray mass transfer element 1015 and then enters the lower vertical channel 1013 through the opening. When the fermentation broth is vertically conducted in the packing 1014, it further stretches the membrane and destroys the cell residue structure, so as to remove carbon dioxide to a greater extent, reduce the alkali consumption of the alkaline hydrolysis process, and the design of the tray mass transfer element 1015 and the vertical channel 1013 can effectively avoid the clogging problem.

[0133] By collecting crude pentanediamine from the top of product tower 302 and partially refluxing it to extraction tower 201, and collecting high-purity pentanediamine product from the side stream, the yield and purity of pentanediamine can be significantly improved, meeting the purity requirements for polymerization grade.

[0134] The 400 crystallization evaporation system can convert the raffinate into inorganic salts, effectively improving the purity of the crystallized salts and precisely controlling the crystal particle size to obtain high-purity inorganic salt crystals. At the same time, the uncrystallized organic protein concentrate mother liquor is dried by spraying to prepare organic fertilizer. Meanwhile, the evaporation condensate is used to regenerate the extractant. In the end, all by-products are fully utilized, solving the waste residue problem and minimizing production costs.

[0135] This application also provides a method for producing bio-based pentanediamine; please refer to [link to method]. Figure 7 , Figure 7 This is a flowchart illustrating one embodiment of the method for producing bio-pentanediamine according to this application.

[0136] like Figure 7 As shown, the method includes:

[0137] S100. Using pentanediamine fermentation broth and / or the treated broth of pentanediamine fermentation broth as raw materials, the raw materials are decarbonized to obtain a decarbonized liquid, and then the decarbonized liquid is subjected to alkaline hydrolysis to obtain an alkaline hydrolysate.

[0138] S200. Extract the alkaline hydrolysate using an extractant and an auxiliary agent, collect the oil phase, and obtain the extract.

[0139] S300: The extract is subjected to light-removal treatment, and then pentanediamine and extractant are separated by distillation to obtain high-purity pentanediamine.

[0140] S400. Collect the aqueous phase of the extraction to obtain the raffinate. Perform fractional evaporation and crystallization on the raffinate to obtain crystalline salt and non-crystalline concentrate.

[0141] S500: The extractant obtained from distillation separation is purified and regenerated for recycling.

[0142] In one embodiment, the condensate vapor generated during the evaporation process of the raffinate can be used to perform reverse extraction washing on the extractant obtained from the distillation separation, so as to achieve purification and regeneration of the extractant without additional energy consumption.

[0143] S600: The concentrate is spray-dried to obtain organic fertilizer.

[0144] The above Figures 1 to 6 The description of the apparatus for producing bio-based pentanediamine is also applicable to the method for producing bio-based pentanediamine of this application, and will not be repeated here.

[0145] The beneficial effects of the technical solution of this application will be further explained in detail below with reference to specific embodiments.

[0146] Example:

[0147] by Figure 1 The production unit shown produces pentanediamine. The raw material is fresh pentanediamine fermentation broth, which is composed of 195 g / L pentanediamine salt (sulfate salt), 225 g / L dry matter such as protein, bacterial cells, residual sugar, and cell metabolites, 8 g / L potassium chloride, and the remaining components are water and a small amount of cell metabolites. Carbon dioxide remains in the fermentation broth, and a small amount exists in the form of pentanediamine carbonate. The pH of the material is 8.

[0148] The operating parameters are as follows: the operating pressure of decarbonization tower 101 is 0.4 bara, the top temperature is 65 ℃, and the bottom temperature is 78 ℃. The operating pressure of alkaline hydrolysis jet reactor 1021 is 4.5 bara, the reaction temperature is 62 ℃, and the operating pressure of alkaline hydrolysis reaction tower 1022 is 1.2 bara, with a top temperature of 65 ℃.

[0149] The operating pressure of extraction column 201 is 1.5 bara, the top temperature is 65 ℃, and the bottom temperature is 70 ℃. The operating pressure of washing column 203 (116) is 1.8 bara, the top temperature is 70 ℃, and the bottom temperature is 70 ℃. The formulation of the extractant is as follows: dodecanol accounts for 80% by volume, hexadecane accounts for 20% by volume, and the auxiliary agent is Span-40, which is added at a ratio of 0.05% by volume of the extractant.

[0150] The operating pressure of the light component removal column 301 is 0.2 bara, the top temperature is 66 ℃, and the bottom temperature is 196 ℃; the operating pressure of the product column 302 is 0.1 bara, the top temperature is 106 ℃, and the bottom temperature is 198 ℃.

[0151] The operating pressure of the primary evaporator 4011 is 1.1 bara, and the operating temperature is 101 ℃; the operating pressure of the secondary evaporator 4014 is 0.5 bara, and the operating temperature is 83 ℃; the operating pressure of the tertiary evaporator 4016 is 0.15 bara, and the operating temperature is 56 ℃; the operating pressure of the primary crystallizer 4021 is 1.2 bara, and the operating temperature is 83 ℃; the operating pressure of the primary crystallizer mother liquor tank 4023 is 1.1 bara, and the operating temperature is 83 ℃; the operating pressure of the secondary crystallizer 4024 is 1.4 bara, and the operating temperature is 55 ℃; the operating pressure of the secondary crystallizer mother liquor tank 4025 is 1.2 bara, and the operating temperature is 55 ℃.

[0152] Testing revealed that after decarbonization treatment, the pH of the pentanediamine fermentation broth increased from 8 to 10, and the carbon dioxide removal rate reached 85%. The yield of pentanediamine reached 97%, and the purity of the high-purity pentanediamine product obtained from the side stream of product tower 302 was 99.85% (mass), with a water content of 480 ppm. The purity of the crystalline salt produced by the crystallization evaporation system 400 reached 98.7%.

[0153] It will be apparent to those skilled in the art that this disclosure is not limited to the details of the exemplary embodiments described above, and that this disclosure can be implemented in other specific forms without departing from its spirit or essential characteristics. Therefore, the embodiments should be considered in all respects as exemplary and non-limiting, and the scope of this disclosure is defined by the appended claims rather than the foregoing description. Thus, all variations falling within the meaning and scope of equivalents of the claims are intended to be included within this disclosure. No reference numerals in the claims should be construed as limiting the scope of the claims.

[0154] Furthermore, it should be understood that although this specification describes embodiments, not every embodiment contains only one independent technical solution. This narrative style is merely for clarity. Those skilled in the art should consider the specification as a whole, and the technical solutions in each embodiment can also be appropriately combined to form other embodiments that can be understood by those skilled in the art.

Claims

1. A method for producing bio-based pentanediamine, characterized in that, include: Using pentanediamine fermentation broth and / or pentanediamine fermentation broth treatment liquid as raw materials, the raw materials are decarbonized to obtain decarbonized liquid, and then the decarbonized liquid is subjected to alkaline hydrolysis to obtain alkaline hydrolysate. The alkaline hydrolysate is extracted using an extractant and an auxiliary agent, and the oil phase is collected to obtain an extract. The extractant includes dodecyl alcohol and hexadecane, the auxiliary agent is Span-40, the volume ratio of dodecyl alcohol to hexadecane is (65~85):(15~35), and the volume ratio of the auxiliary agent to the extractant is (0.02~1):

100. The extract was subjected to a light-removal treatment, and then pentanediamine and the extractant were separated by distillation to obtain pentanediamine; Collect the extracted aqueous phase to obtain the raffinate; The raffinate is subjected to fractional evaporation and crystallization to obtain crystalline salt and non-crystalline concentrate; The concentrated liquid is spray-dried to obtain organic fertilizer; The extractant obtained by distillation separation is washed by reverse extraction using the vapor condensate generated during the fractional evaporation of the raffinate, so as to achieve purification, regeneration and recycling of the extractant; The production equipment used in the production method includes: A decarbonization alkaline hydrolysis system includes a decarbonization tower and an alkaline hydrolysis reactor. The decarbonization tower is used to decarbonize the raw material to obtain a decarbonized liquid, and the alkaline hydrolysis reactor is used to perform an alkaline hydrolysis reaction on the decarbonized liquid to obtain an alkaline hydrolysis liquid. An extraction system includes an extraction tower and a phase separation tank. The extraction tower is used to extract the alkaline hydrolysate, and the phase separation tank is used to collect the top effluent from the extraction tower for phase separation and to return the aqueous phase to the extraction tower to collect the oil phase to obtain the extract. A refining system includes a light-light removal column and a product column. The light-light removal column is used to remove light-light substances from the extract, and the product column is used to distill the light-light removal extract to separate pentanediamine and extractant to obtain pentanediamine product. The product column includes a crude pentanediamine outlet at the top of the column, an extractant outlet at the bottom of the column, and a product outlet at the side of the column. The refining system also includes: The second condenser is used to condense the crude pentanediamine collected from the top of the product column, and to partially return the condensate to the top of the product column and partially return it to the extraction column. The third condenser is used to condense the gaseous output from the top of the light component removal tower, and to return part of the condensate to the light component removal tower, while discharging part of the condensed light component. The decarbonization tower is internally equipped with one or more tray packing elements arranged vertically at intervals, the tray packing elements including: Multiple plates are arranged sequentially at intervals, with at least a portion of the plates extending vertically to form a vertical channel between adjacent plates; The packing material is arranged within the vertical channel; A tray mass transfer element is arranged above the vertical channel, and the tray mass transfer element is used to guide the material into the vertical channel.

2. The production method according to claim 1, characterized in that, The decarbonization alkaline hydrolysis system also includes: The first condenser is used to condense the gaseous discharge from the top of the decarbonization tower and return the condensate to the decarbonization tower. The first heat exchanger is used to exchange heat between a portion of the alkaline hydrolysate and the raw material, and to return the heat-exchanged alkaline hydrolysate to the alkaline hydrolysis reactor. The second heat exchanger is used to exchange heat between the raw material and the decarbonization liquid before the raw material enters the decarbonization tower.

3. The production method according to claim 1, characterized in that, The extraction system also includes: A washing tower is used to perform reverse extraction washing on the extractant obtained from the product tower separation. A mixing tank is used to mix the extractant, which has been washed by reverse extraction, with the auxiliary agent and then introduce it into the extraction tower.

4. The production method according to claim 3, characterized in that, The extraction system also includes: The third heat exchanger is used to exchange heat between the crude pentanediamine refluxed from the product tower and the extractant. The fourth heat exchanger is used to exchange heat between the extract and the extractant obtained from the product tower. The fifth heat exchanger is used to exchange heat between the alkaline hydrolysate and the extractant obtained from the product tower.

5. The production method according to claim 3, characterized in that, The washing tower includes a detergent inlet located at the top of the tower, an extractant inlet located at the bottom of the tower, and a washing outlet located at the top of the tower. The production apparatus also includes: The crystallization evaporation system is used to perform staged evaporation and crystallization of the raffinate from the extraction tower to obtain crystalline salt and non-crystallized concentrate, and to send the condensate vapor generated by evaporation to the detergent inlet of the washing tower.

6. The production method according to claim 5, characterized in that, The extraction system also includes a sixth heat exchanger, which is used to exchange heat between the steam condensate and the extractant obtained from the product tower.

7. The production method according to claim 5, characterized in that, The crystallization evaporation system includes: The three-stage evaporation unit includes an evaporator and a heater in each stage. The heater is used to circulate and heat the material in the evaporator. The evaporator is used to evaporate the material to obtain a concentrated liquid. The first and second stage evaporation units also include separators. The separators are used to separate the vapor phase output from the evaporator into gas and liquid phases, and the separated liquid phase output is returned to the evaporator. The concentrated liquid from the first stage evaporator is fed into the second stage evaporator. The vapor phase output from the first and second stage separators is fed as steam into the heater of the next stage evaporation unit. The secondary crystallization unit includes a crystallization growth device, a crystallization centrifuge, and a crystallization mother liquor tank connected in sequence. The crystallization growth device is used to crystallize to obtain a crystallized liquid. The crystallization centrifuge is used to centrifuge the crystallized liquid. The crystallization mother liquor tank is used to collect the uncrystallized mother liquor. The primary and secondary crystallization growth devices are used to crystallize the concentrates from the secondary and tertiary evaporators, respectively. Part of the mother liquor collected from the primary and secondary crystallization mother liquor tanks is fed into the tertiary evaporator, and part is used for spray drying to prepare organic fertilizer. Multiple preheaters are used to preheat the raffinate by using the vapor phase discharge of the three-stage evaporator, the steam condensate of the second-stage heater and the third-stage heater as heat sources, and to input the preheated raffinate into the first-stage evaporator and the cooled steam condensate into the washing tower.

8. The production method according to claim 7, characterized in that, The operating pressure of the first-stage evaporator is 1~1.2 bara, and the operating temperature is 100~105℃; the operating pressure of the second-stage evaporator is 0.4~0.6 bara, and the operating temperature is 75~88℃; the operating pressure of the third-stage evaporator is 0.12~0.2 bara, and the operating temperature is 49~60℃; and / or, The operating pressure of the primary crystallizer is 1.0~1.6 bara, and the operating temperature is 75~88 ℃. The operating pressure of the primary crystallizer mother liquor tank is 1.0~1.2 bara, and the operating temperature is 75~88 ℃. The operating pressure of the secondary crystallizer is 1.0~1.6 bara, and the operating temperature is 49~60 ℃. The operating pressure of the secondary crystallizer mother liquor tank is 1.0~1.2 bara, and the operating temperature is 49~60 ℃.

9. The production method according to claim 3, characterized in that, The decarbonization tower operates at a pressure of 0.2~0.6 bara and a top temperature of 50~85℃; and / or, The operating pressure of the extraction column is 1.0~4.0 bara, and the top temperature is 45~75 ℃; and / or, The operating pressure of the scrubbing tower is 1.0~4.0 bara, and the top temperature is 50~80 ℃; and / or The operating pressure of the light-weight removal tower is 0.1~0.4 bara, and the top temperature is 50~80 ℃; and / or, The product column operates at a pressure of 0.05~0.2 bara and a top temperature of 85~130 ℃; and / or, The alkaline hydrolysis reactor comprises an alkaline hydrolysis jet reactor and an alkaline hydrolysis reaction tower connected in series. The alkaline hydrolysis jet reactor operates at a pressure of 3.0~5.0 bara and a reaction temperature of 45~75 ℃. The alkaline hydrolysis reaction tower operates at a pressure of 1.0~1.4 bara and a top temperature of 45~75 ℃.

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