WO2021197058A1 - 通过微反应器合成氧杂环丁烷化合物的方法 - Google Patents

通过微反应器合成氧杂环丁烷化合物的方法 Download PDF

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WO2021197058A1
WO2021197058A1 PCT/CN2021/081186 CN2021081186W WO2021197058A1 WO 2021197058 A1 WO2021197058 A1 WO 2021197058A1 CN 2021081186 W CN2021081186 W CN 2021081186W WO 2021197058 A1 WO2021197058 A1 WO 2021197058A1
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microreactor
reaction
synthesizing
carbonate
ppm
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PCT/CN2021/081186
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English (en)
French (fr)
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钱晓春
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常州强力先端电子材料有限公司
常州强力电子新材料股份有限公司
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Priority claimed from CN202110184193.0A external-priority patent/CN113493426A/zh
Application filed by 常州强力先端电子材料有限公司, 常州强力电子新材料股份有限公司 filed Critical 常州强力先端电子材料有限公司
Priority to KR1020227038089A priority Critical patent/KR20220163415A/ko
Priority to JP2022557799A priority patent/JP7438391B2/ja
Priority to US17/907,244 priority patent/US20230150961A1/en
Priority to EP21782226.1A priority patent/EP4129991A4/en
Publication of WO2021197058A1 publication Critical patent/WO2021197058A1/zh

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    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07D—HETEROCYCLIC COMPOUNDS
    • C07D305/00—Heterocyclic compounds containing four-membered rings having one oxygen atom as the only ring hetero atoms
    • C07D305/02—Heterocyclic compounds containing four-membered rings having one oxygen atom as the only ring hetero atoms not condensed with other rings
    • C07D305/04—Heterocyclic compounds containing four-membered rings having one oxygen atom as the only ring hetero atoms not condensed with other rings having no double bonds between ring members or between ring members and non-ring members
    • C07D305/06—Heterocyclic compounds containing four-membered rings having one oxygen atom as the only ring hetero atoms not condensed with other rings having no double bonds between ring members or between ring members and non-ring members with only hydrogen atoms, hydrocarbon or substituted hydrocarbon radicals, directly attached to the ring atoms
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J19/00—Chemical, physical or physico-chemical processes in general; Their relevant apparatus
    • B01J19/0093—Microreactors, e.g. miniaturised or microfabricated reactors
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J23/00—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00
    • B01J23/02—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00 of the alkali- or alkaline earth metals or beryllium
    • B01J23/04—Alkali metals
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J2219/00—Chemical, physical or physico-chemical processes in general; Their relevant apparatus
    • B01J2219/00781—Aspects relating to microreactors
    • B01J2219/00851—Additional features
    • B01J2219/00858—Aspects relating to the size of the reactor
    • B01J2219/0086—Dimensions of the flow channels
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J2219/00—Chemical, physical or physico-chemical processes in general; Their relevant apparatus
    • B01J2219/00781—Aspects relating to microreactors
    • B01J2219/00891—Feeding or evacuation

Definitions

  • the invention relates to the field of organic synthesis, and in particular to a method for synthesizing oxetane compounds through a microreactor.
  • 3-Ethyl-3-hydroxymethyloxetane and bis[1-ethyl(3-oxetanyl)methyl]ether are currently the most used monomers in photocurable cationic systems and are widely used In the fields of light-curable coatings, inks, adhesives, etc., their structural formulas are as follows:
  • Cyclocarbonate cracking method is widely used in industry to prepare oxetane products.
  • the reaction process for producing 3-ethyl-3-hydroxymethyloxetane by the cyclic carbonate cracking method is as follows:
  • R is an alkyl group, generally a methyl group or an ethyl group.
  • the process flow is: the transesterification reaction between trimethylolpropane and carbonate in a rectifying kettle, the temperature is 80°C ⁇ 120°C, the by-product alcohol is continuously fractionated during the reaction process, and the excess carbonate and ester are distilled off after the reaction is completed.
  • the exchange process requires 10h-12h, and then enters the cracking process, cracking at a temperature of 160°C ⁇ 200°C for 12h ⁇ 15h to remove carbon dioxide, and finally rectifying under negative pressure to obtain the finished product.
  • the entire process needs 30h-40h to complete, and the production efficiency is low.
  • the cracking reaction is carried out at high temperature for a long time, which will result in the production of high-boiling by-products, so the yield of the finished product is low (65%-75%). After the rectification is completed, a large amount of distillation residues are produced, which can only be treated as solid waste.
  • the main purpose of the present invention is to provide a method for synthesizing oxetane compounds through a microreactor to solve the problems of low yield and long reaction time in the existing synthetic methods of oxetane compounds. On the basis, it also provides a method for adjusting the product distribution by controlling the relevant parameters in the reaction process to realize the co-production of three oxetane compounds.
  • a method for synthesizing oxetane compounds through a microreactor which includes passing trimethylolpropane and carbonate into the microreactor in the presence of a basic catalyst, and The oxetane compound is synthesized through a micro-reaction continuous flow process under solvent or solvent-free conditions.
  • the basic catalyst includes a first basic catalyst and a second basic catalyst
  • the method for synthesizing an oxetane compound through a microreactor includes: successively combining the first basic catalyst, trimethylolpropane, and carbonate. It is transported to the first microreactor for transesterification reaction to obtain the reaction product system containing the esterification intermediate; the esterification intermediate is extracted from the reaction product system containing the esterification intermediate; the esterification intermediate is combined with the second base
  • the sexual catalyst is transported to the second microreactor for cracking reaction to obtain a cracking reaction product system; the cracking reaction product system is subjected to gas-liquid separation treatment to obtain an oxetane compound.
  • the solvent is one or more of the group consisting of halogenated hydrocarbon, benzene, toluene, xylene, nitrobenzene, and acetonitrile.
  • the temperature of the first microreactor is 50-300°C, and the residence time is 1-60 min; the reaction temperature of the second microreactor is 150-400°C, and the residence time is 1-8 min.
  • the temperature of the first microreactor is 100-200°C; the reaction temperature of the second microreactor is 200-300°C.
  • the molar ratio of trimethylolpropane to carbonate is 1: (1-5), and the content of the basic catalyst is 100 ppm to 50,000 ppm.
  • the molar ratio of trimethylolpropane to carbonate is 1: (1.5-3), and the content of the basic catalyst is 100 ppm to 10000 ppm.
  • the carbonate is selected from one or more of the group consisting of dimethyl carbonate, diethyl carbonate and dipropyl carbonate;
  • the basic catalyst is selected from alkali metal hydroxides, sodium alkoxides, potassium alkoxides or One or more of alkali metal carbonates.
  • first alkaline catalyst and the second alkaline catalyst are each independently selected from one or more of alkali metal hydroxides, sodium alkoxides, potassium alkoxides, or alkali metal carbonates.
  • first The one basic catalyst and the second basic catalyst are independently selected from one selected from the group consisting of sodium methoxide, sodium ethoxide, lithium hydroxide, sodium hydroxide, potassium hydroxide, lithium carbonate, sodium carbonate and potassium carbonate Or more; preferably, the amount of the first alkaline catalyst is 200-500 ppm, and the amount of the second alkaline catalyst is 300-3000 ppm.
  • the synthesis method further includes: adding water to the esterification intermediate, and the water content of the system during the cracking reaction process is 10-100000 ppm.
  • the inner diameter of the reaction channel of the first microreactor is selected from 200 to 10000 ⁇ m
  • the inner diameter of the reaction channel of the second microreactor is independently selected from 200 to 10000 ⁇ m; preferably, the inner diameter of the reaction channel of the first microreactor is selected from 200-2000 ⁇ m, and the inner diameter of the reaction channel of the second microreactor is selected from 500-10000 ⁇ m.
  • the device used in the extraction process is selected from a thin film evaporator or a rectification tower.
  • the microreactor has the advantages of high heat and mass transfer coefficient, good mixing performance, easy temperature control, and safe and controllable process.
  • Using the advantages of the microreactor to produce the above three oxetane products can greatly improve the mass and heat transfer of the reaction system, reduce the reaction time and increase the production efficiency. In particular, it avoids the long-term high-temperature process in the cracking process and reduces The production of high boiling point by-products improves the yield, realizes the continuity and automation of the process, and improves the safety of the process.
  • the reaction device required for the above synthesis process is small in size, the production site occupies a small area, the required human resources are small, and the safety is high.
  • Fig. 1 shows a schematic structural diagram of an oxetane compound synthesis device provided according to a typical embodiment of the present invention.
  • this application provides a method for synthesizing oxetane compounds through a microreactor.
  • the synthesis method includes: in the presence of a basic catalyst, passing trimethylolpropane and carbonate into the microreactor In an inert solvent or solvent-free condition, the oxetane compound is synthesized through a micro-reaction continuous flow process.
  • microreactors Compared with conventional reactors, microreactors have the advantages of high heat and mass transfer coefficient, good mixing performance, easy temperature control, and safe and controllable process.
  • Using the advantages of the microreactor to produce the above three oxetane products can greatly improve the mass and heat transfer of the reaction system, reduce the reaction time and increase the production efficiency. In particular, it avoids the long-term high-temperature process in the cracking process and reduces
  • the production of high boiling point by-products improves the yield, realizes the continuity and automation of the process, and improves the safety of the process.
  • the reaction device required for the above synthesis process is small in size, the production site occupies a small area, the required human resources are small, and the safety is high.
  • the basic catalyst includes a first basic catalyst and a second basic catalyst
  • the method for synthesizing an oxetane compound through a microreactor includes: combining the first basic catalyst, trimethylol Propane and carbonate are continuously transported to the first microreactor for transesterification to obtain a reaction product system containing esterification intermediates; the esterification intermediate is extracted from the reaction product system containing esterification intermediates; The intermediate and the second basic catalyst are transported to the second microreactor for cracking reaction to obtain a cracking reaction product system; the cracking reaction product system is subjected to gas-liquid separation treatment to obtain an oxetane compound.
  • the product obtained is a mixture, and those skilled in the art can further separate it as needed.
  • Methods of separation include but are not limited to distillation.
  • the solvent used in the above-mentioned synthesis method can be of the type commonly used in the art.
  • the solvent includes but is not limited to one or more of the group consisting of halogenated hydrocarbons, benzene, toluene, xylene, nitrobenzene, and acetonitrile. Considering cost factors, solvent-free conditions are preferred.
  • the temperature of the first microreactor is 50-300°C, and the residence time is 1-60 min; the reaction temperature of the second microreactor is 150-400°C, and the residence time is 1-8 min.
  • the process conditions of the transesterification reaction and the cracking reaction process in this application are limited to the above range, which can improve the total yield of oxetane products while adopting the application
  • the method can also realize the control of different product distributions, realize the co-production of three oxetane compounds, and improve its economic value.
  • the temperature of the first microreactor can be 50°C, 80°C, 100°C, 160°C, 200°C, 300°C;
  • the reaction temperature of the second microreactor can be 150°C, 200°C, 260°C, 300°C, 400°C.
  • the temperature of the first microreactor is 100-200°C; the reaction temperature of the second microreactor is 200-300°C.
  • the temperature of the first microreactor and the temperature of the second microreactor include but are not limited to the above range, and limiting it to the above range is beneficial to further increase the yield of the target product and shorten the reaction time.
  • the molar ratio of trimethylolpropane to carbonate during the transesterification reaction is 1: (1-5), and the content of the basic catalyst is 100 ppm to 50,000 ppm.
  • the amount of trimethylolpropane, carbonate ester, and basic catalyst includes but is not limited to the above range, and limiting it to the above range is beneficial to further increase the conversion rate of the reaction raw materials. More preferably, the molar ratio of trimethylolpropane to carbonate is 1: (1.5-3), and the content of the basic catalyst is 100 ppm to 10000 ppm.
  • carbonate and basic catalysts can be of commonly used types in this field.
  • the carbonate includes but is not limited to one or more of the group consisting of dimethyl carbonate, diethyl carbonate, and dipropyl carbonate.
  • the first alkaline catalyst and the second alkaline catalyst are independently selected from one of hydroxides formed by alkali metals, sodium alkoxides, potassium alkoxides, and carbonates formed by alkali metals. Or multiple. More preferably, the first alkaline catalyst and the second alkaline catalyst are independently selected from the group consisting of sodium methoxide, sodium ethoxide, lithium hydroxide, sodium hydroxide, potassium hydroxide, lithium carbonate, sodium carbonate and potassium carbonate. One or more of the group.
  • the first alkaline catalyst and the second alkaline catalyst are independently selected from one of sodium methoxide, sodium ethoxide, sodium hydroxide or potassium hydroxide.
  • the amount of the first basic catalyst is 200-500 ppm, and the amount of the second basic catalyst is 300-3000 ppm.
  • the above-mentioned synthesis method further includes: adding water to the esterification intermediate, and the water content of the system during the cracking reaction process is 10-100000 ppm. Compared with other ranges, limiting the water content of the system during the cleavage reaction process within the above range is beneficial to increase the cleavage rate of the esterification intermediate, and thus the yield of the oxetane compound.
  • the inner diameter of the reaction channel of the first microreactor is selected from 200 to 10000 ⁇ m
  • the inner diameter of the reaction channel of the second microreactor is selected from 200 to 10000 ⁇ m.
  • the inner diameter of the reaction channel of the first microreactor may be selected from 200 ⁇ m, 500 ⁇ m, 1000 ⁇ m, 5000 ⁇ m, and 10000 ⁇ m
  • the inner diameter of the reaction channel of the second microreactor may be selected from 200 ⁇ m, 500 ⁇ m, 1000 ⁇ m, 8000 ⁇ m, and 10000 ⁇ m.
  • the inner diameter of the reaction channel of the first microreactor is selected from 200 to 2000 ⁇ m
  • the inner diameter of the reaction channel of the second microreactor is selected from 500 to 10000 ⁇ m.
  • the above-mentioned synthesis method further includes performing heat exchange of the cracking reaction product system in a micro heat exchanger, and removing carbon dioxide through a gas-liquid separation device. Distillation to obtain the desired product.
  • the micro-channel equipment system used in Examples 1 to 5 was provided by Shanghai Timo Fluid Technology Co., Ltd., model Shanghai Timo TMP/S3047-32-3/A2000, the channel inner diameter of the first microreactor was 1000 ⁇ m, and the second micro The inner diameter of the channel of the reactor is 8000 ⁇ m.
  • the device shown in Figure 1 is used to prepare oxetane compounds, and the synthesis method includes:
  • Transesterification section mix trimethylolpropane (TMP) and dimethyl carbonate (DMC) in a certain molar ratio, add a metered alkaline catalyst according to the weight of TMP, mix and evenly place it in the raw material storage tank 10 for preheating.
  • the first feed pump 11 is added to the first microreactor 20 to perform the transesterification reaction and stays for a certain time to obtain a reaction product system containing esterification intermediates; the above reaction product system containing esterification intermediates is separated by a thin film evaporator 30 Methanol and the remaining DMC, the recovered raw materials are placed in the light boiler collection tank 50, and the esterification intermediate is stored in the esterification intermediate storage tank 40.
  • Cracking section Add a basic catalyst to the ester intermediate, optionally add an appropriate amount of water, mix well and add it to the second microreactor 60 from the second feed pump 41, stay for a certain period of time to proceed with the cracking reaction, and undergo micro heat exchange. After the device 70 is cooled, it enters the gas-liquid separation tank 80 to separate carbon dioxide to obtain a crude product;
  • Product separation the crude product is transported to the rectification device 90 for rectification separation to obtain the corresponding product.
  • Example 1 The difference from Example 1 is: the reaction temperature of the transesterification reaction is 80°C, and the temperature of the cleavage reaction is 300°C.
  • the total yield of the products was 83.4%, and the proportions (%) of products A, B and C were 95.7%, 1.2% and 3.1%, respectively.
  • Example 2 The difference from Example 1 is that the reaction temperature of the transesterification reaction is 160°C, and the temperature of the cleavage reaction is 260°C.
  • the total yield of the products is 90.8%, and the proportions (%) of the products A, B and C are 92.0%, 4.6% and 3.4%, respectively.
  • Example 1 The difference from Example 1 is that the molar ratio of trimethylolpropane to carbonate is 1:1, and the content of the basic catalyst for the transesterification reaction is 800 ppm.
  • the total yield of the products was 80.7%, and the proportions (%) of products A, B and C were 94.5%, 3.3% and 2.2% in sequence.
  • Example 2 The difference from Example 1 is: the molar ratio of trimethylolpropane to carbonate is 1:5, and the content of the basic catalyst for the transesterification reaction is 100 ppm.
  • the total yield of the products was 84.7%, and the proportions (%) of products A, B and C were 95.1%, 2.9% and 2.0% in sequence.
  • Example 1 The difference from Example 1 is: the reaction temperature of the transesterification reaction is 50°C, the residence time of the transesterification reaction is 60 min, and the temperature of the cleavage reaction is 400°C.
  • the total yield of the products was 85.7%, and the proportions (%) of the products A, B and C were 98.3%, 1.3% and 0.4%, respectively.
  • Example 1 The difference from Example 1 is: the reaction temperature of the transesterification reaction is 300°C, the residence time of the transesterification reaction is 1 min, and the temperature of the cleavage reaction is 200°C.
  • the total yield of the products is 86.2%, and the proportions (%) of products A, B and C are 39.5%, 46.9% and 13.6%, respectively.
  • Example 1 The difference from Example 1 is: the reaction temperature of the transesterification reaction is 200°C, the residence time of the transesterification reaction is 30 min, and the temperature of the cleavage reaction is 150°C.
  • the total yield of the products is 96.8%, and the proportions (%) of the products A, B and C are 29.3%, 50.1% and 20.6% in sequence.
  • Example 2 The difference from Example 1 is that the content of the basic catalyst for the transesterification reaction is 50,000 ppm.
  • the total yield of the products is 91.3%, and the proportions (%) of the products A, B and C are 95.2%, 2.0% and 2.8%, respectively.
  • Example 2 The difference from Example 1 is that the content of the basic catalyst for the transesterification reaction is 10000 ppm.
  • the total yield of the products was 93.4%, and the proportions (%) of products A, B and C were 96.8%, 1.2% and 2.0%, respectively.
  • Example 2 The difference from Example 1 is that the content of the basic catalyst for the transesterification reaction is 30000 ppm.
  • the total yield of the products was 91.8%, and the proportions (%) of products A, B and C were 95.3%, 1.6% and 3.1%, respectively.
  • Example 2 The difference from Example 1 is that the content of the basic catalyst for the transesterification reaction is 8000 ppm.
  • the total yield of the products was 93.1%, and the proportions (%) of products A, B and C were 96.0%, 1.8% and 2.2%, respectively.
  • Example 2 The difference from Example 1 is that the content of the basic catalyst for the transesterification reaction is 5000 ppm.
  • the total yield of the products was 92.8%, and the proportions (%) of products A, B and C were 96.2%, 1.9% and 2.3%, respectively.
  • Example 1 The difference from Example 1 is that the inner diameter of the reaction channel of the first microreactor is 100 ⁇ m, and the inner diameter of the reaction channel of the second microreactor is 100 ⁇ m.
  • the total yield of the products was 81.3%, and the proportions (%) of products A, B and C were 95.6%, 1.0% and 3.4%, respectively.
  • Example 1 The difference from Example 1 is: the inner diameter of the reaction channel of the first microreactor is 10000 ⁇ m, and the inner diameter of the reaction channel of the second microreactor is 10000 ⁇ m.
  • the total yield of the products is 91.8%, and the proportions (%) of the products A, B and C are 95.5%, 2.0% and 2.5%, respectively.
  • Example 1 The difference from Example 1 is: the inner diameter of the reaction channel of the first microreactor is 200 ⁇ m, and the inner diameter of the reaction channel of the second microreactor is 200 ⁇ m.
  • the total yield of the products was 90.7%, and the proportions (%) of the products A, B and C were 94.9%, 2.1% and 3.0%, respectively.
  • Example 1 The difference from Example 1 is that the catalyst for the cracking reaction is sodium methoxide with a content of 1000 ppm, and the water content of the cracking reaction is 100000 ppm.
  • the total yield of the products was 93.5%, and the proportions (%) of the products A, B and C were 97.2%, 2.0% and 0.8%, respectively.
  • Example 1 The difference from Example 1 is that the catalyst for the transesterification reaction is sodium methoxide with a content of 10000 ppm, the catalyst for the cracking reaction is sodium methoxide with a content of 1000 ppm, and the water content of the cracking reaction is 100,000 ppm.
  • the total yield of the products was 95.8%, and the proportions (%) of products A, B and C were 97.5%, 1.9% and 0.6%, respectively.
  • Example 1 The difference from Example 1 is that the catalyst for the transesterification reaction is sodium methoxide with a content of 10000 ppm, the catalyst for the cracking reaction is sodium methoxide with a content of 300 ppm, and the water content of the cracking reaction is 10 ppm.
  • the total yield of the products is 96.5%, and the proportions (%) of products A, B and C are 3.5%, 94.6% and 1.9%, respectively.
  • Example 1 The difference from Example 1 is that the temperature of the cracking reaction is 100°C.
  • the total yield of the products is 66.8%, and the proportions (%) of products A, B and C are 3.9%, 1.7% and 94.4%, respectively.
  • Example 2 The difference from Example 1 is that the temperature of the cracking reaction is 450°C.
  • the total yield of the products is 82.9%, and the proportions (%) of products A, B and C are 43.4%, 56.5% and 0.1%, respectively.
  • Example 1 The difference from Example 1 is that the temperature of the transesterification is 30°C.
  • the total yield of the products was 78.6%, and the proportions (%) of products A, B and C were 94.1%, 2.4% and 3.5%, respectively.
  • Example 1 The difference from Example 1 is that the temperature of the transesterification is 350°C.
  • the total yield of the products is 79.0%, and the proportions (%) of products A, B and C are 94.5%, 2.6% and 2.9%, respectively.
  • a conventional reaction device was used to prepare the oxetane compound, and the synthesis method includes:
  • Transesterification section Put TMP, DMC and toluene into a stainless steel stirred tank with a rectification tower and a condenser, stir and mix evenly, add the catalyst, heat up the reaction, collect the methanol at the top of the tower, and there is no methanol at the top of the tower After distillation, continue to heat up and distill to remove the solvent toluene and the remaining DMC to obtain the esterification intermediate;

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Abstract

本发明提供了一种通过微反应器合成氧杂环丁烷化合物的方法。该合成方法包括：在碱性催化剂存在下，将三羟甲基丙烷和碳酸酯通入微反应器中，并在惰性溶剂或无溶剂条件下通过微反应连续流工艺合成所述氧杂环丁烷化合物。相对于常规反应器，微反应器具有传热传质系数高，混合性能好，温度容易控制和过程安全可控等优点。利用微反应器的优势进行上述三种氧杂环丁烷产品的生产，可大幅提高反应体系传质传热性，减少反应时间提高生产效率，尤其避免了裂解工序中长时间的高温过程，减少高沸点副产物的产生，提高了收率，并实现过程的连续化、自动化，提高过程安全性。上述合成过程所需的反应装置体积小，生产场地占用面积小，安全性高。

Description

通过微反应器合成氧杂环丁烷化合物的方法 技术领域
本发明涉及有机合成领域，具体而言，涉及一种通过微反应器合成氧杂环丁烷化合物的方法。
背景技术
3-乙基-3-羟甲基氧杂环丁烷和双[1-乙基(3-氧杂环丁基)甲基]醚是目前光固化阳离子体系使用最多的单体，被广泛应用于光固化涂料、油墨、粘胶剂等领域，它们的结构式如下所示：
Figure PCTCN2021081186-appb-000001
工业上普遍采用环碳酸酯裂解法制备氧杂环丁烷类产品。如环碳酸酯裂解法生产3-乙基-3-羟甲基氧杂环丁烷的反应过程如下：
Figure PCTCN2021081186-appb-000002
式中R为烷基，一般为甲基或乙基。
工艺流程为：三羟甲基丙烷与碳酸酯在精馏釜内进行酯交换反应，温度80℃～120℃，反应过程中不断分馏出副产物醇，反应完成后蒸馏去除过量的碳酸酯，酯交换工序需要10h～12h，之后进入裂解工序，在温度160℃～200℃条件下裂解12h～15h脱除二氧化碳，最后在负压下进行精馏得到成品。整个工序需要30h～40h才能完成，生产效率较低。裂解反应长时间在高温下进行，会导致产生高沸点的副产物，因而成品收率较低(65％～75％)，在精馏完成后产生大量蒸馏残留物，只能作为固废物处理。
鉴于上述问题的存在，有必要提供一种收率较高，且反应时间较短的氧杂环丁烷化合物的合成方法。
发明内容
本发明的主要目的在于提供一种通过微反应器合成氧杂环丁烷化合物的方法，以解决现有的氧杂环丁烷化合物的合成方法存在收率低和反应时间长的问题，在此基础上还提供了通过控制反应过程中的相关参数调节产物分布的方法，实现三种氧杂环丁烷化合物的联产。
为了实现上述目的，根据本发明提供了一种通过微反应器合成氧杂环丁烷化合物的方法，包括在碱性催化剂存在下，将三羟甲基丙烷和碳酸酯通入微反应器中，并在溶剂或无溶剂条件下通过微反应连续流工艺合成氧杂环丁烷化合物。
进一步地，碱性催化剂包括第一碱性催化剂和第二碱性催化剂，通过微反应器合成氧杂环丁烷化合物的方法包括：将第一碱性催化剂、三羟甲基丙烷和碳酸酯连续输送至第一微反应器中进行酯交换反应，得到含酯化中间体的反应产物体系；从含酯化中间体的反应产物体系中提取酯化中间体；使酯化中间体与第二碱性催化剂输送至第二微反应器中进行裂解反应，得到裂解反应产物体系；将裂解反应产物体系进行气液分离处理，得到氧杂环丁烷化合物。
进一步地，溶剂为卤代烃、苯、甲苯、二甲苯、硝基苯、乙腈组成的组中的一种或多种。
进一步地，第一微反应器的温度为50～300℃，停留时间为1～60min；第二微反应器的反应温度为150～400℃，停留时间1～8min。
进一步地，第一微反应器的温度为100～200℃；第二微反应器的反应温度为200～300℃。
进一步地，三羟甲基丙烷与碳酸酯的摩尔比为1:(1～5)，碱性催化剂的含量为100ppm～50000ppm。
进一步地，三羟甲基丙烷与碳酸酯的摩尔比为1:(1.5～3)，碱性催化剂的含量为100ppm～10000ppm。
进一步地，碳酸酯选自碳酸二甲酯、碳酸二乙酯和碳酸二丙酯组成的组中的一种或多种；碱性催化剂选自碱金属的氢氧化物、醇钠、醇钾或碱金属的碳酸盐中的一种或多种。
进一步地，第一碱性催化剂和第二碱性催化剂分别独立地选自碱金属的氢氧化物、醇钠、醇钾或碱金属的碳酸盐中的一种或多种，优选地，第一碱性催化剂和第二碱性催化剂分别独立地选自选自甲醇钠、乙醇钠、氢氧化锂、氢氧化钠、氢氧化钾、碳酸锂、碳酸钠和碳酸钾组成的组中的一种或多种；优选地，第一碱性催化剂的用量为200～500ppm，第二碱性催化剂的用量为300～3000ppm。
进一步地，在进行裂解反应过程中，合成方法还包括：向酯化中间体中加入水，且裂解反应过程中体系的含水量为10～100000ppm。
进一步地，第一微反应器的反应通道内径选自200～10000μm，第二微反应器的反应通道内径分别独立地选自200～10000μm；优选地，第一微反应器的反应通道内径选自200～2000μm，第二微反应器的反应通道内径选自500～10000μm。
进一步地，提取过程采用的装置选自薄膜蒸发器或精馏塔。
应用本发明的技术方案，相对于常规反应器，微反应器具有传热传质系数高，混合性能好，温度容易控制和过程安全可控等优点。利用微反应器的优势进行上述三种氧杂环丁烷产品的生产，可大幅提高反应体系传质传热性，减少反应时间提高生产效率，尤其避免了裂解工序中长时间的高温过程，减少高沸点副产物的产生，提高了收率，并实现过程的连续化、自动化，提高过程安全性。此外，上述合成过程所需的反应装置体积小，生产场地占用面积小，所需人力资源少，安全性高。
附图说明
构成本申请的一部分的说明书附图用来提供对本发明的进一步理解，本发明的示意性实施例及其说明用于解释本发明，并不构成对本发明的不当限定。在附图中：
图1示出了根据本发明的一种典型的实施方式提供的氧杂环丁烷化合物的合成装置的结构示意图。
其中，上述附图包括以下附图标记：
10、原料储罐；11、第一进料泵；20、第一微反应器；30、薄膜蒸发器；40、酯化中间体储罐；41、第二进料泵；50、轻沸物收集罐；60、第二微反应器；70、微换热器；80、气液分离罐；90、精馏装置。
具体实施方式
需要说明的是，在不冲突的情况下，本申请中的实施例及实施例中的特征可以相互组合。下面将结合实施例来详细说明本发明。
正如背景技术所描述的，现有的氧杂环丁烷化合物的合成方法存在收率低和反应时间长的问题。为了解决上述技术问题，本申请提供了一种通过微反应器合成氧杂环丁烷化合物的方法，该合成方法包括：在碱性催化剂存在下，将三羟甲基丙烷和碳酸酯通入微反应器中，并在惰性溶剂或无溶剂条件下通过微反应连续流工艺合成氧杂环丁烷化合物。
相对于常规反应器，微反应器具有传热传质系数高，混合性能好，温度容易控制和过程安全可控等优点。利用微反应器的优势进行上述三种氧杂环丁烷产品的生产，可大幅提高反应体系传质传热性，减少反应时间提高生产效率，尤其避免了裂解工序中长时间的高温过程，减少高沸点副产物的产生，提高了收率，并实现过程的连续化、自动化，提高过程安全性。此外，上述合成过程所需的反应装置体积小，生产场地占用面积小，所需人力资源少，安全性高。
在一种优选的实施例中，碱性催化剂包括第一碱性催化剂和第二碱性催化剂，通过微反应器合成氧杂环丁烷化合物的方法包括：将第一碱性催化剂、三羟甲基丙烷和碳酸酯连续输送至第一微反应器中进行酯交换反应，得到含酯化中间体的反应产物体系；从含酯化中间体的反应产物体系中提取酯化中间体；使酯化中间体与第二碱性催化剂输送至第二微反应器中 进行裂解反应，得到裂解反应产物体系；将裂解反应产物体系进行气液分离处理，得到氧杂环丁烷化合物。
经过上述气液分离后，得到的产物为混合物，本领域技术人员可以根据需要对其进行进一步地分离。分离的方法包括但不限于蒸馏。
上述合成方法中采用的溶剂可以采用本领域常用的种类。在一种优选的实施例中，溶剂包括但不限于卤代烃、苯、甲苯、二甲苯、硝基苯、乙腈组成的组中的一种或多种。综合成本因素考虑，优选无溶剂的条件。
在一种优选的实施例中，第一微反应器的温度为50～300℃，停留时间为1～60min；第二微反应器的反应温度为150～400℃，停留时间1～8min。相比于常规工艺中产物组成的不可控性，本申请中将酯交换反应和裂解反应过程的工艺条件限定在上述范围内能够在提高氧杂环丁烷产物总收率的同时，采用本申请的方法还可实现控制不同的产物分布，实现三种氧杂环丁烷化合物的联产，提高了其经济价值。比如第一微反应器的温度可以为50℃、80℃、100℃、160℃、200℃、300℃；第二微反应器的反应温度可以为150℃、200℃、260℃、300℃、400℃。
在一种优选的实施例中，第一微反应器的温度为100～200℃；第二微反应器的反应温度为200～300℃。第一微反应器的温度和第二微反应器的温度包括但不限于上述范围，而将其限定在上述范围内有利于进一步提高目标产物的收率，并缩短反应时间。
在一种优选的实施例中，酯交换反应过程中三羟甲基丙烷与碳酸酯的摩尔比为1:(1～5)，碱性催化剂含量为100ppm～50000ppm。三羟甲基丙烷与碳酸酯及碱性催化剂的用量包括但不限于上述范围，而将其限定在上述范围内有利于进一步提高反应原料的转化率。更优选地，三羟甲基丙烷与碳酸酯的摩尔比为1:(1.5～3)，碱性催化剂含量为100ppm～10000ppm。
上述合成方法中，碳酸酯、碱性催化剂可以采用本领域常用的种类。在一种优选的实施例中，碳酸酯包括但不限于碳酸二甲酯、碳酸二乙酯和碳酸二丙酯组成的组中的一种或多种。
在一种优选的实施例中，第一碱性催化剂和第二碱性催化剂分别独立地选自碱金属形成的氢氧化物、醇钠、醇钾、碱金属形成的碳酸盐中的一种或多种。更优选地，第一碱性催化剂和第二碱性催化剂分别独立地选自选自甲醇钠、乙醇钠、氢氧化锂、氢氧化钠、氢氧化钾、碳酸锂、碳酸钠和碳酸钾组成的组中的一种或多种。
为了进一步提高其催化效果，缩短反应时间，进一步地优选地，第一碱性催化剂和第二碱性催化剂分别独立地选自甲醇钠、乙醇钠、氢氧化钠或氢氧化钾中的一种或多种；且第一碱性催化剂的用量为200～500ppm，第二碱性催化剂的用量为300～3000ppm。
在一种优选的实施例中，在进行裂解反应过程中，上述合成方法还包括：在向酯化中间体中加入水，裂解反应过程中体系的含水量为10～100000ppm。相比于其它范围，将裂解反应过程中体系的含水量限定在上述范围内有利于提高酯化中间体的裂解率，从而有利于氧杂环丁烷化合物的收率。
在一种优选的实施例中，第一微反应器的反应通道内径选自200～10000μm，第二微反应器的反应通道内径选自200～10000μm。相比于其它范围，将第一微反应器的反应通道和第二微反应器的反应通道限定在上述范围内有利于提高目标产物的收率。比如第一微反应器的反应通道内径可以选自200μm、500μm、1000μm、5000μm、10000μm，第二微反应器的反应通道内径可以选自200μm、500μm、1000μm、8000μm、10000μm。更优选地，第一微反应器的反应通道内径选自200～2000μm，第二微反应器的反应通道内径选自500～10000μm。
为了进一步提高目标产物的纯度，并降低合成法的能量损耗，优选地，上述合成方法还包括将裂解反应产物体系在微换热器中进行换热，并通过气液分离装置去除二氧化碳后，进行精馏，得到所需的产品。
以下结合具体实施例对本申请作进一步详细描述，这些实施例不能理解为限制本申请所要求保护的范围。
实施例1至5中采用的微通道设备系统为上海替末流体技术有限公司提供，型号上海替末TMP/S3047-32-3/A2000，第一微反应器的通道内径为1000μm，第二微反应器的通道内径为8000μm。
实施例中采用图1所示的装置制备氧杂环丁烷化合物，合成方法包括：
酯交换工段：将三羟甲基丙烷(TMP)、碳酸二甲酯(DMC)按一定摩尔比混合，按TMP重量加入计量的碱性催化剂，混合均匀置于原料储罐10中预热，由第一进料泵11加入第一微反应器20进行酯交换反应，停留一定时间，得到含酯化中间体的反应产物体系；将上述含酯化中间体的反应产物体系经薄膜蒸发器30分离甲醇和剩余的DMC，将回收的原料置于轻沸物收集罐50中，酯化物中间体储存在酯化中间体储罐40中。
裂解工段：酯化物中间体中加入碱性催化剂，选择性地加入适量水，混合均匀后由第二进料泵41加入第二微反应器60中，停留一定时间进行裂解反应，经微换热器70冷却后进入气液分离罐80分离二氧化碳，得到粗产品；
产品分离：将粗产品输送至精馏装置90中进行精馏分离，得到相应的产品。
实施例1至5中的工艺参数见表1。
表1
Figure PCTCN2021081186-appb-000003
表1中产物结构如下：产物A：3-乙基-3-羟甲基氧杂环丁烷
Figure PCTCN2021081186-appb-000004
产物B：双[1-乙基(3-氧杂环丁基)甲基]醚
Figure PCTCN2021081186-appb-000005
产物C:碳酸双[1-乙基(3-氧杂环丁基)甲基]酯
Figure PCTCN2021081186-appb-000006
实施例6
与实施例1的区别为：酯交换反应的反应温度为80℃，裂解反应的温度为300℃。
产品的合计收率为83.4％，产物A、B及C的占比(％)依次为95.7％、1.2％及3.1％。
实施例7
与实施例1的区别为：酯交换反应的反应温度为160℃，裂解反应的温度为260℃。
产品的合计收率为90.8％，产物A、B及C的占比(％)依次为92.0％、4.6％及3.4％。
实施例8
与实施例1的区别为：三羟甲基丙烷与碳酸酯的摩尔比为1:1，酯交换反应碱性催化剂含量为800ppm。
产品的合计收率为80.7％，产物A、B及C的占比(％)依次为94.5％、3.3％及2.2％。
实施例9
与实施例1的区别为：三羟甲基丙烷与碳酸酯的摩尔比为1:5，酯交换反应碱性催化剂含量为100ppm。
产品的合计收率为84.7％，产物A、B及C的占比(％)依次为95.1％、2.9％及2.0％。
实施例10
与实施例1的区别为：酯交换反应的反应温度为50℃，酯交换反应的停留时间60min，裂解反应的温度为400℃。
产品的合计收率为85.7％，产物A、B及C的占比(％)依次为98.3％、1.3％及0.4％。
实施例11
与实施例1的区别为：酯交换反应的反应温度为300℃，酯交换反应的停留时间1min，裂解反应的温度为200℃。
产品的合计收率为86.2％，产物A、B及C的占比(％)依次为39.5％、46.9％及13.6％。
实施例12
与实施例1的区别为：酯交换反应的反应温度为200℃，酯交换反应的停留时间30min，裂解反应的温度为150℃。
产品的合计收率为96.8％，产物A、B及C的占比(％)依次为29.3％、50.1％及20.6％。
实施例13
与实施例1的区别为：酯交换反应碱性催化剂含量为50000ppm。
产品的合计收率为91.3％，产物A、B及C的占比(％)依次为95.2％、2.0％及2.8％。
实施例14
与实施例1的区别为：酯交换反应碱性催化剂含量为10000ppm。
产品的合计收率为93.4％，产物A、B及C的占比(％)依次为96.8％、1.2％及2.0％。
实施例15
与实施例1的区别为：酯交换反应碱性催化剂含量为30000ppm。
产品的合计收率为91.8％，产物A、B及C的占比(％)依次为95.3％、1.6％及3.1％。
实施例16
与实施例1的区别为：酯交换反应碱性催化剂含量为8000ppm。
产品的合计收率为93.1％，产物A、B及C的占比(％)依次为96.0％、1.8％及2.2％。
实施例17
与实施例1的区别为：酯交换反应碱性催化剂含量为5000ppm。
产品的合计收率为92.8％，产物A、B及C的占比(％)依次为96.2％、1.9％及2.3％。
实施例18
与实施例1的区别为：第一微反应器反应通道内径为100μm，第二微反应器反应通道内径为100μm。
产品的合计收率为81.3％，产物A、B及C的占比(％)依次为95.6％、1.0％及3.4％。
实施例19
与实施例1的区别为：第一微反应器反应通道内径为10000μm，第二微反应器反应通道内径为10000μm。
产品的合计收率为91.8％，产物A、B及C的占比(％)依次为95.5％、2.0％及2.5％。
实施例20
与实施例1的区别为：第一微反应器反应通道内径为200μm，第二微反应器反应通道内径为200μm。
产品的合计收率为90.7％，产物A、B及C的占比(％)依次为94.9％、2.1％及3.0％。
实施例21
与实施例1的区别为：裂解反应的催化剂为甲醇钠，含量为1000ppm，裂解反应的含水量为100000ppm。
产品的合计收率为93.5％，产物A、B及C的占比(％)依次为97.2％、2.0％及0.8％。
实施例22
与实施例1的区别为：酯交换反应的催化剂为甲醇钠，含量为10000ppm，裂解反应的催化剂为甲醇钠，含量为1000ppm，裂解反应的含水量为100000ppm。
产品的合计收率为95.8％，产物A、B及C的占比(％)依次为97.5％、1.9％及0.6％。
实施例23
与实施例1的区别为：酯交换反应的催化剂为甲醇钠，含量为10000ppm，裂解反应的催化剂为甲醇钠，含量为300ppm，裂解反应的含水量为10ppm。
产品的合计收率为96.5％，产物A、B及C的占比(％)依次为3.5％、94.6％及1.9％。
实施例24
与实施例1的区别为：裂解反应的温度为100℃。
产品的合计收率为66.8％，产物A、B及C的占比(％)依次为3.9％、1.7％及94.4％。
实施例25
与实施例1的区别为：裂解反应的温度为450℃。
产品的合计收率为82.9％，产物A、B及C的占比(％)依次为43.4％、56.5％及0.1％。
实施例26
与实施例1的区别为：酯交换的温度为30℃。
产品的合计收率为78.6％，产物A、B及C的占比(％)依次为94.1％、2.4％及3.5％。
实施例27
与实施例1的区别为：酯交换的温度为350℃。
产品的合计收率为79.0％，产物A、B及C的占比(％)依次为94.5％、2.6％及2.9％。
对比例中采用常规的反应装置制备氧杂环丁烷化合物，合成方法包括：
(1)酯交换工段：在带有精馏塔和冷凝器的不锈钢搅拌釜中投入TMP、DMC和甲苯，搅拌均匀混合，加入催化剂，升温反应，塔顶采集生成的甲醇，至塔顶无甲醇馏出后继续升温蒸馏去除溶剂甲苯以及剩余的DMC，得到酯化中间体；
(2)裂解工段：加入碱性催化剂，选择性地加入适量水，在一定温度和负压下进行裂解，同时蒸馏得到产品。对比例1至5中的工艺参数见表2。
表2
Figure PCTCN2021081186-appb-000007
从对比实施例的结果可以看出：采用常规釜式反应器制备收率低于微通道设备系统的收率，且双[1-乙基(3-氧杂环丁基)甲基]醚选择性明显低于预期，无法达到产物分布的设计要求。
从以上的描述中，可以看出，比较实施例1至27及对比例1至5可知，本发明上述的实施例实现了如下技术效果：采用本申请提供的合成方法能够大大提高三种氧杂环丁烷化合物的收率，缩短反应时间，同时还能够实现三种氧杂环丁烷化合物的联产。
比较实施例1、6、7、10至12、24至27可知，将酯交换反应和裂解反应的温度限定在本申请的优选范围内有利于提高三种氧杂环丁烷化合物的总收率，同时还可以通过控制裂解温度控制三种产物的比例。
比较实施例1、8、9、13至17、21及23可知将三羟甲基丙烷与所述碳酸酯的摩尔比、酯交换反应中使用的碱性催化剂的含量及裂解反应中使用的碱性催化剂的含量限定在本申请优选的范围内有利于提高三种氧杂环丁烷化合物的总收率。
比较实施例1、18、19及20可知将第一微反应器的通道内径和第二微反应器的通道内径限定在本申请优选的范围内有利于提高三种氧杂环丁烷化合物的总收率。
以上所述仅为本发明的优选实施例而已，并不用于限制本发明，对于本领域的技术人员来说，本发明可以有各种更改和变化。凡在本发明的精神和原则之内，所作的任何修改、等同替换、改进等，均应包含在本发明的保护范围之内。

Claims (12)

  1. 一种通过微反应器合成氧杂环丁烷化合物的方法，其特征在于，所述通过微反应器合成氧杂环丁烷化合物的方法在碱性催化剂存在下，将三羟甲基丙烷和碳酸酯通入所述微反应器中，并在溶剂或无溶剂条件下通过微反应连续流工艺合成所述氧杂环丁烷化合物。
  2. 根据权利要求1所述的通过微反应器合成氧杂环丁烷化合物的方法，其特征在于，所述碱性催化剂包括第一碱性催化剂和第二碱性催化剂，所述通过微反应器合成氧杂环丁烷化合物的方法包括：
    将所述第一碱性催化剂、所述三羟甲基丙烷和所述碳酸酯连续输送至第一微反应器中进行酯交换反应，得到含酯化中间体的反应产物体系；
    从所述含酯化中间体的反应产物体系中提取酯化中间体；
    使所述酯化中间体与所述第二碱性催化剂输送至第二微反应器中进行裂解反应，得到裂解反应产物体系；
    将所述裂解反应产物体系进行气液分离处理，得到所述氧杂环丁烷化合物。
  3. 根据权利要求1所述的通过微反应器合成氧杂环丁烷化合物的方法，其特征在于，所述溶剂为卤代烃、苯、甲苯、二甲苯、硝基苯、乙腈组成的组中的一种或多种。
  4. 根据权利要求2所述的通过微反应器合成氧杂环丁烷化合物的方法，其特征在于，所述第一微反应器的温度为50～300℃，停留时间为1～60min；所述第二微反应器的反应温度为150～400℃，停留时间1～8min。
  5. 根据权利要求4所述的通过微反应器合成氧杂环丁烷化合物的方法，其特征在于，所述第一微反应器的温度为100～200℃；所述第二微反应器的反应温度为200～300℃。
  6. 根据权利要求2、4至5中任一项所述的通过微反应器合成氧杂环丁烷化合物的方法，其特征在于，所述三羟甲基丙烷与所述碳酸酯的摩尔比为1:(1～5)，所述碱性催化剂的含量为100ppm～50000ppm。
  7. 根据权利要求6所述的通过微反应器合成氧杂环丁烷化合物的方法，其特征在于，所述三羟甲基丙烷与所述碳酸酯的摩尔比为1:(1.5～3)，所述碱性催化剂的含量为100ppm～10000ppm。
  8. 根据权利要求6或7所述的通过微反应器合成氧杂环丁烷化合物的方法，其特征在于，所述碳酸酯选自碳酸二甲酯、碳酸二乙酯和碳酸二丙酯组成的组中的一种或多种；
    所述碱性催化剂选自碱金属的氢氧化物、醇钠、醇钾、碱金属的碳酸盐中的一种或多种。
  9. 根据权利要求8所述的通过微反应器合成氧杂环丁烷化合物的方法，其特征在于，所述第一碱性催化剂和所述第二碱性催化剂分别独立地选自碱金属的氢氧化物、醇钠、醇钾、碱金属的碳酸盐中的一种或多种，优选地，所述第一碱性催化剂和所述第二碱性催化剂 分别独立地选自选自甲醇钠、乙醇钠、氢氧化锂、氢氧化钠、氢氧化钾、碳酸锂、碳酸钠和碳酸钾组成的组中的一种或多种；
    优选地，所述第一碱性催化剂的用量为200～500ppm，所述第二碱性催化剂的用量为300～3000ppm。
  10. 根据权利要求2所述的通过微反应器合成氧杂环丁烷化合物的方法，其特征在于，在进行所述裂解反应过程中，所述合成方法还包括：向所述酯化中间体中加入水，且裂解反应过程中体系的含水量为10～100000ppm。
  11. 根据权利要求2至5中任一项所述的通过微反应器合成氧杂环丁烷化合物的方法，其特征在于，所述第一微反应器的反应通道内径选自200～10000μm，所述第二微反应器的反应通道内径选自200～10000μm；
    优选地，所述第一微反应器的反应通道内径选自200～2000μm，所述第二微反应器的反应通道内径选自500～10000μm。
  12. 根据权利要求2所述的通过微反应器合成氧杂环丁烷化合物的方法，其特征在于，所述提取过程采用的装置选自薄膜蒸发器或精馏塔。
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