Preparation system and preparation method for 1,3-propanediol

Abstract

The present invention relates to the field of chemical synthesis. Disclosed are a preparation system and a preparation method for 1,3-propanediol (PDO). The preparation system comprises a hydrogenation section, and the hydrogenation section comprises a hydrogenation reactor, a PDO refining tower, a polypropylene glycol recovery tower, a hydrolysis reactor, and a dehydration tower which are sequentially communicated. A tower bottom outlet of the dehydration tower is communicated with an inlet of the PDO refining tower. The PDO refining tower is provided with a 1,3-PDO product outlet. The present invention has the advantages of improving the yield and the purity of 1,3-PDO products.

Description

Preparation system and method of 1,3-propanediol

[0001] Cross-references to related applications

[0002] This application claims the benefit of Chinese Patent Application No. 202510028361.5, filed on January 8, 2025, entitled "Preparation System and Method of 1,3-Propanediol", the contents of which are incorporated herein by reference. Technical Field

[0003] This invention relates to the field of chemical synthesis, and more specifically to a preparation system and method for 1,3-propanediol. Background Technology

[0004] 1,3-Propanediol (1,3-PDO) is an important chemical raw material used in industries such as inks, coatings, lubricants, and antifreeze. It is also used as a pharmaceutical synthesis intermediate, and its applications are quite broad. However, most 1,3-PDO is mainly used in the production of polyethylene terephthalate (PTT), followed by polyethylene naphthalate (PTN), coatings, and cosmetic additives.

[0005] Currently, there are three main production methods for 1,3-PDO: microbial fermentation (MF method), ethylene oxide carbonylation (EO method), and acrolein hydration hydrogenation (AC method).

[0006] The fermentation molecule (MF) method is characterized by mild conditions and independence from petroleum products, but it suffers from bottlenecks such as expensive strains, long fermentation cycles, low product concentration in the fermentation broth, complex composition, and difficulties in product separation, which are challenging to meet the requirements of conventional distillation methods. Its cost is also not yet competitive with chemical methods. See DuPont's MF method patents EP361082A2 and DE3734764A. In China, CN102250968A discloses an MF method for producing fermentation broth containing 1,3-PDO, with a maximum fermentation yield of 52.8%; CN102070402A discloses a subsequent fermentation broth separation process—electrodialysis—for separating 1,3-PDO, with a separation yield of approximately 92.7%. CN104774879A discloses a fermentation broth containing 1,3-PDO prepared by the MF method, with a fermentation yield of up to 50%; CN101012151A discloses a subsequent fermentation broth separation process—aqueous two-phase extraction separation to prepare 1,3-PDO, with a separation yield of about 92%.

[0007] The EO process involves large equipment investment, high technical difficulty, high reaction pressure, complex reactor structure and catalyst system, and a demanding and unstable process. The high aldehyde content in the product makes it difficult to meet the purity requirements of polymer-grade 1,3-PDO for fiber applications; the reported purity is 99.6%, with a product yield of 85-90%. See Shell's EO process patents US5723389A and US5777182A. Foreign EO process plants ceased operation in 2008.

[0008] The AC process has mild conditions, inexpensive and readily available propylene feedstock, and mature propylene oxidation and hydrogenation processes, requiring relatively low equipment specifications. However, the hydration process involves a high water content in the medium, resulting in high separation energy consumption. Degussa's AC process unit is still in operation (see US6232511B1 and US6140543A). However, Degussa's hydration unit has a hydration conversion rate of 85% and a total product yield of 83.7%. The research report "New Technology Development and Research for the Hydrogenation of Acrolein to 1,3-Propanediol" reports a comprehensive energy consumption of 2990.58 kg standard oil / t product for the hydration process, with energy consumption alone accounting for 31% of the total cost, approximately RMB 4272 / t product. Summary of the Invention

[0009] The purpose of this invention is to overcome the problems of low overall yield of 1,3-PDO and product purity that only meets the purity requirements of HG / T4980 (≥99.5%) in the existing technology, and to provide a preparation system and method for 1,3-propanediol, which has the advantage of improving the yield and purity of 1,3-PDO products.

[0010] To achieve the above objectives, the present invention provides a 1,3-propanediol preparation system, the preparation system comprising a hydrogenation section, the hydrogenation section comprising a hydrogenation reactor, a propylene glycol refining tower, a polypropylene glycol recovery tower, a hydrolysis reactor, and a dehydration tower connected in sequence, the bottom outlet of the dehydration tower being connected to the inlet of the propylene glycol refining tower, and the propylene glycol refining tower being provided with a 1,3-propanediol product outlet.

[0011] A second aspect of the present invention provides a method for preparing 1,3-propanediol, wherein the method employs the 1,3-propanediol preparation system described herein, and the method comprises: catalytically hydrogenating 3-hydroxypropanal feedstock in a hydrogen atmosphere in a hydrogenation reactor to obtain a mixture stream containing 1,3-propanediol and its dimer; purifying the mixture stream in a propylene glycol purification tower to obtain 1,3-propanediol product and 1,3-propanediol dimer; and hydrolyzing and dehydrating the 1,3-propanediol dimer in a hydrolysis reactor and a dehydration tower, and then returning the dehydrated product to the propylene glycol purification tower for further purification to obtain the 1,3-propanediol product.

[0012] Through the above technical solution, the present invention uses a polypropylene glycol recovery tower to separate the AC method to prepare the dimer of 1,3-propanediol generated by hydrogenation in 1,3-PDO. The dimer is depolymerized and 1,3-propanediol is recovered by using a hydrolysis reactor and a dehydration tower, so as to increase the yield and purity of 1,3-propanediol product. Attached Figure Description

[0013] Figure 1 is a schematic diagram of a 1,3-propanediol preparation system according to one embodiment of the present invention.

[0014] Explanation of reference numerals in the attached figures

[0015] 1. Acrolein reactor; 2. Acid washing tower; 3. Stripping tower; 4. Absorption tower; 5. Desorption tower; 6. Catalytic oxidation unit; 7. Wastewater recovery unit; 8. 3-HPA reactor; 9. Acrolein recovery tower; 10. Extraction tower; 11. Extract phase dehydration tower; 12. Extractant recovery tower; 13. Hydrogenation reactor; 14. Supplementary purification reactor; 15. Gas-liquid separator; 16. De-alcoholization tower; 17. Propylene glycol purification tower; 18. Polypropylene glycol recovery tower; 19. Hydrolysis reactor; 20. Dehydration tower; 21. Heat pump energy recovery system. Detailed Implementation

[0016] The specific embodiments of the present invention will be described in detail below with reference to the accompanying drawings. It should be understood that the specific embodiments described herein are for illustration and explanation only and are not intended to limit the present invention. Unless otherwise specified, all material proportions in the present invention are based on pure substances.

[0017] This invention discloses a 1,3-propanediol preparation system. The system includes a hydrogenation section comprising a hydrogenation reactor, a propylene glycol refining tower 17, a polypropylene glycol recovery tower 18, a hydrolysis reactor 19, and a dehydration tower 20, connected sequentially. The bottom outlet of the dehydration tower 20 is connected to the inlet of the propylene glycol refining tower 17, and the top of the propylene glycol refining tower 17 has a 1,3-propanediol product outlet. As shown in Figure 1, the top of the propylene glycol refining tower 17 has a 1,3-propanediol product outlet pipeline. The pipeline connecting the top of the polypropylene glycol recovery tower 18 and the inlet of the hydrolysis reactor 19 has a deoxygenated water inlet and can be fitted with a heater. The top of the dehydration tower 20 can be connected to the inlet of the aforementioned heater via a pipeline.

[0018] It should be noted that in the acrolein hydration hydrogenation process for producing 1,3-PDO, the hydrogenation of 3-hydroxypropanal (3-HPA) generates a portion of 1,3-propanediol dimers, affecting the final product yield. This invention employs a polypropylene glycol recovery tower to separate this component, utilizing a hydrolysis reactor and dehydration tower to depolymerize the dimers and recover 1,3-propanediol, thereby increasing the 1,3-propanediol product yield and purity. The 1,3-propanediol dimers are polypropylene glycol-based dimer products formed by the intermolecular condensation of 1,3-propanediol molecules.

[0019] In some embodiments, the hydrogenation reactor includes a hydrogenation reactor 13 and a supplementary purification reactor 14 connected in series, thereby performing 3-HPA hydrogenation in stages, whereby hydrogenation can be carried out at low temperature in the hydrogenation reactor 13 and supplementary purification can be performed at high temperature in the supplementary purification reactor 14. As needed, as shown in FIG1, a feed heater can be installed on the feed line of the hydrogenation reactor 13, and a second feed heater can be installed on the connecting line between the hydrogenation reactor 13 and the supplementary purification reactor 14, with the hydrogen feed line connected to the inlet of each of the first and second feed heaters.

[0020] In some embodiments, a cooler and a gas-liquid separator 15 are preferably installed sequentially on the connecting pipeline between the supplementary refining reactor 14 and the propylene glycol refining tower 17. The liquid phase outlet and gas phase outlet of the gas-liquid separator 15 are each connected to the feed pipeline of the hydrogenation reactor 13. During hydrogenation, the reactants are circulated and mixed with the liquid phase of the gas-liquid separator 15 to reduce the 3-HPA concentration, preventing over-hydrogenation and polymer blockage of the catalyst pores. In the hydrogenation reactor 13, some 1,3-propanediol is generated at low temperature to further reduce the 3-HPA concentration. Then, it enters the supplementary refining reactor 14 at high temperature to improve product selectivity and reactant conversion rate. In addition, a circulating hydrogen compressor is installed on the connecting pipeline between the gas phase outlet and the feed pipeline of the hydrogenation reactor 13. Unreacted hydrogen is returned to the feed inlet of the hydrogenation reactor 13 after being pressurized by the circulating hydrogen compressor from the top of the gas-liquid separator 15 to reduce the amount of fresh hydrogen added.

[0021] In this invention, 3-hydroxypropionaldehyde (3-HPA) can be prepared using the apparatus described below. In some embodiments, the preparation system includes a hydration section, which includes a 3-HPA reactor 8 and an extraction tower 10 connected together. The top of the extraction tower 10 is connected to an extraction phase dehydration tower 11, and the bottom outlet of the extraction phase dehydration tower 11 is connected to the feed line of the hydrogenation reactor 13. Thus, 3-HPA is produced by hydration reaction of acrolein in the 3-HPA reactor 8. After extraction and separation by an alcohol extractant in the extraction tower 10, the alcohol solution rich in 3-HPA obtained at the top of the tower is dehydrated by the extraction phase dehydration tower 11 and then sent to the hydrogenation reactor 13.

[0022] In some embodiments, a cooler and an acrolein recovery tower 9 are installed on the connecting pipeline between the 3-HPA reactor 8 and the extraction tower 10. The top of the acrolein recovery tower 9 is provided with a polymerization inhibitor inlet and the top outlet is connected to the feed inlet of the 3-HPA reactor 8. The acrolein obtained at the top of the tower is returned to the feed of the 3-HPA reactor 8 and mixed with fresh acrolein.

[0023] In some embodiments, the bottom outlet of the extraction tower 10 is connected to the extractant recovery tower 12, and the top outlet of the extractant recovery tower 12 is connected to the bottom inlet of the extraction tower 10.

[0024] In some embodiments, a dealcoholization tower 16 is installed after the gas-liquid separator 15 along the material flow direction. The bottom outlet of the dealcoholization tower 16 is connected to the inlet of the propylene glycol refining tower 17, and the bottom inlet of the extraction tower 10 is connected to the top outlet of the dealcoholization tower 16. After the hydrogenation reaction product is degassed by the gas-liquid separator 15, it enters the dealcoholization tower 16 to remove alcohols, and then enters the propylene glycol refining tower 17 for refining. The alcohols separated by the dealcoholization tower 16 are returned to the extraction tower 10 to realize the recycling of the extractant.

[0025] In some embodiments, the bottom inlet of the extraction tower 10 is connected to the top outlet of the extraction phase dehydration tower 11.

[0026] To improve energy efficiency, in some embodiments, based on the aforementioned improvements, the extraction phase dehydration tower 11 is equipped with a heat pump energy recovery system 21. The heat pump energy recovery system 21 is used to extract heat from the top of the extraction phase dehydration tower 11 to heat the bottom of the tower. Specifically, the heat pump energy recovery system 21 can be an open-loop heat pump: utilizing the heat from the compressed gas phase at the top of the extraction phase dehydration tower 11 to heat the reboiler at the bottom of the tower; the condensed liquid phase is partially returned to the top of the tower, and the other part goes to the extraction tower 10. Alternatively, the heat pump energy recovery system 21 can be a closed-loop heat pump: using propane, butane, or a mixture of both as the heat pump organic working fluid, absorbing heat from the condensed gas at the top of the tower and vaporizing it, then using a compressor to increase the saturated pressure and temperature of the working fluid for heating the reboiler at the bottom of the tower. The heated liquid working fluid is depressurized by a regulating valve and then returned to the top of the tower for heat extraction.

[0027] It should be noted that in existing industrialized technologies, after acrolein is recovered in an acrolein recovery tower, 3-hydroxypropanal and water undergo a hydrogenation reaction. After hydrogenation, water and products are separated in a dehydration tower. The temperature difference between the top and bottom of the acrolein recovery tower is 26°C, and the temperature difference between the top and bottom of the dehydration tower is even greater, reaching 110°C. A larger temperature difference makes it less economical to install a heat pump (heat pump installation involves pressurizing and heating the top gas to heat the bottom; the higher the bottom temperature, the higher the compressor pressure ratio is required, thus a larger temperature difference leads to lower economic efficiency). This process uses an organic extractant to separate 3-hydroxypropanal and water, achieving separation before the hydrogenation reaction. The separation of hydroxypropionaldehyde and water involves separating the organic extractant and hydrogenation product after the hydrogenation reaction. In industrialized technologies, this is equivalent to all the remaining water from the hydration reaction being distilled off the top of the dehydration tower after hydrogenation and returned to the 3-HPA reactor. The new technology uses an extraction route; after the hydration reaction, some water returns from the bottom of the extractant recovery tower to the inlet of the 3-HPA reactor, eliminating the need for evaporation to the top. The majority of the water evaporated from the top of the towers before and after the hydrogenation reactor is the organic extractant. The enthalpy of vaporization of the organic extractant (582 kJ / kg) is smaller than that of water (2256 kJ / kg), making the process relatively energy-efficient. Furthermore, the temperature difference between the top and bottom of the extraction phase dehydration tower is approximately 20°C, making it more suitable for installing a heat pump compared to industrialized technologies.

[0028] Furthermore, since the acrolein recovery tower 9 in this invention operates under reduced pressure and high vacuum, an open-loop heat pump would utilize the overhead gas from the tower to pressurize the compressor and then heat the tower kettle for reboiling. However, the pressurized vacuum operation results in a large volume of overhead gas, naturally leading to a large compressor. In contrast, a closed-loop heat pump uses a compressor to compress butane. Butane is not under vacuum, so its volume is much smaller. A small compressor is sufficient to meet the requirements, resulting in a significant difference in equipment costs. Therefore, a closed-loop heat pump is the most economical option.

[0029] In some embodiments, the bottom outlet of the extractant recovery tower 12 is connected to the feed inlet of the 3-HPA reactor 8 to recycle the aqueous activator solution after alcohol removal.

[0030] Considering that propylene is readily available and much cheaper than acrolein, using propylene as a raw material is more economical and feasible than producing propylene glycol only through subsequent hydration and hydrogenation of acrolein. Therefore, acrolein in the hydration stage can be prepared using the apparatus described below in this invention. In some embodiments, the preparation system includes an oxidation stage, which includes an acrolein reactor 1, a steam generator, an acid washing tower 2, an absorption tower 4, and a desorption tower 5 connected in sequence. The acrolein outlet of the desorption tower 5 is connected to the inlet of the 3-HPA reactor 8. Propylene is oxidized in acrolein reactor 1. The effluent from acrolein reactor 1 first enters acid washing tower 2, where acetic acid, acrylic acid, and a small amount of acrolein are washed away with circulating acidic water. Acid washing tower 2 has an external circulation pipeline at its bottom, and a stripping tower is installed on this pipeline. The acidic water from the bottom of acid washing tower 2 then enters stripping tower 3 to recover acrolein. A low-pressure steam inlet is located at the bottom of stripping tower 3, and a wastewater discharge pipeline with a wastewater recovery device 7 is installed at the bottom of stripping tower 3. The acidic wastewater is sent to the wastewater recovery device 7, and the gas from the top of stripping tower 3 returns to acid washing tower 2 through the external circulation pipeline. The reaction gas, after the acrylic acid and acetic acid have been removed in acid washing tower 2, enters absorption tower 4. A liquid disperser, such as an absorbent sprayer, can be installed at the top of absorption tower 4. The reaction gas from acid washing tower 2 is fed from the bottom of absorption tower 4 and absorbed by the absorbent (water) flowing downwards, generating an acrolein aqueous solution. After being preheated by a heat exchanger, the acrolein aqueous solution enters the desorption tower 5, and the acrolein collected at the top of the tower enters the 3-HPA reactor 8.

[0031] To further reduce energy consumption, a heat exchanger tube side can be installed on the connecting pipeline between absorption tower 4 and desorption tower 5. The shell side of the heat exchanger can be connected in series on the connecting pipeline between the absorbent inlet at the top of absorption tower 4 and the lean liquid outlet of desorption tower 5. While circulating the bottom water of desorption tower 5 back to the absorbent inlet at the top of absorption tower 4, the heat from the discharge of desorption tower 5 is used to preheat the feed of desorption tower 5.

[0032] The present invention provides a polymerization inhibitor inlet at the top of the desorption tower 5. A polymerization inhibitor needs to be added to the top of the desorption tower 5 to prevent polymerization from occurring inside the tower.

[0033] In some embodiments, the wastewater recovery device 7 includes a wastewater extraction tower, a wastewater extractant recovery tower, and an acrylic acid dehydration tower connected in sequence. The wastewater extraction tower can use an ester extractant to extract acrylic acid from the wastewater. The ester extractant can be isopropyl acetate and / or n-butyl acetate. The mass ratio of extractant to acrylic acid in the wastewater extraction tower is 5–10:1. The extract phase after liquid-liquid extraction in the wastewater extraction tower enters the wastewater extractant recovery tower. The extractant is recovered at the top of the tower for recycling, and high-purity acrylic acid is recovered at the bottom. The raffinate phase enters the acrylic acid dehydration tower. The azeotrope of extractant and water is recovered at the top of the tower, and wastewater is discharged at the bottom. It should be noted that the present invention does not have special requirements for the operating conditions of the aforementioned wastewater extraction tower, wastewater extractant recovery tower, and acrylic acid dehydration tower; conventional methods in the prior art can be used.

[0034] In some embodiments, the acrolein reactor 1 can be a tubular reactor, with spherical coated catalysts packed inside the tubes. It consists of a reaction section and a cooling section. The shell side of the reaction section can be filled with molten salt, and the reaction temperature (hot spot temperature 370°C) is controlled by controlling the circulation rate of the molten salt. To suppress excessive oxidation of acrolein in the reaction product, heat is removed by forced circulation using a molten salt pump in the cooling section, for example, rapidly reducing the outlet temperature of the acrolein reactor 1 to below 260°C. The reaction section of the acrolein reactor 1 is equipped with a molten salt heater. Before start-up, the molten salt is heated to the propylene feed temperature using the molten salt heater. During shutdown, the liquid molten salt is heated using the molten salt heater to maintain the temperature of the hot molten salt.

[0035] In some embodiments, the bottom inlet of the absorption tower 4 is connected to the top outlet of the acrolein recovery tower 9, so that the acrolein recovered in the hydration section can be recycled back to the absorption tower 4 in the oxidation section.

[0036] In some embodiments, the acrolein reactor 1 is equipped with a mixing pipeline, the feed end of which is connected to a propylene feed pipeline and an air feed pipeline. Along the material flow direction in the pipeline, a propylene evaporator and a propylene superheater can be installed sequentially on the propylene feed pipeline, so that liquid propylene can be fed into the feed end of the propylene feed pipeline. An air compressor and an air preheater can be installed sequentially on the air feed pipeline. In order to control the concentration of propylene entering the acrolein reactor 1 to be within the explosion limit range, a certain amount of low-pressure steam needs to be introduced to mix with propylene and air to adjust the propylene feed concentration. For this purpose, the feed end of the mixing pipeline is connected to a steam feed pipeline.

[0037] In some embodiments, the absorption tower 4 is provided with a top material discharge pipeline. A compressor and a catalytic oxidation device 6 can be installed sequentially along the material flow direction on the discharge pipeline. A branch extends between the compressor and the catalytic oxidation device 6 and connects to the inlet of the acrolein reactor 1. After the air and carbon oxides at the top of the absorption tower 4 are pressurized by the compressor, part of them are circulated and mixed with fresh air to control the concentration of the propylene in the reaction feed. The other part enters the catalytic oxidation device 6 and is discharged as exhaust gas after catalytic oxidation.

[0038] To further reduce energy consumption, in some embodiments, the bottom material of the desorption tower is introduced into the propylene evaporator through a pipeline after exchanging heat with the feed of the desorption tower via a heat exchanger, and then exchanged heat with the liquid propylene before returning to the top of the acid washing tower 2.

[0039] Based on the foregoing disclosure, this invention discloses a method for preparing 1,3-propanediol. This method employs the preparation system of this invention. The method includes: catalytically hydrogenating 3-hydroxypropanal feedstock in a hydrogenation reactor under a hydrogen atmosphere to obtain a mixture stream containing 1,3-propanediol and its dimer; purifying the mixture stream in a propylene glycol purification tower 17 to obtain 1,3-propanediol product and 1,3-propanediol dimer; and hydrolyzing and dehydrating the 1,3-propanediol dimer through a hydrolysis reactor 19 and a dehydration tower 20, then returning the dehydrated product to the propylene glycol purification tower 17 for further purification to obtain the 1,3-propanediol product.

[0040] In some embodiments, the catalytic hydrogenation reaction includes: mixing 3-hydroxypropionaldehyde feedstock with hydrogen gas, heating the mixture to 30–40°C, and introducing it into the hydrogenation reactor 13 at a pressure of 1–3 MPaG, with a weight hourly space velocity (WHSV) of 0.6–0.8 h⁻¹. -1 The temperature rise of the hydrogenation reaction is controlled by partially recycling the reaction products (not exceeding 25°C). Unreacted 3-hydroxypropanal enters the supplementary purification reactor 14 to continue the reaction, with a weight hourly space velocity of 0.5–2.8 h⁻¹. -1 In the hydrogenation reactor 13, the molar ratio of 3-hydroxypropionaldehyde to hydrogen is 1:40 to 50. The range of hydrogenation catalysts that can be selected is relatively wide, such as precious metal or non-precious metal catalysts commonly used in this field.

[0041] It should be noted that, as shown in Figure 1, hydrogen enters the hydrogenation reactor 13 and the supplementary purification reactor 14. After separation by the gas-liquid separator 15, the hydrogen-containing circulating gas is recycled back to the inlet of the hydrogenation reactor 13 via the circulating hydrogen compressor. In this invention, the main hydrogenation reaction occurs in the first stage hydrogenation reaction. Therefore, the temperature rise in the first stage is mainly controlled. The temperature rise in the second stage is small and does not need to be deliberately controlled. The temperature rise of the reactor can be controlled by adjusting the amount of circulating hydrogen. Similarly, the temperature rise in the first stage requires the ratio of hydrogen to 3-hydroxypropanal to be controlled. In the second stage, the remaining amount of 3-hydroxypropanal is very small and reacts with the remaining hydrogen in the first stage reaction, resulting in a very low temperature rise. Therefore, the ratio does not need to be controlled.

[0042] In some embodiments, the effluent from the supplementary refining reactor 14 undergoes gas-liquid separation and refining sequentially through a gas-liquid separator 15 and a propylene glycol refining tower 17 to obtain 1,3-propanediol product and 1,3-propanediol dimer. The operating conditions for gas-liquid separation and refining can be selected over a wide range. The following is an illustrative description, but it does not limit the scope of the invention. For this invention, the operating conditions of the gas-liquid separator 15 can be 40–50°C and a pressure of 1.5–2.5 MPa. The operating conditions of the propylene glycol refining tower 17 include the tower being equipped with structured packing, a tower pressure controlled at 5–15 kPaA, and a tower top temperature of 110–170°C.

[0043] In some embodiments, the gas phase and / or liquid phase after gas-liquid separation are returned to the hydrogenation reactor 13. It should be noted that the return of the gas and liquid phases only affects the reactor temperature rise and reactor configuration, while the yield and purity will change to some extent. Conventional gas phase circulation requires a trickle bed with a circulating hydrogen compressor, while liquid phase circulation requires a liquid bed with a micro-dispersor and a liquid phase circulation pump. The circulating liquid phase:feed mass ratio is (0.9-1.1):1, for example, it can be 1:1. Both forms are feasible, but the liquid phase is more economical. However, in this invention, the total conversion rate and selectivity are quite high after two stages of hydrogenation, close to 100%, indicating that the reaction is relatively easy to carry out. After subsequent distillation processes, the changes in reactor outlet yield and purity have a relatively small impact on the final product outlet. Therefore, this invention does not particularly limit the return ratio of the gas phase and / or liquid phase after gas-liquid separation. For example, the return ratio can be a molar ratio of 3-hydroxypropanal to circulating hydrogen of 1:40-50.

[0044] In some embodiments, the dimer of 1,3-propanediol is mixed with water and heated to 230–250°C, and introduced into the hydrolysis reactor 19 at a pressure of 5–6 MPaG, with a weight hourly space velocity of 0.01–1 h⁻¹. -1 The water in the hydrolysis reactor is in excess; for example, the water circulation rate to feed rate mass ratio can be 3 to 4:1. The operating conditions of the hydrolysis reactor 19 can be selected from a wide range. The following is an illustrative description, but it does not limit the scope of the invention. For the present invention, the hydrolysis catalyst can be selected from a wide range, for example, it can be at least one of molecular sieves or resin catalysts commonly used in the art.

[0045] The operating conditions of the dehydration tower 20 can be selected from a wide range. The following is an illustrative description, but it does not limit the scope of the present invention. For the present invention, the operating conditions of the dehydration tower 20 include that the tower is equipped with structured packing, the tower pressure is controlled at 10 to 30 kPaA, and the tower top temperature is 50 to 60°C.

[0046] In some embodiments, in the 3-HPA reactor 8, acrolein is hydrated in the presence of an activator to obtain 3-HPA. The effluent from the 3-HPA reactor 8 enters the extraction tower 10, where the hydrated product is extracted using an alcohol extractant. The extracted phase enters the extraction phase dehydration tower 11 for dehydration, and the bottom of the tower yields an alcoholic solution of 3-hydroxypropionaldehyde, which is then fed into the hydrogenation reactor 13 as the 3-hydroxypropionaldehyde feedstock.

[0047] When the 3-hydroxypropanal feedstock is an alcoholic solution of 3-hydroxypropanal, in some embodiments, the effluent from the supplementary refining reactor 14 undergoes gas-liquid separation, de-alcoholization, and refining sequentially through a gas-liquid separator 15, a dealcoholization tower 16, and a propylene glycol refining tower 17 to obtain 1,3-propanediol product and 1,3-propanediol dimer. The 1,3-propanediol dimer then flows through a subsequent polypropylene glycol recovery tower 18, a hydrolysis reactor 19, and a dehydration tower 20 before returning to the propylene glycol refining tower 17 for further refining to improve the 1,3-propanediol product yield. The top material from the dealcoholization tower 16 is returned to the extraction tower 10. The polypropylene glycol from the polypropylene glycol recovery tower 18 is mixed with water and heated to 230–270°C (e.g., 230–250°C) and a pressure of 5–6 MPaG before entering the hydrolysis reactor 19 at a weight hourly space velocity (WHSV) of 0.01–1 h⁻¹. -1 The crude 1,3-propanediol is then separated by dehydration tower 20 and then recycled to propylene glycol refining tower 17 to recover 1,3-propanediol.

[0048] The 3-HPA generated by hydration using this invention exhibits high selectivity. After extraction with an alcohol extractant, the effluent from the 3-HPA reactor 8 is separated into 3-HPA and an activator. In some embodiments, the aqueous solution obtained at the bottom of the extraction tower 10 enters the extractant recovery tower 12 to remove alcohol, and the activator aqueous solution is returned to the 3-HPA reactor 8 for recycling. The introduction of 3-HPA and the alcohol extractant into the hydrogenation reactor 13 prevents the loss of active components from the hydrogenation catalyst present in the conventional aqueous phase hydrogenation process, thus improving the catalytic efficiency and lifespan of the hydrogenation catalyst. Furthermore, the latent heat of vaporization of the alcohol separated at the top of the subsequent hydrogenation section's alcohol removal tower is lower than that of the conventional process for separating water, resulting in relatively lower energy consumption.

[0049] The operating conditions of the dealcoholization tower 16 can be selected from a wide range. The following is an illustrative description, but it does not limit the scope of the present invention. For the present invention, the operating conditions of the dealcoholization tower 16 include: the tower is equipped with structured packing, the tower pressure is controlled at 10-20 kPaA, and the tower top temperature is 60-80°C.

[0050] In some embodiments, the feed temperature of the 3-HPA reactor 8 is 50–60°C, the pressure is 0.2–0.5 MPaG, and the weight hourly space velocity is 0.08–1 h⁻¹. -1A wide range of hydration catalysts can be selected. The following is an illustrative example, but it does not limit the scope of the invention. For example, the hydration catalyst disclosed in CN 114409518 A can be used for this invention. The acrolein mass concentration in the feed to reactor 8 of the 3-HPA reactor is controlled to be 10% to 20%.

[0051] In the 3-HPA reactor 8, acrolein is hydrated to generate 3-HPA in the presence of an activator. The reaction to generate 3-HPA has high selectivity. In some embodiments, the activator accounts for 1‰ to 15‰ of the acrolein feed mass, for example, 1‰ to 10‰. For the present invention, the activator can be a salt, such as sodium chloride.

[0052] In some embodiments, the effluent from the 3-HPA reactor 8 is fed to the top of the acrolein recovery tower 9 to obtain unreacted acrolein, which is then returned to the feed of the 3-HPA reactor 8 as recycled acrolein and mixed with fresh acrolein. The effluent from the bottom of the acrolein recovery tower 9 is then fed into the extraction tower 10 for extraction. In some embodiments, the 3-hydroxypropionaldehyde (3-HPA) solution from the bottom of the acrolein recovery tower 9 is cooled to 40–50°C and then fed into the extraction tower 10 for extraction. The operating conditions of the extraction tower 10 can be selected over a wide range. The following is an illustrative description, but it does not limit the scope of the invention. For the present invention, the operating conditions of the extraction tower 10 include: temperature 40–50°C, pressure 0.2 MPG, and conventional structured packing material.

[0053] It is understandable that the crude product contains water after the hydration reaction. The concentration of the alcohol extractant only needs to meet the two principles of alcohol-water miscibility and material balance. In some embodiments, the alcohol extractant can be an aqueous alcohol solution with a concentration of 75-80%.

[0054] In some embodiments, the alcohol extractant includes at least one of isopropanol, n-propanol, or n-butanol.

[0055] In some embodiments, the mass ratio of alcohol extractant to 3-hydroxypropionaldehyde is 1 to 10:1.

[0056] Considering that the product of the 3-HPA reactor 8 is prone to polymerization at high temperatures, in some embodiments, the operating conditions of the acrolein recovery tower 9 include controlling the bottom temperature of the tower to not exceed 70°C under vacuum conditions. It is understood that the tower pressure is determined only to match the boiling point of the tower bottom, or the condensation point of the tower top, with the temperatures of the plant's heat and cold sources, or to avoid coking or polymerization of the medium in the tower bottom, or solidification of the medium at the top due to excessive cooling. There are no special requirements for vacuum pressure in this invention.

[0057] To prevent polymerization within the tower, in some embodiments, a polymerization inhibitor is added to the top of the acrolein recovery tower 9, preferably with a polymerization inhibitor to 3-hydroxypropionaldehyde mass ratio of 5 to 10:100.

[0058] To reduce energy consumption, in some embodiments, the heat from the compression of the gas phase at the top of the extraction phase dehydration tower 11 is used to heat the reboiler at the bottom of the tower, and part of the condensed liquid phase is returned to the top of the tower, while the other part goes to the extraction tower 10.

[0059] The operating conditions of the extraction phase dehydration tower 11 can be selected from a wide range. The following is an illustrative description, but it does not limit the scope of the present invention. For the present invention, the operating conditions of the extraction phase dehydration tower 11 include: the tower is equipped with structured packing, the tower pressure is controlled at 10 to 20 kPaA, and the tower top temperature is 43 to 50°C.

[0060] To reduce energy consumption, in some embodiments, the top discharge of the extractant recovery tower 12 is returned to the extraction tower 10 for recycling.

[0061] The operating conditions of the extractant recovery tower 12 can be selected from a wide range. The following is an illustrative description, but it does not limit the scope of the present invention. For the present invention, the operating conditions of the extractant recovery tower 12 include: the tower is equipped with structured packing, the tower pressure is controlled at 20-30 kPaA, and the tower top temperature is 55-70°C.

[0062] The 3-HPA generated by hydration using the present invention has high selectivity, hydration selectivity and hydrogenation effect, and can meet the purity requirements (≥99.9%) of "Polymer Grade 1,3-Propanediol for Fibers (T / CCFA 01028-2017)".

[0063] In some embodiments, in the oxidation section, propylene is oxidized in acrolein reactor 1 to obtain acrolein-containing material. The temperature of the effluent from acrolein reactor 1 (including acrolein and a small amount of byproducts such as acrylic acid, acetaldehyde, and acetic acid) is rapidly reduced to below 260°C. After passing through a steam generator between acrolein reactor 1 and acid washing tower 2, the temperature is further reduced to below 220°C, while high-pressure steam is generated as a byproduct. The cooled reaction product first enters acid washing tower 2, where acetic acid and acrylic acid are washed away with circulating acidic water. The product then exits from the top of acid washing tower 2 and enters the bottom of absorption tower 4 to be absorbed by low-temperature water (10-20°C) flowing down from top to bottom, generating an acrolein aqueous solution which enters desorption tower 5 to desorb acrolein and enters 3-HPA reactor 8.

[0064] The bottom material of the acid washing tower 2 includes acetic acid, acrylic acid and a small amount of acrolein washed out from the acrolein reactor 1. In order to recover this part of acrolein, it is preferable to introduce the bottom material of the acid washing tower 2 into the stripping tower 3 through the external circulation pipeline to strip and recover acrolein, return the acrolein at the top of the stripping tower 3 to the acid washing tower 2, and send the acid-containing wastewater at the bottom of the stripping tower 3 to the wastewater recovery device 7.

[0065] In some embodiments, the operating conditions of the acrolein reactor 1 include a reaction temperature of 155–300°C and a pressure of 0.2–0.3 MPa.

[0066] The operating conditions of the acid washing tower 2 can be selected from a wide range. The following is an illustrative description, but it does not limit the scope of the present invention. For the present invention, the operating conditions of the acid washing tower 2 include using a floating valve plate tower, controlling the pressure at 0.04 to 0.07 MPaG, and the tower top temperature at 80 to 85°C.

[0067] The operating conditions of stripping tower 3 can be selected from a wide range. The following is an illustrative description, but it does not limit the scope of the present invention. For the present invention, the operating conditions of stripping tower 3 include using a floating valve plate tower, controlling the pressure at 0.045 to 0.075 MPaG, and the tower top temperature at 85 to 100°C.

[0068] The operating conditions of the absorption tower 4 can be selected from a wide range. The following is an illustrative description, but it does not limit the scope of the present invention. For the present invention, the operating conditions of the absorption tower 4 include using a floating valve plate tower, controlling the pressure at 0.02 to 0.025 MPaG, and the tower top temperature at 10 to 25°C.

[0069] The operating conditions of the desorption tower 5 can be selected from a wide range. The following is an illustrative description, but it does not limit the scope of the present invention. For the present invention, the operating conditions of the desorption tower 5 include using a floating valve plate tower, controlling the pressure at 0.02 to 0.025 MPaG, and the tower top temperature at 50 to 60°C.

[0070] To prevent polymerization within the tower, in some embodiments, a polymerization inhibitor is added to the top of the desorption tower 5, with the mass ratio of the polymerization inhibitor added to the desorption tower 5 to acrolein being 1 to 20:100.

[0071] The range of polymerization inhibitors selected in this invention is relatively wide. The following is an illustrative description, but it does not limit the scope of the invention. In some embodiments, the polymerization inhibitors are selected from phenolic polymerization inhibitors and / or quinone polymerization inhibitors.

[0072] In some embodiments, the feed to the acrolein reactor is a mixture comprising propylene and air.

[0073] To control the concentration of propylene entering the acrolein reactor 1 below the explosion limit, a certain amount of low-pressure steam needs to be introduced to mix with propylene and air to reduce the feed concentration of propylene. In some embodiments, the feed to the acrolein reactor is a mixture of propylene, air and steam, wherein the molar ratio of propylene, air, steam and circulating gas from the top of the absorption tower 4 can be 1:(1.65-1.75):(0.9-1.2):(2.3-2.6), for example, 1:1.7:1:2.4. It should be noted that the above ratio is determined by plotting the reactor explosion zone, the under-air protection zone, and the catalyst moisture protection zone curves.

[0074] In some embodiments, the air and carbon oxides at the top of the absorption tower 4 are pressurized by a compressor and then partially circulated and mixed with fresh air to control the concentration of propylene in the acrolein reactor feed to the range of 10-20%. The other part enters the catalytic oxidation device 6, and after catalytic oxidation, the exhaust gas is discharged in compliance with standards. The catalytic oxidation device 6 uses a conventional commercially available precious metal catalyst.

[0075] Compared to traditional production processes, this invention uses propylene, a simpler and more readily available raw material. After oxidizing to acrolein, the acrolein is separated in four stages: acid washing, stripping, absorption, and desorption, resulting in a high acrolein yield. Furthermore, the gaseous phase after the reaction, after the above separation and washing process, undergoes catalytic oxidation to ensure that the waste gas meets emission standards. Acrylic acid, after extraction and separation in a wastewater recovery device, can directly produce acrylic acid as a byproduct. This invention satisfies both economic and environmental requirements.

[0076] The advantages of the present invention will be illustrated by the following examples, but the present invention is not limited thereto.

[0077] Example 1

[0078] Using the 1,3-propanediol preparation system shown in Figure 1, propylene, air, water vapor, and circulating gas (top gas from absorber 4) were mixed in a molar ratio of 1:1.7:1:2.4 and introduced into acrolein reactor 1 at 163°C and 0.2 MPaG. The acrolein reactor was a tubular reactor, with spherical coated composite metal catalysts packed in the tubes, and a weight hourly space velocity (WHSV) of 90 h⁻¹. -1The reactor consists of a reaction section and a cooling section. The shell side of the reaction section is filled with molten salt, and the reaction temperature (hot spot temperature 370℃) is controlled by regulating the circulation rate of the molten salt. To suppress excessive oxidation of acrolein in the reaction product, heat is removed by forced circulation using a molten salt pump in the cooling section, rapidly reducing the outlet temperature of the acrolein reactor to 255℃. The reaction produces acrolein and small amounts of acrylic acid, acetaldehyde, acetic acid, and other byproducts. Analysis of the products shows a propylene conversion rate of 98.3%, an acrolein yield of 85%, and an acrylic acid yield of 8.2%. The gas phase exiting the acrolein reactor 1 is then cooled to 220℃ by a steam generator to produce steam before entering the bottom of the acid washing tower 2. The bottom of the acid washing tower 2 is equipped with a bottom liquid circulation spray to prevent overheating of the lower tower internals. The mass ratio of the circulation spray rate to the bottom liquid phase is 1:1. A circulating acidic solution (5% wt acrylic acid, 94.5% wt water, 0.5% wt acrolein) is added to the top of the column. The pressure at the top of this column is controlled at 0.04 MPaG, and the temperature is controlled at 82℃. The liquid phase from the bottom of the column enters the stripping column 3. Saturated steam above 0.1 MPaG is added to the bottom of the stripping column. The pressure at the top of the stripping column is controlled at 0.05 MPaG, and the temperature is controlled at 97℃. The liquid phase from the bottom of the stripping column is sent to the wastewater recovery unit, and the gas from the top of the stripping column is returned to the bottom of the acid washing column 2. The gas phase from the top of the acid washing column 2 enters the bottom of the absorption column 4 and is absorbed by the low-temperature water flowing down from above, generating an acrolein aqueous solution that enters the desorption column 5. The pressure at the top of the absorption column 4 is controlled at 0.02 MPaG, and the temperature is controlled at 15℃. The mass ratio of circulating water to gas entering from the bottom of the column is 1:5. The air and carbon oxides at the top of absorption tower 4 are pressurized by a compressor, and a portion (50%) is circulated and mixed with fresh air, while the other portion (50%) enters the catalytic oxidation unit 6. After catalytic oxidation, the exhaust gas meets the emission standards. The feed temperature of the catalytic oxidation reactor is 390℃, and the volumetric space velocity is 1000 h⁻¹. -1 The acrolein content in the treated waste gas is less than 3 mg / m³. 3 Acrylic acid is less than 20 mg / m³ 3 Acetaldehyde below 50 mg / m³ 3 The emissions meet GB31571 "Emission Standard for Pollutants from Petrochemical Industry". The top pressure of desorption tower 5 is controlled at 0.02 MPaG, and the temperature is controlled at 60℃. The mass ratio of the polymerization inhibitor hydroquinone to acrolein is 8:100. The acrolein solution at the top of the tower enters the 3-HPA reactor 8, and the water collected from the bottom of the tower is cooled and recycled back to the top of absorption tower 4.

[0079] Acrylic acid-containing wastewater from the bottom of stripping tower 3 is sent to wastewater recovery unit 7. The wastewater recovery unit uses an ester extractant to extract acrylic acid from the wastewater; n-butyl acetate can be used as the ester extractant. The mass ratio of extractant to acrylic acid is 10:1. The extraction tower uses a Mellapak 250Y with a bed height of 20m. The extract phase after liquid-liquid extraction in the extraction tower enters the extractant recovery tower, where the extractant is recovered at the top and recycled. The pressure at the top of the extractant recovery tower is controlled at -0.08 MPaG, and the temperature is controlled at 37℃. Acrylic acid is recovered at the bottom of the tower with a purity of 99%. The raffinate phase enters the acrylic acid dehydration tower, where the azeotrope of extractant and water is recovered at the top, and wastewater is discharged at the bottom. The pressure at the top of the acrylic acid dehydration tower is controlled at -0.09 MPaG, and the temperature is controlled at 40℃.

[0080] Acrolein from the top of desorption tower 5 is mixed with recycled acrolein, and then mixed with a recycled aqueous solution containing sodium chloride as an activator, wherein the activator accounts for 10‰ of the acrolein feed mass. After being heated by a heater, the mixture enters 3-HPA reactor 8 for reaction, using the catalyst disclosed in Example 1 of CN 114409518 A. Acrolein and water are used to synthesize 3-HPA at a reaction temperature of 60°C, a pressure of 0.5 MPaG, and a weight hourly space velocity of 0.08 h⁻¹. -1 The product analysis showed a single-pass conversion rate of 80% and a 3-HPA selectivity of 94%. Unreacted acrolein was recovered in acrolein recovery tower 9, with a polymerization inhibitor hydroquinone to 3-hydroxypropionaldehyde mass ratio of 1:10. The top pressure of acrolein recovery tower 9 was controlled at -0.075 MPaG, and the temperature was controlled at 63℃. The acrolein collected from the top of the tower was returned to the inlet of the 3-HPA reactor 8, and the 3-HPA solution collected from the bottom of the tower was cooled to 40℃ and then entered the extraction tower 10. The extractant was n-butanol, and the extractant to 3-hydroxypropionaldehyde mass ratio was 10:1. The extraction tower used a Mellapak 250Y with a bed height of 12m. After extraction and separation, the alcohol solution rich in 3-HPA obtained from the top of the tower was dehydrated in the extraction phase dehydration tower 11, and the 3-HPA alcohol solution in the bottom of the tower was sent to the hydrogenation section. The liquid collected from the top of the tower was returned to the inlet of the extraction tower 10. The top pressure of the extraction phase dehydration tower 11 is controlled at -0.09 MPaG, and the temperature is controlled at 43℃. The extraction phase dehydration tower 11 is equipped with a heat pump energy recovery system, which can be an open-loop heat pump: after the gas phase at the top of the extraction phase dehydration tower 11 is compressed, the exhaust temperature is 84.5℃, and the pressure is -0.033 MPa. The heat contained in this exhaust heats the reboiler at the bottom of the tower. The condensed liquid phase is returned to the top of the tower to the extraction tower 10. The aqueous solution obtained at the bottom of the extraction tower 10 enters the extractant recovery tower 12 to remove alcohol, and the activator aqueous solution in the bottom of the tower is recycled. The extractant solution at the top of the tower is returned to the extraction tower 10. The top pressure of the extractant recovery tower 12 is controlled at -0.07 MPaG, and the temperature is controlled at 66℃.

[0081] An alcoholic solution rich in 3-HPA was mixed with hydrogen gas (molar ratio of 3-hydroxypropanal to hydrogen 1:45), heated to 40°C, and introduced into hydrogenation reactor 13 at a pressure of 2 MPaG, with a weight hourly space velocity of 0.6 h⁻¹. -1 The temperature rise of the hydrogenation reaction is controlled to 56°C at the outlet by partially recycling the reaction products. Unreacted 3-HPA enters the supplementary purification reactor 14 to continue the reaction, with a weight hourly space velocity of 0.5 h⁻¹. -1 Analysis of the products showed that 3-HPA achieved a conversion rate of 99.7% and a product selectivity of 99.8% after hydrogenation and supplementary purification. The supplementally purified product, after cooling and gas-liquid separation in gas-liquid separator 15, sequentially entered a dealcoholization tower 16, a propylene glycol purification tower 17, and a polypropylene glycol recovery tower 18, yielding alcohol extractant, 1,3-propanediol product, and polypropylene glycol. Specifically, the top pressure of dealcoholization tower 16 was controlled at -0.09 MPaG, and the temperature at 65℃; the top pressure of propylene glycol purification tower 17 was controlled at -0.09 MPaG, and the temperature at 150℃; and the top pressure of polypropylene glycol recovery tower 18 was controlled at -0.095 MPaG, and the temperature at 185℃. The n-butanol extractant is collected from the top of the dealcoholization column 16 and returned to the extraction column 10. The bottom liquid of the dealcoholization column 16 enters the propylene glycol purification column 17, and the top product 1,3-propanediol is collected. The bottom liquid enters the polypropylene glycol recovery column 18. After removing the heavy component residue from the bottom liquid, the polypropylene glycol collected from the top of the column is mixed with excess water and heated to 250°C and 5 MPaG. It then enters the hydrolysis reactor 19 with a weight hourly space velocity of 0.01 h⁻¹. -1 The product is then separated in dehydration tower 20 to obtain crude 1,3-propanediol, which is then recycled to propylene glycol refining tower 17 to recover 1,3-propanediol and improve the 1,3-propanediol product yield. The pressure at the top of dehydration tower 20 is controlled at -0.08 MPaG, and the temperature is controlled at 60℃. The final results show that the 1,3-propanediol product yield is 99.5%, which is higher than the 83.7% reported in the Degussa hydration unit patent. The purity of the product, after analysis and testing, meets the purity requirements (≥99.9%) of "Polymerization Grade 1,3-Propanediol for Fibers (T / CCFA 01028-2017)". Compared with the traditional hydration process without a heat pump and after hydrogenation and dehydration separation of 3-HPA aqueous solution (reported in "Research on New Technology Development of Acrolein Hydration Hydrogenation Synthesis of 1,3-Propanediol"—the comprehensive energy consumption of the hydration process is 2990.58 kg standard oil / t product), the overall energy consumption is reduced by 41%.

[0082] Example 2

[0083] Propylene, air, water vapor, and circulating gas (top gas from absorber 4) are mixed in a ratio of 1:1.7:1:2.4 and sequentially fed into acrolein reactor 1, acid washing tower 2, stripping tower 3, absorber 4, and desorption tower 5. During the process, the exhaust gas from the top of absorber 4 is pressurized by a compressor, and part of it is circulated and mixed with fresh air, while the other part enters catalytic oxidation device 6. After catalytic oxidation, the exhaust gas meets emission standards. The acrylic acid-containing wastewater from the bottom of stripping tower 3 is sent to wastewater recovery device 7 to obtain acrylic acid and wastewater with a purity of 99%. The process parameters for reactor 1, acid washing tower 2, stripping tower 3, absorber 4, and desorption tower 5 are set as in Example 1.

[0084] Acrolein from the top of desorption tower 5 is mixed with recycled acrolein, and then mixed with a recycled aqueous solution containing sodium chloride as an activator, wherein the activator accounts for 5‰ of the acrolein feed mass. After being heated by a heater, the mixture enters 3-HPA reactor 8 for reaction. The catalyst disclosed in Example 1 of CN 114409518 A is used to synthesize 3-HPA from acrolein and water. The reaction temperature is 50°C, the pressure is 0.2 MPaG, and the weight hourly space velocity is 1 h⁻¹. -1 The product analysis showed a single-pass conversion rate of 79% and a 3-HPA selectivity of 95%. Unreacted acrolein was recovered in acrolein recovery tower 9, with a polymerization inhibitor hydroquinone to 3-hydroxypropionaldehyde mass ratio of 5:100. The top pressure of acrolein recovery tower 9 was controlled at -0.075 MPaG, and the temperature was controlled at 63℃. The acrolein collected from the top of the tower was returned to the inlet of the 3-HPA reactor, and the 3-HPA solution collected from the bottom of the tower was cooled to 40℃ and then entered the extraction tower. The extractant was n-butanol, and the extractant to 3-hydroxypropionaldehyde mass ratio was 5:1. The extraction tower used a Mellapak 250Y with a bed height of 12m. After extraction and separation, the alcohol solution rich in 3-HPA obtained from the top of the tower was dehydrated in the extraction phase dehydration tower 11. The 3-HPA alcohol solution in the bottom of the tower was sent to the hydrogenation section, and the top product was returned to the inlet of the extraction tower. The top pressure of the extraction phase dehydration tower 11 was controlled at -0.09 MPaG, and the temperature was controlled at 43℃. The extraction phase dehydration tower 11 is equipped with a heat pump energy recovery system, which can be an open-loop heat pump: After the gas phase at the top of the extraction phase dehydration tower 11 is compressed, the exhaust temperature is 84.5℃ and the pressure is -0.033MPa. The heat contained in the exhaust is used to heat the reboiler at the bottom of the tower. The condensed liquid phase is returned to the top of the tower to the extraction tower 10. The aqueous solution obtained at the bottom of the extraction tower 10 enters the extractant recovery tower 12 to remove alcohol, and the activator aqueous solution at the bottom of the tower is recycled. The extractant solution at the top of the tower is returned to the extraction tower 10. The pressure at the top of the extractant recovery tower 12 is controlled at -0.07MPaG, and the temperature is controlled at 66℃.

[0085] An alcoholic solution rich in 3-HPA was mixed with hydrogen gas (molar ratio of 3-hydroxypropionaldehyde to hydrogen 1:50), heated to 37°C, and introduced into hydrogenation reactor 13 at a pressure of 2 MPaG, with a weight hourly space velocity of 0.8 h⁻¹. -1The temperature rise of the hydrogenation reaction is controlled to 53°C at the outlet by partially recycling the reaction products. Unreacted 3-HPA enters the supplementary purification reactor 14 to continue the reaction, with a weight hourly space velocity of 2.8 h⁻¹. -1 Analysis of the products showed that 3-HPA achieved a conversion rate of 99.7% and a product selectivity of 99.7% after hydrogenation and supplementary purification. The supplementally purified product, after cooling and gas-liquid separation in gas-liquid separator 15, sequentially entered a dealcoholization tower 16, a propylene glycol purification tower 17, and a polypropylene glycol recovery tower 18, yielding alcohol extractant, 1,3-propanediol product, and polypropylene glycol. Specifically, the top pressure of dealcoholization tower 16 was controlled at -0.09 MPaG, and the temperature at 65℃; the top pressure of propylene glycol purification tower 17 was controlled at -0.09 MPaG, and the temperature at 150℃; and the top pressure of polypropylene glycol recovery tower 18 was controlled at -0.095 MPaG, and the temperature at 185℃. The n-butanol extractant is collected from the top of the dealcoholization column 16 and returned to the extraction column 10. The bottom liquid of the dealcoholization column 16 enters the propylene glycol purification column 17, and the top product 1,3-propanediol is collected. The bottom liquid enters the polypropylene glycol recovery column 18. After removing the heavy component residue from the bottom liquid, the polypropylene glycol collected from the top of the column is mixed with excess water and heated to 250°C and 5 MPaG. It then enters the hydrolysis reactor 19 with a weight hourly space velocity of 1 h⁻¹. -1 The product is then separated in dehydration tower 20 to obtain crude 1,3-propanediol, which is then recycled to propylene glycol refining tower 17 to recover 1,3-propanediol and improve the 1,3-propanediol product yield. The pressure at the top of dehydration tower 20 is controlled at -0.08 MPaG, and the temperature is controlled at 60℃.

[0086] The final results show that the yield of 1,3-propanediol was 99.4%, higher than the 83.7% reported in the Degussa hydration unit patent. Analysis of the product showed that its purity met the purity requirements (≥99.9%) of the standard "Polymerization Grade 1,3-Propanediol for Fibers (T / CCFA 01028-2017)". Compared to the traditional hydration process without a heat pump and involving dehydration separation of 3-HPA aqueous solution after hydrogenation (reported in "Research and Development of New Technology for the Hydrogenation Synthesis of 1,3-Propanediol from Acrolein Hydration"—the overall energy consumption of the hydration process was 2990.58 kg standard oil / t product), the overall energy consumption is reduced by 47%.

[0087] Example 3

[0088] Unlike Example 1, the extraction tower 10, the extractant dehydration tower 11, and the extractant recovery tower 12 are not included. The discharge end of the 3-HPA reactor 8 is connected to the feed end of the hydrogenation reactor 13. That is, 3-HPA, water, and hydrogen are mixed (molar ratio of 3-hydroxypropionaldehyde to hydrogen is 1:50), heated to 37°C, and introduced into the hydrogenation reactor 13 at a pressure of 7 MPaG, with a weight hourly space velocity of 0.8 h⁻¹. -1 The rest is the same as in Example 1.

[0089] Results: The selectivity of hydrogenated products decreased by 1%, and the degree of decrease increased over time. In addition, the pressure needs to be increased to above 7 MPa, and the equipment investment is relatively large.

[0090] Comparative Example 1

[0091] Unlike Example 1, the hydrogenation section does not include a polypropylene glycol recovery tower (18), a hydrolysis reactor (19), and a dehydration tower (20), but the rest is the same as in Example 1.

[0092] Results: Compared with Example 1, the yield of 1,3-propanediol decreased by 6%, the purity was less than 99.5%, and the energy saving was reduced by 5% compared with Example 1.

[0093] Comparative Example 2

[0094] Propylene, air, water vapor, and circulating gas (top gas from absorber 4) are mixed in a ratio of 1:1.7:1:2.4 and fed into the 1,3-propanediol preparation system shown in Figure 1. The process parameters are set as in Example 1. The product purity, yield, and energy consumption of this example are the same as in Example 1. However, the heat pump energy recovery system is a closed-loop heat pump: butane is used as the heat pump organic working fluid, absorbing heat from the top gas phase condensation of the extraction phase dehydration tower 11 and vaporizing. The working fluid saturation pressure and temperature are then increased by a compressor for heating the reboiler in the tower bottom. The heated liquid working fluid is depressurized by a regulating valve and then returned to the top of the tower for heat extraction. This heat pump system uses a compressor gas flow rate that is 0.4 times that of the heat pump compressor gas flow rate in Example 1, and the investment cost of the heat pump equipment will be 0.5 times that of the heat pump equipment in Example 1.

[0095] The preferred embodiments of the present invention have been described in detail above; however, the present invention is not limited thereto. Within the scope of the inventive concept, various simple modifications can be made to the technical solutions of the present invention, including combinations of various technical features in any other suitable manner. These simple modifications and combinations should also be considered as the content disclosed in the present invention and are all within the protection scope of the present invention.

Claims

1. A system for preparing 1,3-propanediol, characterized in that, The preparation system includes a hydrogenation section, which includes a hydrogenation reactor, a propylene glycol refining tower (17), a polypropylene glycol recovery tower (18), a hydrolysis reactor (19), and a dehydration tower (20) connected in sequence. The bottom outlet of the dehydration tower (20) is connected to the inlet of the propylene glycol refining tower (17), and the propylene glycol refining tower (17) is provided with a 1,3-propanediol product outlet.

2. The preparation system according to claim 1, characterized in that, The hydrogenation reactor includes a hydrogenation reactor (13) connected in series and a supplementary refining reactor (14).

3. The preparation system according to claim 2, characterized in that, A gas-liquid separator (15) is installed on the connecting pipeline between the supplementary refining reactor (14) and the propylene glycol refining tower (17). The liquid phase outlet and gas phase outlet of the gas-liquid separator (15) are respectively connected to the feed inlet of the hydrogenation reactor (13).

4. The preparation system according to claim 1, characterized in that, The preparation system includes a hydration section, which includes a 3-HPA reactor (8) and an extraction tower (10) connected together. The top of the extraction tower (10) is connected to an extraction phase dehydration tower (11), and the bottom outlet of the extraction phase dehydration tower (11) is connected to the feed end of the hydrogenation reactor.

5. The preparation system according to claim 4, characterized in that, An acrolein recovery tower (9) is installed on the connecting pipeline between the 3-HPA reactor (8) and the extraction tower (10), and the top outlet of the acrolein recovery tower (9) is connected to the feed inlet of the 3-HPA reactor (8). And / or, the bottom outlet of the extraction tower (10) is connected to an extractant recovery tower (12), and the bottom outlet of the extractant recovery tower (12) is connected to the feed inlet of the 3-HPA reactor (8); And / or, the extraction phase dehydration tower (11) is equipped with a heat pump energy recovery system (21) for extracting heat from the top of the extraction phase dehydration tower (11) to heat the bottom of the tower; And / or, along the material flow direction, a dealcoholization tower (16) is installed after the gas-liquid separator (15), the bottom outlet of the dealcoholization tower (16) is connected to the inlet of the propylene glycol refining tower (17), and the top outlet of the dealcoholization tower (16) is connected to the bottom inlet of the extraction tower (10).

6. The preparation system according to claim 1 or 4, characterized in that, The preparation system includes an oxidation section, which includes an acrolein reactor (1), an acid washing tower (2), an absorption tower (4), and a desorption tower (5) connected in sequence. The acrolein outlet of the desorption tower (5) is connected to the inlet of the 3-HPA reactor (8).

7. The preparation system according to claim 6, characterized in that, The acid washing tower is equipped with an external circulation pipeline, and a stripping tower is installed on the external circulation pipeline; And / or, the bottom inlet of the absorption tower (4) is connected to the top outlet of the acrolein recovery tower (9); And / or, the acrolein reactor (1) is equipped with a mixing line, the feed end of which is connected to a propylene feed line and an air feed line.

8. The preparation system according to claim 7, characterized in that, The feed end of the mixing pipeline is connected to a steam feed pipeline.

9. A method for preparing 1,3-propanediol, characterized in that, The preparation method employs the preparation system described in any one of claims 1-8. The method comprises: in a hydrogenation reactor, under a hydrogen atmosphere, catalytically hydrogenating 3-hydroxypropanal feedstock to obtain a mixture stream containing 1,3-propanediol and its dimer; refining the mixture stream in a propylene glycol refining tower (17) to obtain 1,3-propanediol product and 1,3-propanediol dimer; and hydrolyzing and dehydrating the 1,3-propanediol dimer through a hydrolysis reactor (19) and a dehydration tower (20), and returning the dehydrated product to the propylene glycol refining tower (17) for further refining to obtain the 1,3-propanediol product.

10. The preparation method according to claim 9, characterized in that, The catalytic hydrogenation reaction includes: mixing 3-hydroxypropionaldehyde feedstock with hydrogen gas, heating the mixture to 30–40°C, and introducing it into the hydrogenation reactor (13) at a pressure of 1–3 MPaG, with a weight hourly space velocity of 0.6–0.8 h⁻¹. -1 Unreacted 3-hydroxypropanal enters the supplementary purification reactor (14) at a weight hourly space velocity (WHSV) of 0.5–2.8 h⁻¹. - 1 .

11. The preparation method according to claim 10, characterized in that, The effluent from the supplementary refining reactor (14) is sequentially separated and refined by a gas-liquid separator (15) and a propylene glycol refining tower (17) to obtain 1,3-propanediol product and 1,3-propanediol dimer.

12. The preparation method according to claim 11, characterized in that, The gas phase and / or liquid phase after gas-liquid separation are returned to the hydrogenation reactor (13).

13. The preparation method according to claim 9 or 11, characterized in that, The dimer of 1,3-propanediol was mixed with water and heated to 230–250°C at a pressure of 5–6 MPaG before being introduced into the hydrolysis reactor (19) at a weight hourly space velocity of 0.01–1 h⁻¹. - 1 .

14. The preparation method according to claim 10, characterized in that, In the 3-HPA reactor (8), acrolein is hydrated in the presence of an activator and then enters the extraction tower (10). The hydrated product is extracted using an alcohol extractant. The extracted phase is then dehydrated in the extraction phase dehydration tower (11) to obtain a 3-hydroxypropionaldehyde alcohol solution, which is used as the 3-hydroxypropionaldehyde raw material for the catalytic hydrogenation reaction.

15. The preparation method according to claim 14, characterized in that, The effluent from the supplementary refining reactor (14) is sequentially subjected to gas-liquid separation, de-alcoholization, and refining through a gas-liquid separator (15), a dealcoholization tower (16), and a propylene glycol refining tower (17) to obtain 1,3-propanediol product and 1,3-propanediol dimer.

16. The preparation method according to claim 15, characterized in that, The feed temperature of the 3-HPA reactor (8) is 50-60℃, the pressure is 0.2-0.5 MPaG, and the weight hourly space velocity is 0.08-1 h⁻¹. -1 ; And / or, after the acrolein is recovered by feeding the 3-HPA reactor (8) into the acrolein recovery tower (9), the recovered acrolein is returned to the 3-HPA reactor (8), and the bottom of the acrolein recovery tower (9) is fed into the extraction tower (10). And / or, the activator entering the 3-HPA reactor accounts for 1 to 10‰ of the acrolein feed mass.

17. The preparation method according to claim 14, characterized in that, The raffinate in the extraction tower (10) enters the extractant recovery tower (12) for de-alcoholization to obtain an activator aqueous solution, which is then returned to the 3-HPA reactor (8). And / or, the operating conditions of the acrolein recovery tower (9) include: controlling the bottom temperature of the tower to not exceed 70°C under vacuum conditions; And / or, a polymerization inhibitor is added to the top of the acrolein recovery tower (9); And / or, the alcohol extractant is an aqueous alcohol solution; and / or, the alcohol extractant includes at least one of isopropanol, n-propanol, or n-butanol; And / or, the mass ratio of the alcohol extractant to 3-hydroxypropionaldehyde is 1 to 10:1; And / or, propylene is oxidized in an acrolein reactor (1) to obtain acrolein-containing material, and the effluent from the acrolein reactor (1) is sequentially acid-washed, absorbed and desorbed through an acid washing tower (2), an absorption tower (4) and a desorption tower (5) to obtain the acrolein; the bottom material of the acid washing tower (2) includes acetic acid, acrylic acid and acrolein washed out from the effluent from the acrolein reactor (1).

18. The preparation method according to claim 17, characterized in that, The mass ratio of the polymerization inhibitor added to the acrolein recovery tower (9) to 3-hydroxypropionaldehyde is 5-10:100; And / or, the polymerization inhibitor is selected from phenolic polymerization inhibitors and / or quinone polymerization inhibitors; And / or, the bottom material of the acid washing tower (2) is introduced into the stripping tower (3) through the external circulation pipeline to strip and recover acrolein, and the acrolein of the stripping tower (3) is returned to the acid washing tower (2); And / or, in the absorption tower (4), the absorbent is in countercurrent contact with the overhead stream from the acid washing tower (2); And / or, a polymerization inhibitor is added to the top of the desorption tower (5), wherein the mass ratio of the polymerization inhibitor added to the desorption tower (5) to acrolein is 1 to 20:100; And / or, the feed to the acrolein reactor is a mixture comprising propylene and air.

19. The preparation method according to claim 18, characterized in that, The feed to the acrolein reactor is a mixture of propylene, air, and water vapor.

Images