CN117960046A - Reaction process for preparing succinic anhydride - Google Patents
Reaction process for preparing succinic anhydride Download PDFInfo
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- CN117960046A CN117960046A CN202211301012.9A CN202211301012A CN117960046A CN 117960046 A CN117960046 A CN 117960046A CN 202211301012 A CN202211301012 A CN 202211301012A CN 117960046 A CN117960046 A CN 117960046A
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- 238000006243 chemical reaction Methods 0.000 title claims abstract description 168
- FALRKNHUBBKYCC-UHFFFAOYSA-N 2-(chloromethyl)pyridine-3-carbonitrile Chemical compound ClCC1=NC=CC=C1C#N FALRKNHUBBKYCC-UHFFFAOYSA-N 0.000 title claims abstract description 29
- 229940014800 succinic anhydride Drugs 0.000 title claims abstract description 29
- 238000005984 hydrogenation reaction Methods 0.000 claims abstract description 200
- FPYJFEHAWHCUMM-UHFFFAOYSA-N maleic anhydride Chemical compound O=C1OC(=O)C=C1 FPYJFEHAWHCUMM-UHFFFAOYSA-N 0.000 claims abstract description 127
- 229910052739 hydrogen Inorganic materials 0.000 claims abstract description 97
- 239000001257 hydrogen Substances 0.000 claims abstract description 97
- UFHFLCQGNIYNRP-UHFFFAOYSA-N Hydrogen Chemical compound [H][H] UFHFLCQGNIYNRP-UHFFFAOYSA-N 0.000 claims abstract description 88
- 239000003054 catalyst Substances 0.000 claims abstract description 51
- 239000007795 chemical reaction product Substances 0.000 claims abstract description 42
- 239000000047 product Substances 0.000 claims abstract description 23
- 239000007788 liquid Substances 0.000 claims abstract description 20
- 239000007791 liquid phase Substances 0.000 claims abstract description 20
- 238000002156 mixing Methods 0.000 claims abstract description 13
- 238000000926 separation method Methods 0.000 claims abstract description 13
- 238000005194 fractionation Methods 0.000 claims abstract description 7
- YEJRWHAVMIAJKC-UHFFFAOYSA-N 4-Butyrolactone Chemical compound O=C1CCCO1 YEJRWHAVMIAJKC-UHFFFAOYSA-N 0.000 claims description 40
- 239000000463 material Substances 0.000 claims description 32
- 239000002904 solvent Substances 0.000 claims description 25
- CSCPPACGZOOCGX-UHFFFAOYSA-N Acetone Chemical compound CC(C)=O CSCPPACGZOOCGX-UHFFFAOYSA-N 0.000 claims description 9
- UHOVQNZJYSORNB-UHFFFAOYSA-N Benzene Chemical compound C1=CC=CC=C1 UHOVQNZJYSORNB-UHFFFAOYSA-N 0.000 claims description 6
- XEKOWRVHYACXOJ-UHFFFAOYSA-N Ethyl acetate Chemical compound CCOC(C)=O XEKOWRVHYACXOJ-UHFFFAOYSA-N 0.000 claims description 6
- YXFVVABEGXRONW-UHFFFAOYSA-N Toluene Chemical compound CC1=CC=CC=C1 YXFVVABEGXRONW-UHFFFAOYSA-N 0.000 claims description 6
- WYURNTSHIVDZCO-UHFFFAOYSA-N Tetrahydrofuran Chemical compound C1CCOC1 WYURNTSHIVDZCO-UHFFFAOYSA-N 0.000 claims description 4
- JHIVVAPYMSGYDF-UHFFFAOYSA-N cyclohexanone Chemical compound O=C1CCCCC1 JHIVVAPYMSGYDF-UHFFFAOYSA-N 0.000 claims description 4
- 230000000694 effects Effects 0.000 claims description 3
- XNWFRZJHXBZDAG-UHFFFAOYSA-N 2-METHOXYETHANOL Chemical compound COCCO XNWFRZJHXBZDAG-UHFFFAOYSA-N 0.000 claims description 2
- DKMROQRQHGEIOW-UHFFFAOYSA-N Diethyl succinate Chemical compound CCOC(=O)CCC(=O)OCC DKMROQRQHGEIOW-UHFFFAOYSA-N 0.000 claims description 2
- CTQNGGLPUBDAKN-UHFFFAOYSA-N O-Xylene Chemical compound CC1=CC=CC=C1C CTQNGGLPUBDAKN-UHFFFAOYSA-N 0.000 claims description 2
- 230000035484 reaction time Effects 0.000 claims description 2
- YLQBMQCUIZJEEH-UHFFFAOYSA-N tetrahydrofuran Natural products C=1C=COC=1 YLQBMQCUIZJEEH-UHFFFAOYSA-N 0.000 claims description 2
- 239000008096 xylene Substances 0.000 claims description 2
- 230000001105 regulatory effect Effects 0.000 abstract description 8
- 230000001360 synchronised effect Effects 0.000 abstract description 6
- 238000000605 extraction Methods 0.000 abstract description 5
- 238000002360 preparation method Methods 0.000 abstract description 4
- 230000001276 controlling effect Effects 0.000 abstract description 3
- 230000002779 inactivation Effects 0.000 abstract description 2
- 238000000034 method Methods 0.000 description 41
- 239000000243 solution Substances 0.000 description 31
- 239000002994 raw material Substances 0.000 description 19
- 150000002431 hydrogen Chemical class 0.000 description 9
- 230000009849 deactivation Effects 0.000 description 7
- 238000007086 side reaction Methods 0.000 description 7
- 239000000203 mixture Substances 0.000 description 6
- 238000005516 engineering process Methods 0.000 description 5
- 238000009472 formulation Methods 0.000 description 5
- 238000004519 manufacturing process Methods 0.000 description 5
- 230000003068 static effect Effects 0.000 description 5
- 239000000919 ceramic Substances 0.000 description 4
- 239000012295 chemical reaction liquid Substances 0.000 description 4
- 239000012528 membrane Substances 0.000 description 4
- 238000010008 shearing Methods 0.000 description 4
- 230000000052 comparative effect Effects 0.000 description 3
- 239000007789 gas Substances 0.000 description 3
- 230000002035 prolonged effect Effects 0.000 description 3
- 239000000126 substance Substances 0.000 description 3
- OKTJSMMVPCPJKN-UHFFFAOYSA-N Carbon Chemical compound [C] OKTJSMMVPCPJKN-UHFFFAOYSA-N 0.000 description 2
- PXHVJJICTQNCMI-UHFFFAOYSA-N Nickel Chemical compound [Ni] PXHVJJICTQNCMI-UHFFFAOYSA-N 0.000 description 2
- GWEVSGVZZGPLCZ-UHFFFAOYSA-N Titan oxide Chemical compound O=[Ti]=O GWEVSGVZZGPLCZ-UHFFFAOYSA-N 0.000 description 2
- YZCKVEUIGOORGS-NJFSPNSNSA-N Tritium Chemical compound [3H] YZCKVEUIGOORGS-NJFSPNSNSA-N 0.000 description 2
- PNEYBMLMFCGWSK-UHFFFAOYSA-N aluminium oxide Inorganic materials [O-2].[O-2].[O-2].[Al+3].[Al+3] PNEYBMLMFCGWSK-UHFFFAOYSA-N 0.000 description 2
- 230000003197 catalytic effect Effects 0.000 description 2
- 238000004939 coking Methods 0.000 description 2
- 229910052593 corundum Inorganic materials 0.000 description 2
- 238000007599 discharging Methods 0.000 description 2
- 239000012263 liquid product Substances 0.000 description 2
- 239000007858 starting material Substances 0.000 description 2
- KDYFGRWQOYBRFD-UHFFFAOYSA-N succinic acid Chemical compound OC(=O)CCC(O)=O KDYFGRWQOYBRFD-UHFFFAOYSA-N 0.000 description 2
- 238000006276 transfer reaction Methods 0.000 description 2
- 238000009827 uniform distribution Methods 0.000 description 2
- 229910001845 yogo sapphire Inorganic materials 0.000 description 2
- ZWEHNKRNPOVVGH-UHFFFAOYSA-N 2-Butanone Chemical compound CCC(C)=O ZWEHNKRNPOVVGH-UHFFFAOYSA-N 0.000 description 1
- YZCKVEUIGOORGS-OUBTZVSYSA-N Deuterium Chemical compound [2H] YZCKVEUIGOORGS-OUBTZVSYSA-N 0.000 description 1
- 229910004298 SiO 2 Inorganic materials 0.000 description 1
- 241000219793 Trifolium Species 0.000 description 1
- 230000003321 amplification Effects 0.000 description 1
- 238000004364 calculation method Methods 0.000 description 1
- 238000009903 catalytic hydrogenation reaction Methods 0.000 description 1
- 239000000084 colloidal system Substances 0.000 description 1
- 208000012839 conversion disease Diseases 0.000 description 1
- 230000007547 defect Effects 0.000 description 1
- 230000018044 dehydration Effects 0.000 description 1
- 238000006297 dehydration reaction Methods 0.000 description 1
- 238000010586 diagram Methods 0.000 description 1
- 239000006185 dispersion Substances 0.000 description 1
- 238000000855 fermentation Methods 0.000 description 1
- 230000004151 fermentation Effects 0.000 description 1
- 239000012530 fluid Substances 0.000 description 1
- 239000011344 liquid material Substances 0.000 description 1
- 230000014759 maintenance of location Effects 0.000 description 1
- 238000005259 measurement Methods 0.000 description 1
- 239000002184 metal Substances 0.000 description 1
- 229910052751 metal Inorganic materials 0.000 description 1
- 239000011259 mixed solution Substances 0.000 description 1
- 239000002808 molecular sieve Substances 0.000 description 1
- 239000002105 nanoparticle Substances 0.000 description 1
- 229910052759 nickel Inorganic materials 0.000 description 1
- 238000003199 nucleic acid amplification method Methods 0.000 description 1
- 239000012071 phase Substances 0.000 description 1
- 229910052573 porcelain Inorganic materials 0.000 description 1
- 239000011148 porous material Substances 0.000 description 1
- URGAHOPLAPQHLN-UHFFFAOYSA-N sodium aluminosilicate Chemical compound [Na+].[Al+3].[O-][Si]([O-])=O.[O-][Si]([O-])=O URGAHOPLAPQHLN-UHFFFAOYSA-N 0.000 description 1
- 239000001384 succinic acid Substances 0.000 description 1
Classifications
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- Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
- Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
- Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
- Y02P20/00—Technologies relating to chemical industry
- Y02P20/50—Improvements relating to the production of bulk chemicals
- Y02P20/584—Recycling of catalysts
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- Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)
Abstract
The invention discloses a reaction process for preparing succinic anhydride, which comprises the following steps: (1) Uniformly mixing maleic anhydride solution and hydrogen to obtain mixed feed I, and feeding the mixed feed I into an up-flow fixed bed reactor A to perform a first hydrogenation reaction to obtain a first hydrogenation reaction product; (2) The first hydrogenation reaction product is heated and then is uniformly mixed with supplementary hydrogen to obtain mixed feed II, and the mixed feed II enters an up-flow fixed bed reactor B to carry out a second hydrogenation reaction to obtain a second hydrogenation reaction product; (3) And (3) carrying out gas-liquid separation on the second hydrogenation reaction product after heat extraction, wherein a part of liquid phase circulates, and the other part enters a subsequent product fractionation unit. The invention can always keep higher maleic anhydride hydrogenation conversion rate and selectivity in the succinic anhydride preparation process by regulating and controlling the hydrogen supply ratio and the microbubble content of the reactor A/B, and simultaneously can realize synchronous inactivation of all catalysts in a reaction system, greatly prolong the total operation period of the device and improve the economical efficiency of industrial devices.
Description
Technical Field
The invention belongs to the technical field of degradable material production, and particularly relates to a reaction process for preparing succinic anhydride.
Background
At present, the production method of succinic anhydride is mainly divided into a succinic anhydride dehydration method, a biological fermentation method and a maleic anhydride catalytic hydrogenation method, wherein the maleic anhydride hydrogenation method is the method with the highest conversion rate of succinic anhydride production and the highest product, and is most suitable for large-scale industrialization, but the maleic anhydride hydrogenation production of succinic anhydride is not industrialized in a large quantity, and the following problems mainly exist at present: (1) The reaction for preparing succinic anhydride by maleic anhydride hydrogenation is a strong exothermic reaction (delta H=128 kJ/mol), and the conventional trickle bed hydrogenation and the conventional liquid phase hydrogenation can not timely remove the reaction heat in time, so that the temperature in the reaction process can not be controlled, and the problems of local hot spots of a catalyst bed, serious side reactions, coking and hardening of the catalyst and the like are caused; (2) The maleic anhydride hydrogenation reaction is based on low maleic anhydride concentration in the later reaction period, so that the contact mass transfer rate of hydrogen and maleic anhydride is very low, the conversion rate of more than 99% is achieved, very long residence time is required, and side reactions are increased; (3) The general maleic anhydride hydrogenation reaction adopts two stages of reactors connected in series, so that the reaction rate of a first reactor, the reaction rate of a second reactor, the temperature rise and the like are different, the catalysts of the two reactors cannot be deactivated synchronously, and the total operation period and the economy of the device are affected.
CN103570650a proposes a technological process for continuously producing succinic anhydride and co-producing succinic acid by maleic anhydride hydrogenation, the method adopts a two-stage hydrogenation reactor, the first-stage hydrogenation reactor is a fixed bed reactor for feeding hydrogen and reaction liquid downwards and discharging upwards, the second-stage hydrogenation reactor is a trickle bed reactor for feeding hydrogen and reaction liquid upwards and discharging downwards, and an external circulation heat removal mode is adopted to remove reaction heat, so as to control the average operation temperature of the whole reactor and equalize the temperature in the reactor. In the method, a primary reactor adopts a parallel flow upward flow mode of hydrogen and reaction liquid, and based on the specificity of large heat release of maleic anhydride hydrogenation reaction, the conventional technology cannot ensure uniform material mixing and uniform distribution, and cannot ensure uniform reaction and solve the problem of local hot spots; the secondary reactor adopts a parallel-flow downward trickle bed reactor flow mode, so that the timely taking away of the reaction heat can not be ensured, and the problem of local hot spots can be solved.
CN 105801536B proposes a method for preparing succinic anhydride by maleic anhydride liquid phase selective hydrogenation, the liquid phase hydrogenation reaction adopts a two-stage low-temperature low-pressure reaction process method to prepare succinic anhydride, two reactors are adopted, a first-stage reactor and a second-stage reactor are respectively adopted, and the first-stage reactor and the second-stage reactor are used in series; the maleic anhydride, the solvent and the hydrogen enter a first-stage reactor to carry out partial catalytic selective hydrogenation, after the reaction, the residual maleic anhydride, the generated succinic anhydride and the solvent mixed liquid material enter a second-stage reactor to carry out complete catalytic selective hydrogenation, and the succinic anhydride product is obtained after gas-liquid separation and rectification of the product of the second-stage reactor. In the method, the two-stage reactor adopts a liquid-phase hydrogenation method of hydrogen and reaction liquid, and based on the specificity of large heat release of maleic anhydride hydrogenation reaction, the conventional liquid-phase hydrogenation mixing and reaction technology cannot ensure uniform material mixing and uniform distribution, and cannot ensure uniform reaction and solve the problem of local hot spots.
In summary, the existing maleic anhydride hydrogenation technology for preparing succinic anhydride has few related patents, and most of the related patents are concentrated in the technological processes of catalyst preparation and reaction, so that the method has important significance on how to solve the problem of concentrated exothermic reaction, ensure higher conversion rate and selectivity, synchronous inactivation of the catalyst and the like in the maleic anhydride hydrogenation reaction process, and has long-period and high-efficiency operation for the succinic anhydride production device.
Disclosure of Invention
Aiming at the defects of the prior art, the invention discloses a reaction process for preparing succinic anhydride. The invention can control the hydrogen supply ratio (K) and the microbubble content of the upflow fixed bed series reactor A/B in the maleic anhydride hydrogenation reaction process, and deeply couple with the earlier stage/later stage of the maleic anhydride hydrogenation reaction and the whole reaction process, so that the higher maleic anhydride hydrogenation conversion rate and selectivity can be kept all the time in the succinic anhydride preparation process, and simultaneously, all the catalysts in the reaction system can be synchronously deactivated, the total operation period of the device is greatly prolonged, and the economical efficiency of the industrial device is improved.
The reaction process for preparing succinic anhydride of the invention comprises the following steps:
(1) Uniformly mixing maleic anhydride solution and hydrogen to obtain mixed feed I, and feeding the mixed feed I into an up-flow fixed bed reactor A to perform a first hydrogenation reaction to obtain a first hydrogenation reaction product;
(2) The first hydrogenation reaction product is heated and then is uniformly mixed with supplementary hydrogen to obtain mixed feed II, and the mixed feed II enters an up-flow fixed bed reactor B to carry out a second hydrogenation reaction to obtain a second hydrogenation reaction product;
(3) After the second hydrogenation reaction product is heated, gas-liquid separation is carried out, one part of liquid phase circulates, and the other part enters a subsequent product fractionation unit; wherein the ratio K of the flow rate (Nm 3/h) of the hydrogen of the upflow fixed bed reactor A in the step (1) to the flow rate (Nm 3/h) of the supplementary hydrogen of the upflow fixed bed reactor B in the step (2) is 1: 10-10: 1.
In the method, the maleic anhydride content in the maleic anhydride solution is 0.03-0.3 g/mL, preferably 0.05-0.15 g/mL, and the solvent adopted in the maleic anhydride solution is one or more of benzene, toluene, xylene, acetone, tetrahydrofuran, gamma-butyrolactone, methyl-acetone, cyclohexanone, ethyl acetate, diethyl succinate or ethylene glycol monomethyl ether and the like.
In the process of the invention, hydrogen may generally be used in a purity of more than 90 (v)%, preferably 99.9% pure hydrogen.
In the method of the invention, in the mixed material I, hydrogen is uniformly dispersed in maleic anhydride solution, hydrogen bubbles mainly have millimeter-level size (the diameter d1 is generally 0.01-20 mm), namely, more than or equal to 70% of hydrogen bubbles are millimeter-level, and the rest is nanometer or micrometer-level; the mixing process may generally employ one or more combinations of static mixers, jet mixers, mechanical shear mixers, impingement mixers, and the like. The process based on the upflow fixed bed reactor A is in the early stage of reaction, the concentration of maleic anhydride is relatively high, the chemical hydrogen consumption and mass transfer driving force are relatively high, and when hydrogen is introduced into the reactor for hydrogenation reaction, the hydrogen is broken into bubbles mainly with millimeter-sized bubbles, so that the sufficiently high reaction rate can be ensured. The method for determining the size and content of bubbles used herein is as follows: and (3) introducing a mixed solution of maleic anhydride solution and hydrogen into a transparent reactor according to a certain proportion, adopting any region in a micro-fluid high-frequency camera shooting device under a stable flow state, measuring the size of bubbles in the shooting range after amplification, calculating the proportion of various types of bubbles, and obtaining the average value of the multi-region photo as a final calculation result. The early stage of the reaction refers to the stage in which the concentration of maleic anhydride in the maleic anhydride solution material in the reactor is greater than or equal to the concentration of succinic anhydride, and the later stage of the reaction refers to the stage in which the concentration of maleic anhydride in the maleic anhydride solution material in the reactor is less than the concentration of succinic anhydride.
In the method of the invention, the first hydrogenation reaction conditions are as follows: the reaction temperature is 40-200 ℃, preferably 50-150 ℃; the reaction pressure is 0.5-10.0 MPa, preferably 1-5.0 MPa; the liquid hourly space velocity is 0.5-15.0 h -1, preferably 3.0-8.0 h -1; the ratio of the volume flow rate of hydrogen (Nm 3/h) to the volume flow rate of maleic anhydride solution (m 3/h) in the upflow fixed bed reactor A was 5:1 to 100:1, preferably 10:1 to 60:1.
In the method, in the mixed material II, hydrogen is uniformly mixed and dispersed in a first hydrogenation product, and equipment capable of generating a large amount of micro bubbles is adopted in the mixing process, for example, one or a plurality of combinations of a dissolved air pump, a colloid mill, a micro-pore plate nano/micro hydrogen dispersing component, a micro bubble generator, a ceramic membrane nano/micro hydrogen dispersing component, a metal membrane nano/micro hydrogen dispersing component, a micro channel mixer, a venturi ejector, high-speed shearing equipment and the like can be adopted; the hydrogen bubbles can be all nano-micron size (diameter d2 is generally 10 nm-1000 μm) or mainly nano-micron size (diameter d2 is generally 10 nm-1000 μm), namely, more than or equal to 70% of the hydrogen bubbles are nano-micron size, and the rest is millimeter size. Here, the process based on the upflow fixed bed reactor B is in the later stage of reaction, the maleic anhydride concentration is relatively low, the chemical hydrogen consumption and the mass transfer driving force are relatively low, when hydrogen is introduced into the reactor to carry out hydrogenation reaction, the hydrogen is required to be broken into bubbles mainly with nano-micron size, the gas-liquid contact area is greatly increased, the mass transfer in the reaction process is enhanced, and the higher mass transfer reaction rate, the high conversion rate and the high selectivity can be ensured. In the method of the invention, the second hydrogenation reaction conditions are as follows: the reaction temperature is 40-200 ℃, preferably 50-150 ℃; the reaction pressure is 0.5-10.0 MPa, preferably 1-5.0 MPa; the liquid hourly space velocity is 0.1 to 8.0h -1, preferably 0.5 to 3.0h -1; the ratio of the volume flow rate of hydrogen (Nm 3/h) in the upflow fixed bed reactor B to the volume flow rate of maleic anhydride feed (m 3/h) in the upflow fixed bed reactor A was 1:1 to 60:1, preferably 5: 1-40: 1.
In the method of the invention, 1 or more catalyst beds can be arranged in the upflow fixed bed reactors A and B according to the requirement, the catalyst can be a catalyst with a hydrogenation function commonly used in maleic anhydride hydrogenation reaction in the field, preferably a supported nickel-based catalyst, wherein the catalyst carrier can be one or more of SiO 2、Al2O3、SiO2-Al2O3、TiO2, activated carbon or molecular sieve, and the catalyst can be one of sphere, bar, clover or tooth sphere, and the like, preferably a sphere or tooth sphere catalyst.
In the method, the maleic anhydride conversion rate of the first hydrogenation reaction is generally 41-99%, the maleic anhydride conversion rate at the initial stage of the reaction is generally 71-99%, and the maleic anhydride conversion rate at the final stage of the reaction is generally 41-70%; the maleic anhydride conversion rate of the second hydrogenation reaction is generally 1-40%, the maleic anhydride conversion rate at the initial stage of the reaction is generally 1-20%, and the maleic anhydride conversion rate at the final stage of the reaction is generally 21-40%. The initial stage and the final stage of the reaction are divided according to the total operation period of the catalyst, the initial stage of the reaction refers to the stage that the inlet temperature of the reactor meets the requirement of maleic anhydride conversion, the final stage of the reaction refers to the stage that the inlet temperature of the reactor does not meet the requirement of maleic anhydride conversion along with the extension of the reaction time and the reduction of the activity of the catalyst, the inlet temperature of the reactor needs to be increased to meet the requirement of maleic anhydride conversion, and the total operation period of the catalyst is reached after the inlet temperature of the reactor reaches the upper limit.
In the total operation period of the catalyst, on one hand, the maleic anhydride conversion rate in the reactor A is gradually reduced and the maleic anhydride conversion rate in the second reactor B is gradually increased by regulating and controlling the hydrogen ratio (K) in the reactor A and the reactor B to gradually change, the hydrogenation conversion rate is transferred from the first reactor to the second reactor, the reaction conversion rate and the catalyst activity of the two reactors are balanced, and the catalysts of the two reactors reach the end of the reaction and are synchronously deactivated; meanwhile, as the conversion rate of the maleic anhydride hydrogenation reaction is transferred, the retention time of materials in the reactor B is gradually shortened, and in order to ensure the conversion rate of the reaction, when the feeding amount of hydrogen in the reactor B is gradually increased, namely the proportion of the hydrogen to the total hydrogen is gradually increased due to the fact that the hydrogen is crushed into nano-micro bubbles, the mass transfer reaction rate and the conversion rate of the second hydrogenation reaction are greatly enhanced and ensured.
In the method of the invention, the first circulating material which circulates back to the upflow fixed bed reactor A accounts for 15 to 90 weight percent, preferably 30 to 80 weight percent of the fresh feed (maleic anhydride solution) of the upflow fixed bed reactor A; the second recycle to the upflow reactor B represents from 0 to 80% by weight, preferably from 0 to 40% by weight, of the fresh feed to the upflow reactor B.
The existing reaction process for preparing succinic anhydride mainly has the following problems: (1) The reaction for preparing succinic anhydride by maleic anhydride hydrogenation is a strong exothermic reaction, and the conventional trickle bed hydrogenation and the conventional liquid phase hydrogenation can not timely remove the reaction heat in time, so that the problems of local hot spot of a catalyst bed, serious side reaction, coking and hardening of the catalyst and the like are easily caused; (2) The maleic anhydride hydrogenation reaction is based on low maleic anhydride concentration in the later reaction period, so that the contact mass transfer rate of hydrogen and maleic anhydride is very low, particularly, the conversion rate of more than 99% is achieved, and the long residence time is required, which is also an important reason for increasing side reaction and low selectivity; (3) The general maleic anhydride hydrogenation reaction adopts a two-stage hydrogenation reactor series process, so that the reaction rate of a first reactor and the reaction rate of a second reactor, the catalyst deactivation rate and the like are different, the catalysts of the two reactors cannot be deactivated synchronously, and the total operation period and the economy of the device are affected.
The system of the invention is coupled with the front-stage/rear-stage maleic anhydride hydrogenation reaction and the depth of the whole reaction process by regulating and controlling the hydrogen supply ratio (K) and the microbubble content of the upflow fixed bed serial reactor A/B in the maleic anhydride hydrogenation reaction process, so that the higher maleic anhydride hydrogenation conversion rate (more than or equal to 99.8%) and selectivity (more than or equal to 98.5%) are always maintained in the succinic anhydride preparation process, and simultaneously, all catalysts in the reaction system can be synchronously deactivated, the total operation period of the device is greatly prolonged, and the economical efficiency of the industrial device is improved. Here, based on the problem that the reaction rate difference between the front-stage and the rear-stage of the maleic anhydride hydrogenation reaction is relatively high, the reaction rate is relatively fast, the reaction mass transfer driving force is relatively large, and the front-stage maleic anhydride concentration is relatively low, the reaction rate is relatively slow, the reaction mass transfer driving force is relatively small, namely, on one hand, the residence time required by the front-stage and the rear-stage reaction is different, the residence time required by the rear-stage reaction to reach more than or equal to 99% conversion rate is long, so that the side reaction is more and the selectivity is low; on the other hand, the catalyst deactivation rates in the early and late stages of the reaction are inconsistent, i.e., the catalyst deactivation in the reaction system is not synchronized. The invention sets two-stage serial continuous hydrogenation reactor, which makes the hydrogenation reaction concentrated mainly in the upflow hydrogenation reactor A in the early stage of reaction, and transfers to the upflow hydrogenation reactor B along with the hydrogenation reaction, until the catalyst of the two reactors reach synchronous deactivation, in this process, the transfer of hydrogenation reaction is controlled by hydrogen feeding, and the microbubble proportion is controlled to achieve the purpose of high conversion rate and selectivity of hydrogenation reaction.
Drawings
FIG. 1 is a schematic diagram of a reaction process for preparing succinic anhydride according to the present invention.
Wherein 1 is maleic anhydride solution, 2 is first hydrogen, 3 is supplementary hydrogen, 4 is mixer A,5 is upflow fixed bed reactor A,6 is porcelain ball, 7 is catalyst bed, 8 is first hydrogenation reaction product, 9 is heat-collecting device A,10 is liquid feed of upflow fixed bed reactor B, 11 is upflow fixed bed reactor B,12 is mixer B,13 is catalyst bed, 14 is second hydrogenation reaction product, 15 is heat-collecting device B,16 is gas-liquid separator feed, 17 is gas-liquid separator, 18 is separated gas, 19 is separated liquid phase hydrogenation product, 20 is liquid phase product entering fractionation unit, 21 is circulating pump, 22 is circulating material entering upflow fixed bed reactor B, 23 is circulating material entering upflow fixed bed reactor A.
Detailed Description
The invention will now be described in more detail with reference to the accompanying drawings and examples, which are not intended to limit the invention thereto.
Taking the attached figure 1 as an example, the application process of the reaction process for preparing succinic anhydride of the invention is as follows: firstly, maleic anhydride solution 1, circulating material 23 and first hydrogen 2 are uniformly mixed by a mixer A4 to obtain mixed feed I, the mixed feed I enters an up-flow fixed bed hydrogenation reactor A5, ceramic balls 6 are uniformly distributed again and then are downwards and upwards in a catalyst bed 7 to perform up-flow hydrogenation reaction, a first hydrogenation reaction product 9 is obtained, after leaving the up-flow fixed bed hydrogenation reactor A5, heat is taken by a heat taking device A9, then the mixed feed I is uniformly mixed with circulating material 22 and supplementary hydrogen 3, the mixed feed II is obtained by firstly uniformly mixing the mixed feed II with the circulating material with supplementary hydrogen 3, meanwhile, the supplementary hydrogen is broken into nano/micro bubbles, the nano/micro bubbles enter the catalyst bed 13 and are upwards subjected to up-flow hydrogenation reaction, a first hydrogenation reaction product 14 is obtained, after leaving the up-flow fixed bed hydrogenation reactor B11, heat is taken by a heat taking device B15, then the first hydrogenation reaction product 9 enters a gas-liquid separator 17, gas 18 after gas-liquid separation is led out of a reaction system, the separated liquid product 18 is separated into two paths, and one path enters a subsequent separation unit, and the other path of the separated liquid product 18 is circulated by a circulating pump 21 to the reaction system.
The method is applied to the technological process of preparing succinic anhydride by maleic anhydride hydrogenation. Maleic anhydride starting material and gamma-butyrolactone solvent were commercially available, the specific properties are shown in tables 1 and 2, respectively, and the catalyst properties are shown in table 3.
TABLE 1 maleic anhydride raw material Properties
TABLE 2 solvent Properties of gamma butyrolactone
TABLE 3 physical and chemical indicators of catalyst
Comparative example 1
The conventional fixed bed hydrogenation process is adopted, and maleic anhydride is subjected to maleic anhydride hydrogenation reaction in the first hydrogenation reactor A and the second hydrogenation reactor B in sequence in a mode of connecting an up-flow fixed bed hydrogenation reactor A and a down-flow fixed bed hydrogenation reactor B in series. Firstly, maleic anhydride raw materials are dissolved in gamma-butyrolactone solvent and uniformly mixed to prepare maleic anhydride solution, the maleic anhydride solution is regulated to the temperature of an inlet of a reactor and then mixed with hydrogen, the maleic anhydride solution enters from the bottom of an up-flow hydrogenation reactor A, hydrogenation reaction is carried out through a catalyst bed layer from bottom to top, the obtained hydrogenation product enters from the top of a down-flow hydrogenation reactor after being regulated to be mixed with supplementary hydrogen, hydrogenation reaction is carried out through the catalyst bed layer from top to bottom, the maleic anhydride solution leaves the reactor after the hydrogenation reaction is completed, gas-liquid separation is carried out through a separator, and separated materials are partially circulated, and the other part of the material enters a separation unit.
The operating conditions of the first hydrogenation reactor a were as follows:
The reaction temperature is 50-150 ℃;
the reaction pressure is 6.0-6.5 MPaG;
The height-diameter ratio of the reactor is as follows: 2.5
Volume space velocity: 3.0h -1
Maleic anhydride formulation concentration: 12g/mL
The volume ratio of hydrogen (Nm 3/h) to fresh raw material (m 3/h) (solution of maleic anhydride dissolved in gamma-butyrolactone solvent) was 300:1, a step of;
the mass ratio of the recycle amount of the reaction product entering the primary reaction to the fresh raw materials: 30%;
the operating conditions of the second hydrogenation reactor B are as follows:
The reaction temperature is 50-150 ℃;
the reaction pressure is 6.0-6.5 MPaG;
volume space velocity: 1.0h -1;
the height-diameter ratio of the reactor is as follows: 2.5;
note that: the heat of the first effluent is taken to enter the second reactor, and the second reactor inlet is not supplemented with hydrogen and is not reintroduced into the circulating material.
Under the reaction conditions, maleic anhydride and gamma-butyrolactone solvents shown in table 1 and table 2 are taken as raw materials, and enter a first hydrogenation reactor A and a second hydrogenation reactor B to continuously carry out hydrogenation reaction to obtain hydrogenation products, and the reaction results are shown in table 4.
Comparative example 2
The conventional fixed bed hydrogenation process is adopted, and maleic anhydride hydrogenation reaction is sequentially carried out in a first hydrogenation reactor A and a second hydrogenation reactor B by adopting a mode that two upflow fixed bed hydrogenation reactors A and a upflow fixed bed hydrogenation reactor B are connected in series. Firstly, maleic anhydride raw materials are dissolved in gamma-butyrolactone solvent and uniformly mixed to prepare maleic anhydride solution, the maleic anhydride solution is regulated to the temperature of an inlet of a reactor and then mixed with hydrogen, the maleic anhydride solution enters from the bottom of an up-flow hydrogenation reactor A, hydrogenation reaction is carried out through a catalyst bed layer from bottom to top, the obtained hydrogenation product enters from the bottom of an up-flow hydrogenation reactor B after being regulated to be mixed with supplementary hydrogen, hydrogenation reaction is carried out through the catalyst bed layer from bottom to top, the maleic anhydride solution leaves the reactor after the hydrogenation reaction is finished, gas-liquid separation is carried out through a separator, part of separated liquid phase materials circulate, the other part of separated liquid phase materials enter a separation unit, and the separated gas phase materials are compressed by a circulating hydrogen compressor and then recycled.
The operating conditions of the first hydrogenation reactor a were as follows:
The reaction temperature is 50-150 ℃;
the reaction pressure is 6.0-6.5 MPaG;
The height-diameter ratio of the reactor is as follows: 2.5
Volume space velocity: 3.0h -1
Maleic anhydride formulation concentration: 12g/mL
The volume ratio of hydrogen (Nm 3/h) to fresh raw material (m 3/h) (solution of maleic anhydride dissolved in gamma-butyrolactone solvent) was 100:1, a step of;
the mass ratio of the recycle amount of the reaction product entering the primary reaction to the fresh raw materials: 30%;
the operating conditions of the second hydrogenation reactor B are as follows:
The reaction temperature is 50-150 ℃;
the reaction pressure is 6.0-6.5 MPaG;
volume space velocity: 1.0h -1;
the height-diameter ratio of the reactor is as follows: 2.5;
The volume ratio of make-up hydrogen (Nm 3/h) to fresh starting material (m 3/h) (solution of maleic anhydride dissolved in gamma-butyrolactone solvent) was 400:1, a step of;
note that: the heat of the first effluent is taken to enter the second reactor, and the second reactor inlet is not used for introducing circulating materials.
Under the reaction conditions, maleic anhydride and gamma-butyrolactone solvents shown in table 1 and table 2 are taken as raw materials, and enter a first hydrogenation reactor A and a second hydrogenation reactor B to continuously carry out hydrogenation reaction to obtain hydrogenation products, and the reaction results are shown in table 4.
Example 1
By adopting the process, the maleic anhydride hydrogenation reaction zone is provided with two upflow fixed bed hydrogenation reactors A and B. Firstly, uniformly mixing a pre-prepared 15% maleic anhydride (gamma-butyrolactone solvent) solution and hydrogen by adopting a conventional SV type static mixer to form a mixed feed I, and feeding the mixed feed I into an up-flow fixed bed hydrogenation reactor A for a first hydrogenation reaction to obtain a first hydrogenation reaction product; the first hydrogenation reaction product is heated and then is uniformly mixed with supplementary hydrogen to obtain mixed feed II, and the mixed feed II enters an up-flow fixed bed hydrogenation reactor B to carry out a second hydrogenation reaction to obtain a second hydrogenation reaction product; the first hydrogenation reaction product is mixed with the supplementary hydrogen by a Venturi mixer, namely, the hydrogen is dispersed into nano-micro bubbles and then mixed with the first hydrogenation reaction product, so that a liquid phase material containing the nano-micro bubbles is formed and enters an up-flow fixed bed hydrogenation reactor B. And (3) carrying out gas-liquid separation on the second hydrogenation reaction product after heat extraction, wherein a part of liquid phase circulates, and the other part enters a subsequent product fractionation unit.
The operating conditions for upflow fixed bed hydrogenation reactor A were as follows:
The reaction temperature is 50-150 ℃;
The reaction pressure is 4.0-4.5 MPaG;
The height-diameter ratio of the reactor is as follows: 2.5
Volume space velocity: 6.0h -1
Maleic anhydride formulation concentration: 15g/mL
The volume ratio of hydrogen (Nm 3/h) to fresh raw material (m 3/h) (solution of maleic anhydride dissolved in gamma-butyrolactone solvent) was 60:1, a step of;
the mass ratio of the circulation amount of the reaction product entering the reactor A to the fresh raw materials: 30%;
In the upflow fixed bed hydrogenation reactor A, maleic anhydride solution and hydrogen are mixed by adopting a conventional SV type static mixer, and the sizes of bubbles in the mixed materials of the mixer are measured to be all in the range of 1-20 mm in diameter.
The operating conditions of the upflow fixed bed hydrogenation reactor B are as follows:
The reaction temperature is 50-150 ℃;
The reaction pressure is 4.0-4.5 MPaG;
The height-diameter ratio of the reactor is as follows: 2.5
Volume space velocity: 1.5h -1
The ratio of the volume flow of make-up hydrogen (Nm 3/h) to the fresh feed (m 3/h) (sum of maleic anhydride and solvent) in the upflow fixed bed reactor A was 6:1.
The mass ratio of the circulation amount of the reaction product entering the reactor B to the fresh raw materials: 20% of a base;
In the upflow fixed bed hydrogenation reactor B, the first hydrogenation product and the supplementary hydrogen are mixed in a Venturi mixer, and the mixture mixed by the mixer has a bubble size ranging from 50 mu m to 1000 mu m, wherein the rest about 28% of bubbles are in millimeter level.
The ratio K of the hydrogen supply to the upflow fixed-bed reactor A to the flow of make-up hydrogen supply to the upflow fixed-bed reactor B (Nm 3/h) was 10:1.
Under the reaction conditions, maleic anhydride and gamma-butyrolactone solvents shown in table 1 and table 2 are taken as raw materials, and enter an up-flow fixed bed hydrogenation reactor A and an up-flow fixed bed hydrogenation reactor B to continuously carry out hydrogenation reaction to obtain hydrogenation products, and the reaction results are shown in table 4.
Example 2
By adopting the method, two upflow fixed bed hydrogenation reactors A and B are arranged in the maleic anhydride hydrogenation reaction zone. Firstly, uniformly mixing a pre-prepared 15% maleic anhydride (gamma-butyrolactone solvent) solution and hydrogen by adopting a conventional SV type static mixer to form a mixed feed I, and feeding the mixed feed I into an up-flow fixed bed hydrogenation reactor A for a first hydrogenation reaction to obtain a first hydrogenation reaction product; the first hydrogenation reaction product is heated and then is uniformly mixed with supplementary hydrogen to obtain mixed feed II, and the mixed feed II enters an up-flow fixed bed hydrogenation reactor B to carry out a second hydrogenation reaction to obtain a second hydrogenation reaction product; the first hydrogenation reaction product is mixed with supplementary hydrogen by a rotational flow shearing type bubble generator, and the formed liquid phase material containing nano-sized and micro-sized bubbles enters an up-flow fixed bed hydrogenation reactor B. And (3) carrying out gas-liquid separation on the second hydrogenation reaction product after heat extraction, wherein a part of liquid phase circulates, and the other part enters a subsequent product fractionation unit.
The operating conditions for upflow fixed bed hydrogenation reactor A were as follows:
The reaction temperature is 50-150 ℃;
The reaction pressure is 4.0-4.5 MPaG;
The height-diameter ratio of the reactor is as follows: 4.0
Volume space velocity: 5.0h -1
Maleic anhydride formulation concentration: 15g/mL
The volume ratio of hydrogen (Nm 3/h) to fresh feed (m 3/h) (solution of maleic anhydride in gamma-butyrolactone solvent) was 30:1, a step of;
the mass ratio of the circulation amount of the reaction product entering the reactor A to the fresh raw materials: 40%;
In the upflow fixed bed hydrogenation reactor A, maleic anhydride solution and hydrogen are mixed in a conventional SV type static mixer, and the size of bubbles in the mixed material is measured to be in the range of 1-20 mm in diameter.
The operating conditions of the upflow fixed bed hydrogenation reactor B are as follows:
The reaction temperature is 50-150 ℃;
The reaction pressure is 4.0-4.5 MPaG;
The height-diameter ratio of the reactor is as follows: 2.0
Volume space velocity: 1.7h -1
The ratio of the volume flow of make-up hydrogen (Nm 3/h) to the fresh feed (m 3/h) (sum of maleic anhydride and solvent) in the upflow fixed bed reactor A was 30:1.
The mass ratio of the circulation amount of the reaction product entering the reactor B to the fresh raw materials: 25%;
In the upflow fixed bed hydrogenation reactor B, the first hydrogenation product and the supplementary hydrogen are mixed in gas-liquid mode by a rotational flow shearing type bubble generator based on the shearing principle, and the measurement shows that about 86% of the total bubbles in the mixed material have the bubble size range of 50-1000 μm in diameter, and the rest about 14% of the bubbles are in millimeter level.
The ratio K of the hydrogen supply to the upflow fixed-bed reactor A to the flow of make-up hydrogen supply to the upflow fixed-bed reactor B (Nm 3/h) was 1:1.
Under the reaction conditions, maleic anhydride and gamma-butyrolactone solvents shown in table 1 and table 2 are taken as raw materials, and enter an up-flow fixed bed hydrogenation reactor A and an up-flow fixed bed hydrogenation reactor B to continuously carry out hydrogenation reaction to obtain hydrogenation products, and the reaction results are shown in table 4.
Example 3
By adopting the method, two upflow fixed bed hydrogenation reactors A and B are arranged in the maleic anhydride hydrogenation reaction zone. Firstly, uniformly mixing a pre-prepared 15% maleic anhydride (gamma-butyrolactone solvent) solution and hydrogen to form a mixed feed I, and feeding the mixed feed I into an up-flow fixed bed hydrogenation reactor A for a first hydrogenation reaction to obtain a first hydrogenation reaction product; uniformly mixing the first hydrogenation reaction product after heat extraction with supplementary hydrogen by a Venturi mixer to obtain mixed feed II, and feeding the mixed feed II into an up-flow fixed bed reactor B to perform a second hydrogenation reaction to obtain a second hydrogenation reaction product; the first hydrogenation reaction product is mixed with the supplementary hydrogen by adopting a ceramic membrane mixer, namely, the hydrogen is dispersed into nano-micro bubbles and then mixed with the first hydrogenation reaction product, so that a liquid phase material containing the nano-micro bubbles is formed and enters an up-flow fixed bed hydrogenation reactor B. And (3) carrying out gas-liquid separation on the second hydrogenation reaction product after heat extraction, wherein a part of liquid phase circulates, and the other part enters a subsequent product fractionation unit.
The operating conditions for upflow fixed bed hydrogenation reactor A were as follows:
The reaction temperature is 50-150 ℃;
The reaction pressure is 4.0-4.5 MPaG;
The height-diameter ratio of the reactor is as follows: 6.0
Volume space velocity: 4.0h -1
Maleic anhydride formulation concentration: 15g/mL
The volume ratio of hydrogen (Nm 3/h) to fresh feed (m 3/h) (solution of maleic anhydride in gamma-butyrolactone solvent) was 25:1, a step of;
the mass ratio of the circulation amount of the reaction product entering the reactor A to the fresh raw materials: 30%;
In the upflow fixed bed hydrogenation reactor A, maleic anhydride solution and hydrogen are mixed by a Venturi ejector, the size of bubbles in the mixed material is 87% in the range of 1-20 mm in diameter, and the rest 13% are micron-sized bubbles.
The operating conditions of the upflow fixed bed hydrogenation reactor B are as follows:
The reaction temperature is 50-150 ℃;
The reaction pressure is 4.0-4.5 MPaG;
The height-diameter ratio of the reactor is as follows: 2.0
Volume space velocity: 1.8h -1
The ratio of the volume flow of make-up hydrogen (Nm 3/h) to the fresh feed (m 3/h) (sum of maleic anhydride and solvent) in the upflow fixed bed reactor A was 50:1.
The mass ratio of the circulation amount of the reaction product entering the reactor B to the fresh raw materials: 30%;
in the upflow fixed bed hydrogenation reactor B, the first hydrogenation product and the supplementary hydrogen are mixed in a ceramic membrane mixer based on the micropore dispersion principle, and about 95% of the mixed materials in the mixer have bubble sizes in the range of 50-1000 μm in diameter, and the rest about 5% of bubbles are in millimeter level.
The ratio K of the hydrogen feed to the upflow fixed bed reactor A to the flow of make-up hydrogen feed to the upflow fixed bed reactor B (Nm 3/h) was 0.5:1.
Under the reaction conditions, maleic anhydride and gamma-butyrolactone solvents shown in table 1 and table 2 are taken as raw materials, and enter an up-flow fixed bed hydrogenation reactor A and an up-flow fixed bed hydrogenation reactor B to continuously carry out hydrogenation reaction to obtain hydrogenation products, and the reaction results are shown in table 4.
TABLE 4 reaction results
Note that: the total run time is the sum of the run times at the beginning and end of the reaction. Starting with a total conversion of the reaction below 99.8% and a total selectivity below 95% as nodes for termination of the run.
As can be seen from the above comparative examples and examples: (1) When maleic anhydride hydrogenation reaction is carried out by adopting a conventional trickle bed and conventional liquid phase hydrogenation, the reaction is uneven, the reaction efficiency is low, the problems of local hot spots of a catalyst bed layer, more side reactions and the like are easily caused, on one hand, the total operation time of the catalyst is influenced, and on the other hand, the selectivity of the catalyst is also influenced; (2) In the conventional technology, maleic anhydride hydrogenation reaction is based on low maleic anhydride concentration in the later reaction period, so that the contact mass transfer rate of hydrogen and maleic anhydride is very low, and the conversion rate of more than 99% is achieved, so that very long residence time is required, and the method is one of reasons for increasing side reactions; (3) When two-stage reactors are connected in series in the conventional technology, the reaction rate of the first reactor and the reaction rate, the temperature rise, the catalyst deactivation rate and the like of the second reactor are different, so that the catalysts of the two reactors cannot be deactivated synchronously, and the total operation period and the economy of the device are affected. Therefore, the invention sets two-stage serial continuous hydrogenation reactors, so that the hydrogenation reaction is mainly concentrated in the upflow fixed bed hydrogenation reactor A in the early stage of the reaction, and the hydrogenation reaction is transferred to the upflow hydrogenation reactor B along with the reaction, until the catalysts of the two reactors reach synchronous deactivation, in the process, the transfer of the hydrogenation reaction is regulated and controlled by hydrogen feeding, the proportion of microbubbles in the two reactors is cooperatively controlled, the conversion rate (more than or equal to 99.8%) and the selectivity (more than or equal to 98.5%) of the maleic anhydride hydrogenation reaction are ensured, the synchronous deactivation of the catalysts of the two reactors A and B is realized, the total operation period of the device is greatly prolonged, and the economical efficiency of the industrial device is improved.
Claims (11)
1. A reaction process for preparing succinic anhydride, comprising the following contents: (1) Uniformly mixing maleic anhydride solution and hydrogen to obtain mixed feed I, and feeding the mixed feed I into an up-flow fixed bed reactor A to perform a first hydrogenation reaction to obtain a first hydrogenation reaction product; (2) The first hydrogenation reaction product is heated and then is uniformly mixed with supplementary hydrogen to obtain mixed feed II, and the mixed feed II enters an up-flow fixed bed reactor B to carry out a second hydrogenation reaction to obtain a second hydrogenation reaction product; (3) After the second hydrogenation reaction product is heated, gas-liquid separation is carried out, one part of liquid phase circulates, and the other part enters a subsequent product fractionation unit; wherein the ratio K of the flow rate (Nm 3/h) of the hydrogen of the upflow fixed bed reactor A in the step (1) to the flow rate (Nm 3/h) of the supplementary hydrogen of the upflow fixed bed reactor B in the step (2) is 1: 10-10: 1.
2. The reaction process according to claim 1, wherein: the maleic anhydride content in the maleic anhydride solution is 0.03-0.3 g/mL; the maleic anhydride solution adopts one or more solvents selected from benzene, toluene, xylene, acetone, tetrahydrofuran, gamma-butyrolactone, methyl-grade acetone, cyclohexanone, ethyl acetate, diethyl succinate or ethylene glycol monomethyl ether.
3. The reaction process according to claim 1, wherein: in the mixed material I, hydrogen is uniformly dispersed in maleic anhydride solution, wherein more than or equal to 70% of hydrogen bubbles are in millimeter level, and the diameter d1 of the hydrogen bubbles is preferably 0.01-20 mm.
4. The reaction process according to claim 1, wherein: the first hydrogenation reaction conditions are as follows: the reaction temperature is 40-200 ℃, preferably 50-150 ℃; the reaction pressure is 0.5-10.0 MPa, preferably 1-5.0 MPa; the liquid hourly space velocity is 0.5 to 15.0h -1, preferably 3.0 to 8.0h -1.
5. The reaction process according to claim 1, wherein: the ratio of the volume flow rate of hydrogen (Nm 3/h) to the volume flow rate of maleic anhydride solution (m 3/h) in the upflow fixed bed reactor A was 5:1 to 100:1, preferably 10:1 to 60:1.
6. The reaction process according to claim 1, wherein: in the mixed material II, hydrogen is uniformly mixed and dispersed in the first hydrogenation product, wherein more than or equal to 70% of hydrogen bubbles are nano-micron-sized, and the diameter d2 of the hydrogen bubbles is preferably 10 nm-1000 mu m.
7. The reaction process according to claim 1, wherein: the second hydrogenation reaction conditions are as follows: the reaction temperature is 40-200 ℃, preferably 50-150 ℃; the reaction pressure is 0.5-10.0 MPa, preferably 1-5.0 MPa; the liquid hourly space velocity is 0.1 to 8.0h -1, preferably 0.5 to 3.0h -1.
8. The reaction process according to claim 1, wherein: the ratio of the volume flow rate of hydrogen (Nm 3/h) in the upflow fixed bed reactor B to the volume flow rate of maleic anhydride feed (m 3/h) in the upflow fixed bed reactor A was 1:1 to 60:1, preferably 5: 1-40: 1.
9. The reaction process according to claim 1, wherein: and 1 or more catalyst beds are arranged in the upflow fixed bed reactors A and B according to the requirement, and the catalyst adopts a catalyst with a hydrogenation function.
10. The reaction process according to claim 1, wherein: the maleic anhydride conversion rate of the first hydrogenation reaction is 41-99%, the maleic anhydride conversion rate at the initial stage of the reaction is 71-99%, and the maleic anhydride conversion rate at the final stage of the reaction is 41-70%; the maleic anhydride conversion rate of the second hydrogenation reaction is 1-40%, the maleic anhydride conversion rate at the initial stage of the reaction is 1-20%, and the maleic anhydride conversion rate at the final stage of the reaction is 21-40%; the initial stage and the final stage of the reaction are divided according to the total operation period of the catalyst, the initial stage of the reaction refers to the stage that the inlet temperature of the reactor meets the requirement of maleic anhydride conversion rate, the final stage of the reaction refers to the stage that the inlet temperature of the reactor does not meet the requirement of maleic anhydride conversion rate along with the extension of the reaction time and the reduction of the activity of the catalyst, the inlet temperature of the reactor needs to be increased to meet the requirement of maleic anhydride conversion rate, and the total operation period of the catalyst is reached after the inlet temperature of the reactor reaches the upper limit.
11. The reaction process according to claim 1, wherein: the first circulating material which circulates back to the upflow fixed bed reactor A accounts for 15 to 90 weight percent, preferably 30 to 80 weight percent of the fresh feed (maleic anhydride solution) of the upflow fixed bed reactor A; the second recycle to the upflow reactor B represents from 0 to 80% by weight, preferably from 0 to 40% by weight, of the fresh feed to the upflow reactor B.
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