CN114518050B - Dehydrogenation reaction heating system - Google Patents
Dehydrogenation reaction heating system Download PDFInfo
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- CN114518050B CN114518050B CN202011304673.8A CN202011304673A CN114518050B CN 114518050 B CN114518050 B CN 114518050B CN 202011304673 A CN202011304673 A CN 202011304673A CN 114518050 B CN114518050 B CN 114518050B
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- dehydrogenation reaction
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- 238000006356 dehydrogenation reaction Methods 0.000 title claims abstract description 145
- 238000010438 heat treatment Methods 0.000 title claims abstract description 44
- YNQLUTRBYVCPMQ-UHFFFAOYSA-N Ethylbenzene Chemical compound CCC1=CC=CC=C1 YNQLUTRBYVCPMQ-UHFFFAOYSA-N 0.000 claims abstract description 38
- NNPPMTNAJDCUHE-UHFFFAOYSA-N isobutane Chemical compound CC(C)C NNPPMTNAJDCUHE-UHFFFAOYSA-N 0.000 claims abstract description 30
- ATUOYWHBWRKTHZ-UHFFFAOYSA-N Propane Chemical compound CCC ATUOYWHBWRKTHZ-UHFFFAOYSA-N 0.000 claims abstract description 28
- 239000003085 diluting agent Substances 0.000 claims abstract description 26
- 238000000034 method Methods 0.000 claims abstract description 24
- 229930195733 hydrocarbon Natural products 0.000 claims abstract description 21
- 150000002430 hydrocarbons Chemical class 0.000 claims abstract description 21
- KVNYFPKFSJIPBJ-UHFFFAOYSA-N 1,2-diethylbenzene Chemical compound CCC1=CC=CC=C1CC KVNYFPKFSJIPBJ-UHFFFAOYSA-N 0.000 claims abstract description 20
- 239000004215 Carbon black (E152) Substances 0.000 claims abstract description 19
- 239000001282 iso-butane Substances 0.000 claims abstract description 15
- 239000001294 propane Substances 0.000 claims abstract description 14
- 239000000446 fuel Substances 0.000 claims description 84
- 239000002994 raw material Substances 0.000 claims description 65
- 238000006243 chemical reaction Methods 0.000 claims description 57
- 238000002485 combustion reaction Methods 0.000 claims description 24
- 239000002912 waste gas Substances 0.000 claims description 22
- 239000007789 gas Substances 0.000 claims description 19
- 239000000203 mixture Substances 0.000 claims description 12
- 150000001335 aliphatic alkanes Chemical class 0.000 claims description 7
- 150000004945 aromatic hydrocarbons Chemical class 0.000 claims description 4
- HYFLWBNQFMXCPA-UHFFFAOYSA-N 1-ethyl-2-methylbenzene Chemical compound CCC1=CC=CC=C1C HYFLWBNQFMXCPA-UHFFFAOYSA-N 0.000 claims description 3
- IJDNQMDRQITEOD-UHFFFAOYSA-N n-butane Chemical compound CCCC IJDNQMDRQITEOD-UHFFFAOYSA-N 0.000 claims description 3
- 239000000463 material Substances 0.000 abstract description 30
- 238000007086 side reaction Methods 0.000 abstract description 11
- 238000005336 cracking Methods 0.000 abstract description 9
- PPBRXRYQALVLMV-UHFFFAOYSA-N Styrene Chemical compound C=CC1=CC=CC=C1 PPBRXRYQALVLMV-UHFFFAOYSA-N 0.000 description 16
- VQTUBCCKSQIDNK-UHFFFAOYSA-N Isobutene Chemical compound CC(C)=C VQTUBCCKSQIDNK-UHFFFAOYSA-N 0.000 description 14
- 238000005265 energy consumption Methods 0.000 description 10
- 238000007599 discharging Methods 0.000 description 9
- CURLTUGMZLYLDI-UHFFFAOYSA-N Carbon dioxide Chemical compound O=C=O CURLTUGMZLYLDI-UHFFFAOYSA-N 0.000 description 8
- VNWKTOKETHGBQD-UHFFFAOYSA-N methane Chemical compound C VNWKTOKETHGBQD-UHFFFAOYSA-N 0.000 description 8
- MYRTYDVEIRVNKP-UHFFFAOYSA-N 1,2-Divinylbenzene Chemical compound C=CC1=CC=CC=C1C=C MYRTYDVEIRVNKP-UHFFFAOYSA-N 0.000 description 6
- 229910052799 carbon Inorganic materials 0.000 description 6
- XLYOFNOQVPJJNP-UHFFFAOYSA-N water Chemical compound O XLYOFNOQVPJJNP-UHFFFAOYSA-N 0.000 description 6
- 229910002092 carbon dioxide Inorganic materials 0.000 description 4
- 239000003054 catalyst Substances 0.000 description 4
- 239000003345 natural gas Substances 0.000 description 4
- 239000000047 product Substances 0.000 description 4
- QQONPFPTGQHPMA-UHFFFAOYSA-N propylene Natural products CC=C QQONPFPTGQHPMA-UHFFFAOYSA-N 0.000 description 4
- 125000004805 propylene group Chemical group [H]C([H])([H])C([H])([*:1])C([H])([H])[*:2] 0.000 description 4
- 239000001569 carbon dioxide Substances 0.000 description 3
- 239000000376 reactant Substances 0.000 description 3
- 238000004227 thermal cracking Methods 0.000 description 3
- 150000001336 alkenes Chemical class 0.000 description 2
- 125000000217 alkyl group Chemical group 0.000 description 2
- 230000000052 comparative effect Effects 0.000 description 2
- VXNZUUAINFGPBY-UHFFFAOYSA-N 1-Butene Chemical compound CCC=C VXNZUUAINFGPBY-UHFFFAOYSA-N 0.000 description 1
- JLBJTVDPSNHSKJ-UHFFFAOYSA-N 4-Methylstyrene Chemical compound CC1=CC=C(C=C)C=C1 JLBJTVDPSNHSKJ-UHFFFAOYSA-N 0.000 description 1
- 125000003118 aryl group Chemical group 0.000 description 1
- 239000001273 butane Substances 0.000 description 1
- IAQRGUVFOMOMEM-UHFFFAOYSA-N butene Natural products CC=CC IAQRGUVFOMOMEM-UHFFFAOYSA-N 0.000 description 1
- 230000003197 catalytic effect Effects 0.000 description 1
- 239000007795 chemical reaction product Substances 0.000 description 1
- 238000004891 communication Methods 0.000 description 1
- 230000000694 effects Effects 0.000 description 1
- 238000005485 electric heating Methods 0.000 description 1
- 238000004519 manufacturing process Methods 0.000 description 1
- 238000012986 modification Methods 0.000 description 1
- 230000004048 modification Effects 0.000 description 1
- OFBQJSOFQDEBGM-UHFFFAOYSA-N n-pentane Natural products CCCCC OFBQJSOFQDEBGM-UHFFFAOYSA-N 0.000 description 1
- JRZJOMJEPLMPRA-UHFFFAOYSA-N olefin Natural products CCCCCCCC=C JRZJOMJEPLMPRA-UHFFFAOYSA-N 0.000 description 1
- 238000003303 reheating Methods 0.000 description 1
- 238000012827 research and development Methods 0.000 description 1
- 239000007858 starting material Substances 0.000 description 1
- 239000000126 substance Substances 0.000 description 1
Classifications
-
- F—MECHANICAL ENGINEERING; LIGHTING; HEATING; WEAPONS; BLASTING
- F28—HEAT EXCHANGE IN GENERAL
- F28D—HEAT-EXCHANGE APPARATUS, NOT PROVIDED FOR IN ANOTHER SUBCLASS, IN WHICH THE HEAT-EXCHANGE MEDIA DO NOT COME INTO DIRECT CONTACT
- F28D21/00—Heat-exchange apparatus not covered by any of the groups F28D1/00 - F28D20/00
-
- F—MECHANICAL ENGINEERING; LIGHTING; HEATING; WEAPONS; BLASTING
- F28—HEAT EXCHANGE IN GENERAL
- F28D—HEAT-EXCHANGE APPARATUS, NOT PROVIDED FOR IN ANOTHER SUBCLASS, IN WHICH THE HEAT-EXCHANGE MEDIA DO NOT COME INTO DIRECT CONTACT
- F28D21/00—Heat-exchange apparatus not covered by any of the groups F28D1/00 - F28D20/00
- F28D2021/0019—Other heat exchangers for particular applications; Heat exchange systems not otherwise provided for
- F28D2021/0022—Other heat exchangers for particular applications; Heat exchange systems not otherwise provided for for chemical reactors
-
- 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
- Y02P20/10—Process efficiency
Landscapes
- Engineering & Computer Science (AREA)
- Physics & Mathematics (AREA)
- Thermal Sciences (AREA)
- Mechanical Engineering (AREA)
- General Engineering & Computer Science (AREA)
- Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)
- Low-Molecular Organic Synthesis Reactions Using Catalysts (AREA)
Abstract
The invention discloses a heater and a use method thereof, and adopts the technical scheme that the temperature of the wall of a heat exchange tube contacted with raw hydrocarbon is lower than 750 ℃, side reactions such as cracking and the like are less, the selectivity of hydrocarbon dehydrogenation products can be improved by 0.5-5 percent, the pressure drop of the heater is reduced by more than 50 percent compared with that of a traditional heating furnace, the heater can be used for the direct dehydrogenation reaction of propane and isobutane, and the dehydrogenation reaction of ethylbenzene, diethylbenzene and the like which need diluents, and the problem that the materials of high-temperature pipelines of the traditional heating furnace with too little diluent consumption are limited can be solved.
Description
Technical Field
The invention relates to a dehydrogenation reaction heating system, in particular to a heat exchanger for a raw material of hydrocarbon dehydrogenation reaction and a heating method using the heat exchanger.
Background
The hydrocarbon dehydrogenation reaction is common and important in the chemical industry, and the industrial dehydrogenation reaction is mainly applied to prepare lower olefins by dehydrogenation of lower alkanes and alkyl aromatics by dehydrogenation of lower alkanes, wherein the lower alkanes are prepared by dehydrogenation of propane to prepare propylene, the isobutane to prepare isobutene, the butane to prepare normal butene and the like, the alkyl aromatics are prepared by dehydrogenation of ethylbenzene to prepare styrene, the diethylbenzene is prepared by dehydrogenation of diethylbenzene, and the methyl ethylbenzene is prepared by dehydrogenation of methyl styrene and the like. The common characteristic of this type of reaction is that the reaction needs to be dehydrogenated at high temperature, low pressure and in the presence of a catalyst, so that the reaction raw materials must be heated to a relatively high temperature before entering the reactor, and the pressure drop of the system is kept low. In order to obtain high conversion and olefin selectivity, research and development of efficient reaction process conditions, particularly heating methods of the feedstock, is one of the key factors. The heating process of the proper raw materials can not only improve the utilization rate of energy, but also reduce side reactions such as thermal cracking and the like caused by overhigh temperature of the tube wall of the heat exchange equipment, reduce the pressure drop of the system and improve the selectivity of target products.
The patent CN208098029U proposes a heat exchanger method using dehydrogenation reaction products and raw materials, the reaction raw materials at 26-40 ℃ can be heated to 420-460 ℃ by the reaction material inlet and outlet heat exchanger, so as to avoid bias flow problem among a plurality of heat exchangers, and improve stability of the propane dehydrogenation device, but because the reaction temperature of the propane dehydrogenation reactor is 595-610 ℃, the outlet material flow of the heat exchanger still needs to be heated to the reaction temperature by the heating furnace and depressurized to 0.044-0.066 MPa, and the wall temperature of the furnace tube of the common heating furnace reaches more than 800 ℃, so that side reactions such as cracking of propane are aggravated, and the pressure drop of the heating furnace is large, so that energy consumption of a subsequent compressor is increased.
Patent CN104072325a discloses a method for improving the dehydrogenation reaction performance of low-carbon alkane, which adopts a fixed bed reactor with an internal electric heating tube in the dehydrogenation process, provides heat for a catalyst in the dehydrogenation reaction process of low-carbon alkane, reduces the temperature drop of a catalyst bed layer caused by strong endothermic dehydrogenation reaction, and reduces the heat load of an electric heater in front of the reactor, thereby reducing the thermal cracking of low-carbon alkane in the electric heater, but the wall temperature of the heating tube of the electric heater is very high, and the thermal cracking side reaction of low-carbon alkane in the electric heater is unavoidable.
The patent CN110903155A proposes a method, a device and a reaction system for a low-carbon alkane dehydrogenation process, which preheat C3-C5 low-carbon alkane feed gas, CO and/or CO2 process gas at 200-500 ℃ and then enter a reactor to be converted with a dehydrogenation catalyst under the conditions of the reaction temperature of 500-700 ℃ and the reaction pressure of 10-100 kPa.
The patent CN103664497B proposes a method for producing styrene by ethylbenzene catalytic dehydrogenation, and the method provided by the invention is characterized in that raw material ethylbenzene is divided into two streams, water vapor and a first stream of ethylbenzene are mixed and enter a first dehydrogenation reactor, and outlet gas of the first dehydrogenation reactor and a second stream of raw material ethylbenzene are mixed and enter a second dehydrogenation reactor. The conventional ethylbenzene dehydrogenation reaction process adopts steam as a diluent, a heating furnace is used for heating the steam to more than 800 ℃, then the steam and ethylbenzene are mixed to 600-650 ℃ and enter a dehydrogenation reactor, and the dehydrogenation reactor is limited by the material temperature of a high-temperature steam pipeline, so that the energy consumption of the whole device is very high due to the higher water ratio (steam/ethylbenzene).
Disclosure of Invention
The invention aims to solve the technical problems that the existing hydrocarbon dehydrogenation reaction device adopts a fuel heating furnace to directly heat raw materials to high temperature, and the temperature of the tube wall of a furnace tube is too high to cause the aggravation of hydrocarbon cracking side reaction; or the fuel heating furnace is adopted to heat the diluent such as water vapor and the like, and then the diluent is used for reheating the raw materials, so that the problems of low heat efficiency, high energy consumption and large pressure drop exist.
To this end, the present invention provides a novel heater and a method for heating a hydrocarbon dehydrogenation reaction feedstock using the same. The method uses the heater using fuel as a heat source to directly provide heat for the reaction raw materials, and has the advantages of low wall temperature of a heat exchange tube, less cracking side reaction, low pressure drop and applicability to dehydrogenation reactions such as diluent or not.
To solve the above technical problems, a first aspect of the present invention provides a heater, which includes a fuel chamber, an exhaust gas chamber, an air chamber, and a heat exchange tube box in which at least one heat exchange tube is disposed, the heat exchange tube being of a sleeve structure including a fuel tube, a combustion zone, and an air preheating zone from inside to outside,
wherein the fuel tube is in communication with the fuel chamber; fuel enters from a fuel feed port provided in the fuel chamber, and enters the fuel pipe through the fuel chamber;
the air preheating zone is communicated with the air chamber, air enters from an air feed port arranged in the air chamber, and enters the air preheating zone through the air chamber;
the combustion zone is connected with the exhaust gas chamber; the fuel and air are combusted in the combustion zone, and combusted exhaust gas is discharged from the exhaust gas chamber and from a discharge port provided in the exhaust gas chamber.
In some preferred embodiments of the invention, the fuel chamber has a space of 10% -150% of the fuel tube space.
In some preferred embodiments of the invention, the exhaust chamber is 10% -200% of the fuel pipe.
In some preferred embodiments of the invention, the air chamber has a space of 10% -200% of the space of the air preheating zone.
In some preferred embodiments of the invention, the air chamber and the fuel chamber are provided with a single stage baffle distributor, the distributor being perpendicular to the feed direction.
A second aspect of the present invention provides a heating method for a hydrocarbon dehydrogenation reaction feedstock comprising the use of the heater, wherein
The heater is arranged in the dehydrogenation reaction zone, dehydrogenation reaction raw materials or a mixture of the dehydrogenation reaction raw materials and a diluent are heated to a temperature required by reaction and then introduced into the dehydrogenation reaction zone, fuel enters the fuel pipe from the fuel feed inlet through the fuel chamber, air enters the air preheating zone from the air feed inlet through the air chamber, fuel and air are combusted in the combustion process, and combusted waste gas is discharged from the waste gas chamber through the waste gas discharge port;
the combustion zone height Wen Guanbi is not in direct contact with the dehydrogenation reaction raw material, but is heated by air in the air preheating zone with lower temperature, and the temperature of the heat exchange tube wall of the air preheating zone in contact with the dehydrogenation reaction raw material is lower than 750 ℃.
The heater is of a vertical container structure, and the inside of the container is composed of a fuel chamber, an exhaust gas chamber, an air chamber and a heat exchange tube box from top to bottom; the diameter of the heat exchange tube box of the heater is the diameter required by the dehydrogenation reaction raw material to be reduced to 10kPa through the equipment, preferably the diameter required by the dehydrogenation reaction raw material to be reduced to 5 kPa; the diameter of a fuel feed pipe in the heater cluster heat exchange pipe sleeve is the diameter required by the fuel passing flow speed of 1-30 m/s; the diameter of the combustion zone is the diameter required by the flow rate of the gas after combustion to be 5-150 m/s; the diameter of the air preheating zone is the diameter required by the air passing flow rate of 1-100 m/s; the fuel chamber space of the heater is 10% -150% of the fuel pipe space of all the cluster heat exchange pipes, the waste gas chamber space is 10% -200% of the waste gas space of all the cluster heat exchange pipes, and the air chamber space is 10% -200% of the air preheating zone space of all the cluster heat exchange pipes; the temperature of the tube wall of the air preheating zone of the cluster heat exchange tube, which is in contact with the hydrocarbon raw material, is 300-750 ℃.
In some preferred embodiments of the invention, the dehydrogenation feed is an alkane and/or an arene. Further, the alkane is preferably at least one of propane, n-butane and isobutane; the aromatic hydrocarbon is preferably at least one of aromatic ethylbenzene, diethylbenzene and methyl ethylbenzene.
In a further preferred embodiment of the present invention, the heater has a dehydrogenation reaction raw material inlet temperature of 20 to 500 ℃ and a dehydrogenation reaction raw material outlet temperature of 300 to 800 ℃; the inlet pressure of the dehydrogenation reaction raw material is 20-300 kPaA, and the outlet pressure of the dehydrogenation reaction raw material is 10-300 kPaA.
The invention adopts a heater which uses fuel as a heat source and is arranged in front of a hydrocarbon dehydrogenation reactor to heat dehydrogenation reaction raw materials to the temperature required by reaction, the heater is of a container structure, the inside of the heater is divided into a fuel feeding chamber, a waste gas discharging chamber, an air feeding chamber and a heat exchange tube box, the heat exchange tube box is fully distributed with clustered heat exchange tubes, the outside of the heat exchange tube is a dehydrogenation reaction raw material channel, each heat exchange tube is of a sleeve structure, and a fuel feeding tube, an air preheating zone and a combustion zone are respectively arranged in the sleeve from inside to outside and are respectively communicated with the fuel chamber, the waste gas chamber and the air chamber. The air and the fuel are combusted in a combustion zone in the clustered heat exchange tubes of the sleeve structure, the heat combusted in the combustion zone does not directly heat the raw materials, but firstly heats the air of an air preheating zone positioned outside the sleeve combustion zone, and then the dehydrogenation reaction raw materials in the raw material channel between the interior of the container and the clustered heat exchange tubes are reheated by the air of the air preheating zone. Thus, the high Wen Guanbi of the combustion zone is not in direct contact with the material flow, but the raw materials are heated by the air of the air preheating zone with lower temperature, so that the side reaction of cracking caused by hydrocarbon contacting with the high-temperature pipe wall is reduced; meanwhile, the number of the clustered heat exchange tubes is large, the dehydrogenation reaction raw material can be heated to the required temperature in a short time through the heater, the pressure drop is small, and the higher conversion rate and selectivity of the hydrocarbon dehydrogenation reaction are ensured.
By adopting the technical scheme of the invention, the wall temperature of the heat exchange tube contacted with the dehydrogenation reaction raw material is lower than 750 ℃, the side reactions such as cracking are less, the selectivity of hydrocarbon dehydrogenation products can be improved by 0.5-5 percent, the pressure drop of the heater is reduced by more than 50 percent compared with that of the traditional heating furnace, the heater can be used for the direct dehydrogenation reaction of propane and isobutane, and the dehydrogenation reaction of ethylbenzene, diethylbenzene and the like which need diluents, and the problem that the materials of high-temperature pipelines are limited due to the fact that the consumption of the diluents is too small in the traditional heating furnace can be solved. By applying the technical scheme, a better technical effect is achieved.
Drawings
FIG. 1 is a typical prior art diluent-free single stage hydrocarbon dehydrogenation process.
FIG. 2 is a typical process for two-stage hydrocarbon dehydrogenation using a diluent.
FIG. 3 shows a heater structure for hydrocarbon dehydrogenation according to the present invention.
FIG. 4 shows a technical process of the present invention in the dehydrogenation of single-stage hydrocarbons without diluent.
FIG. 5 shows a technical process of the present invention in the dehydrogenation of two-stage hydrocarbons with a diluent.
Detailed Description
For a better understanding of the present invention, reference will now be made in detail to the present invention, examples of which are illustrated in the accompanying drawings and described in the accompanying drawings.
The endpoints and any values of the ranges disclosed herein are not limited to the precise range or value, and are understood to encompass values approaching those ranges or values. For numerical ranges, one or more new numerical ranges may be found between the endpoints of each range, between the endpoint of each range and the individual point value, and between the individual point value, in combination with each other, and are to be considered as specifically disclosed herein.
The labels in each of the figures are independent in the present invention, and even if identical labels are present in different figures, these identical labels do not represent the same meaning, but merely represent the respective meanings in the respective figures.
In fig. 1, A1 is a heating furnace; a2 is a dehydrogenation reactor; 101 is a dehydrogenation reaction feedstock; 102 is the dehydrogenation reaction raw material discharge of the heating furnace; 103 is the dehydrogenation reactor discharge.
The dehydrogenation reaction raw material 101 is heated to a reaction temperature by a heating furnace A1 to obtain a material flow 102, the material flow 102 enters a dehydrogenation reactor A2 for dehydrogenation reaction, and the material flow 103 after the reaction enters a subsequent working section.
In fig. 2, B1 is a first heating furnace; b2 is an intermediate reheater; b3 is a second heating furnace; b4 is a first dehydrogenation reactor; b5 is a second dehydrogenation reactor; b6 is a reaction feeding and discharging heat exchanger; 201 is a diluent; 202 is the diluent heated by the first heating furnace B1; 203 is the diluent subjected to heat exchange by the intermediate reheater B2; 204 is the diluent heated by the second heating furnace B3; 205 is a dehydrogenation feed; 206 feed the first dehydrogenation reactor B4; 207 is the discharge of the first dehydrogenation reactor B4; 208 is the discharge of the second dehydrogenation reactor B5; 209 is the reactant which exchanges heat through the reaction feeding and discharging heat exchanger B6.
After heat exchange is carried out on dehydrogenation reaction raw materials 205 through a reaction material inlet and outlet heat exchanger B6, the dehydrogenation raw materials are mixed with a diluent 204 from a second heating furnace B3 to obtain a material flow 206, the material flow 206 enters a first dehydrogenation reactor B4, a material outlet 207 exchanges heat with the diluent 202 from the first heating furnace B1 in an intermediate reheater B2, the material flow enters a second dehydrogenation reactor B5 after the required reaction temperature is reached, and a material flow 209 is obtained after heat exchange is carried out on a material flow 208 after dehydrogenation through the reaction material inlet and outlet heat exchanger B6 and enters a subsequent working section.
In fig. 3, i is a fuel chamber; II is an exhaust gas chamber; III is an air chamber; IV is a heat exchange tube box; v is a fuel pipe; VI is a combustion zone of the heat exchange sleeve; VII is an air preheating zone of the heat exchange sleeve; VIII is a heat exchange sleeve; IX is a single stage baffle distributor. C1 is a fuel feed inlet; c2 is an exhaust gas discharge port; c3 is an air feed inlet; c4 is a dehydrogenation reaction raw material feed inlet; c5 is the discharge port of the dehydrogenation reaction raw material.
The fuel enters the fuel pipe V from the fuel chamber I through the fuel inlet C1, the air enters the heat exchange sleeve air preheating zone VII from the air inlet C3 through the air chamber III, the fuel and the air are combusted in the heat exchange sleeve combustion zone VI, and the combusted waste gas is discharged from the waste gas chamber II through the waste gas discharge port C2. The dehydrogenation reaction raw material enters a heat exchange tube box IV from a dehydrogenation reaction raw material feed port C4, exchanges heat with an air preheating zone VII in a heat exchange sleeve VIII, and the dehydrogenation reaction raw material after heat exchange is discharged from a dehydrogenation reaction raw material discharge port C5.
The high Wen Guanbi of the heat exchange sleeve combustion zone VI is not in direct contact with the dehydrogenation reaction raw material, but the raw material is heated by the air of the heat exchange sleeve air preheating zone VII with lower temperature, so that the side reaction of cracking caused by the contact of the dehydrogenation reaction raw material with the high-temperature pipe wall is reduced; meanwhile, the number of the clustered heat exchange tubes is large, the dehydrogenation reaction raw material can be heated to the required temperature in a short time through the heater, the pressure drop is small, and the dehydrogenation reaction raw material dehydrogenation energy is ensured to obtain higher conversion rate and selectivity.
The wall temperature of the heat exchange tube contacted with the dehydrogenation reaction raw material is lower than 750 ℃, the side reactions such as cracking are few, the selectivity of dehydrogenation products can be improved by 0.5-5%, and the pressure drop of the heater is reduced by more than 50% compared with that of a traditional heating furnace.
In fig. 4, D1 is a heater of the present invention, D2 is a dehydrogenation reactor, 401 is a dehydrogenation reaction raw material, 402 is a dehydrogenation reaction raw material which is discharged after being heated by the heater D1, 403 is a dehydrogenation reactor discharge, 404 is fuel, 405 is exhaust gas, and 406 is air.
The dehydrogenation reaction raw material 401 is heated to a reaction temperature by a heater D1 to obtain a material flow 402, the material flow 402 enters a dehydrogenation reactor D2 for dehydrogenation reaction, and the material flow 403 after the reaction enters a subsequent working section.
In fig. 5, E1 is a first heater, E2 is a second heater, E3 is a first dehydrogenation reactor, E4 is a second dehydrogenation reactor, and E5 is a reaction charge-discharge heat exchanger; 501 is a diluent, 502 is a dehydrogenation reaction raw material, 503 is a dehydrogenation reaction raw material heated by a feeding and discharging heat exchanger E5, 504 is fuel entering a first heater E1, 505 is waste gas discharged by the first heater E1, 506 is air entering the first heater E1, 507 is a first dehydrogenation reactor E3 feed, 508 is a first dehydrogenation reactor E3 discharge, 509 is a dehydrogenation reactant heated by a second heater, 510 is a second dehydrogenation reactor E4 discharge, 511 is a reactant subjected to heat exchange by the reaction feeding and discharging heat exchanger E5, 512 is fuel entering a second heater E2, 513 is waste gas discharged by the second heater E2, and 514 is air entering the second heater E2.
The dehydrogenation reaction raw material 502 and the diluent 501 are mixed and then subjected to heat exchange through a reaction material inlet and outlet heat exchanger E5, the mixture is heated through a first heater E1 to obtain a material flow 507, the material flow enters a first dehydrogenation reactor E3, a material flow 509 is obtained after a material outlet 508 is heated through a second heater E2, the material flow enters a second dehydrogenation reactor E4, and the material flow 510 after dehydrogenation enters a subsequent working section after heat exchange through the reaction material inlet and outlet heat exchanger E5 to obtain a material flow 511.
Example 1
The device for preparing isobutene by dehydrogenating 10 ten thousand tons/year isobutane adopts the reaction flow shown in figure 4, and the heater adopts the structure shown in figure 3. The flow of the raw material isobutane is 32 tons/hour, the temperature of the preheated isobutane is 400 ℃, the preheated isobutane enters a heater for heating, the preheated isobutane enters a dehydrogenation reactor after being heated to the temperature of 500 ℃, the reaction inlet temperature is 500 ℃, the reaction outlet temperature is 450 ℃, and the reaction pressure is 150kPaA. The diameter of the heater tube box is 2000mm, the length of the heat exchange sleeve is 3 m, the number of the heat exchange sleeves is 580, the diameter of the fuel tube in the heat exchange sleeve is 19mm, the outer diameter of the combustion zone is 38mm, the outer diameter of the air preheating zone is 50mm, the diameter of the fuel chamber is 2000mm, the height of the fuel chamber is 200mm, the diameter of the waste gas chamber is 2000mm, the height of the waste gas chamber is 400mm, the diameter of the air chamber is 2000mm, the height of the air chamber is 300mm, single-stage baffles are arranged in the fuel chamber and the air chamber, the type is flat plate type, and the direction of the single-stage baffles is perpendicular to the feed inlet. The fuel used was natural gas in an amount of 0.85 ton/hr and air in an amount of 25 ton/hr.
The total pressure drop of the heat exchange tube box of the heater is 5kPa, and the tube wall temperature of the air preheating zone of the clustered heat exchange tube contacted with isobutane is 650 ℃. The isobutene selectivity of the reaction unit is 92% and the energy consumption is 105 kg of standard oil per ton of isobutene.
Example 2
The apparatus for producing propylene by dehydrogenating propane of 20 ten thousand tons/year adopts the reaction flow shown in fig. 4, and the heater adopts the structure shown in fig. 3. The flow rate of the raw material propane is 55 tons/hour, the preheated propane temperature is 400 ℃, the raw material propane enters a heater for heating, the raw material propane enters a dehydrogenation reactor after being heated to 600 ℃, the reaction inlet temperature is 600 ℃, the reaction outlet temperature is 550 ℃, and the reaction pressure is 50kPaA. The diameter of the heater tube box is 3000mm, the length of the heat exchange sleeve is 3 m, the diameter of the fuel tube in the heat exchange sleeve is 19mm, the outer diameter of the combustion zone is 38mm, the outer diameter of the air preheating zone is 50mm, the diameter of the fuel chamber is 3000mm, the height of the fuel chamber is 400mm, the diameter of the waste gas chamber is 3000mm, the height of the waste gas chamber is 600mm, the diameter of the air chamber is 3000mm, the height of the air chamber is 500mm, single-stage baffles are arranged in the fuel chamber and the air chamber, the type is conical, and the direction is perpendicular to the feed inlet. The fuel used was natural gas in an amount of 1.8 tons/hour and air in an amount of 55 tons/hour.
The total pressure drop of the heat exchange tube box of the heater is 10kPa, and the tube wall temperature of the air preheating zone of the clustered heat exchange tube contacted with propane is 700 ℃. The propylene selectivity of the reaction unit is 91% and the energy consumption is 95 kg of standard oil/ton of propylene.
Example 3
The device for preparing styrene by ethylbenzene dehydrogenation of 10 ten thousand tons/year adopts a reaction flow shown in fig. 5, adopts water vapor as a diluent, and the structures of the first heater and the second heater are shown in fig. 3. The flow rate of ethylbenzene as a raw material is 20.5 tons/hour, the flow rate of water vapor as a diluent is 25 tons/hour, the temperature of the ethylbenzene and the water vapor is 500 ℃ after being mixed by a reaction feeding and discharging heat exchanger, the mixture is heated to 610 ℃ by a first heater and enters a first dehydrogenation reactor, the outlet temperature of the reactor is 530 ℃, the mixture is heated to 615 ℃ by a second heater and enters a second dehydrogenation reactor, the outlet temperature of the reactor is 540 ℃, the mixture is heated to 340 ℃ by a reaction feeding and discharging heat exchanger, and the mixture enters a subsequent working section. The first heater and the second heater are the same in size, the diameter of the tube box is 3200mm, the length of the heat exchange sleeve is 4 m, 1560 heat exchange sleeves are arranged in total, the diameter of the fuel tube in the heat exchange sleeve is 25mm, the outer diameter of the combustion area is 50mm, the outer diameter of the air preheating area is 70mm, the diameter of the fuel chamber is 3200mm, the height of the fuel chamber is 500mm, the diameter of the waste gas chamber is 3200mm, the height of the waste gas chamber is 800mm, the diameter of the air chamber is 3200mm, the height of the air chamber is 800mm, the fuel chamber and the air chamber are provided with single-stage baffles, the type is conical, and the direction of the fuel chamber and the air chamber are perpendicular to the feed inlet. The fuel used was natural gas in an amount of 1.6 tons/hour and air in an amount of 48 tons/hour.
The total pressure drop of the heat exchange tube boxes of the first heater and the second heater is 3kPa and 4kPa respectively, and the temperature of the tube wall of the air preheating zone of the clustered heat exchange tube contacted with ethylbenzene is 680 ℃. The selectivity of the styrene in the reaction unit is 98%, and the energy consumption is 80 kg of standard oil per ton of styrene.
Example 4
The device for preparing divinylbenzene by 10 ten thousand tons/year diethylbenzene dehydrogenation adopts the reaction flow shown in fig. 5, adopts carbon dioxide as a diluent, and has the structures shown in fig. 3. The flow rate of diethylbenzene serving as a raw material is 22 tons/hour, the flow rate of carbon dioxide serving as a diluent is 30 tons/hour, the mixture of diethylbenzene and carbon dioxide is subjected to heat exchange by a reaction feeding and discharging heat exchanger to 520 ℃, the mixture is heated to 620 ℃ by a first heater and enters a first dehydrogenation reactor, the outlet temperature of the reactor is 550 ℃, the mixture is heated to 630 ℃ by a second heater and enters a second dehydrogenation reactor, the outlet temperature of the reactor is 560 ℃, the mixture is subjected to heat exchange by a reaction feeding and discharging heat exchanger to 360 ℃, and the mixture enters a subsequent working section. The first heater and the second heater are the same in size, the diameter of the tube box is 3600mm, the length of the heat exchange sleeve is 5m, the total number of the heat exchange sleeves is 1980, the diameter of the fuel tube in the heat exchange sleeve is 19mm, the outer diameter of the combustion area is 30mm, the outer diameter of the air preheating area is 58mm, the diameter of the fuel chamber is 3600mm, the height is 600mm, the diameter of the waste gas chamber is 3600mm, the height is 900mm, the diameter of the air chamber is 3600mm, the height is 800mm, the fuel chamber and the air chamber are provided with single-stage baffles, the type is flat plates, and the direction of the fuel chamber and the air chamber are perpendicular to the feed inlet. The fuel used was natural gas in an amount of 2.8 tons/hour and air in an amount of 70 tons/hour.
The total pressure drop of the heat exchange tube boxes of the first heater and the second heater is 5kPa and 6kPa respectively, and the tube wall temperature of the air preheating zone of the clustered heat exchange tube contacted with diethylbenzene is 720 ℃. The selectivity of the divinylbenzene as a reaction unit is 96% and the energy consumption is 220 kg of standard oil per ton of divinylbenzene.
Comparative example 1
The device for preparing isobutene by dehydrogenating 10 ten thousand tons/year isobutane adopts the reaction flow shown in figure 1, and the heating furnace is a cylindrical heating furnace. The flow rate of isobutane as the starting material and the reaction conditions were the same as in example 1.
The total pressure drop of the isobutane material in the heating furnace is 50kPa, and the pipe wall temperature of the contact between the isobutane and the pipe wall of the furnace pipe is 820 ℃. The isobutene selectivity of the reaction unit is 90%, and the energy consumption is 120 kg of standard oil per ton of isobutene.
Comparative example 2
The device for preparing styrene by ethylbenzene dehydrogenation of 10 ten thousand tons/year adopts a reaction flow shown in figure 2, and the heating furnace is a square box furnace. The flow rate of ethylbenzene as a raw material and the reaction conditions were the same as in example 3.
The steam pressure drop of the diluent of the heating furnace is 200kPa, and the pipe wall temperature of the contact between the ethylbenzene and the pipe wall of the furnace pipe is 850 ℃. The selectivity of the styrene in the reaction unit is 96%, and the energy consumption is 95 kg of standard oil per ton of styrene.
Compared with the prior art, the technical scheme of the invention has the advantages that the temperature of the heat exchange tube wall contacted with the dehydrogenation reaction raw material is lower than 750 ℃, the side reactions such as cracking are less, the selectivity of the hydrocarbon dehydrogenation product can be improved by 0.5-5 percent, and the energy consumption is obviously reduced.
While the invention has been described with reference to a preferred embodiment, various modifications may be made and equivalents may be substituted for elements thereof without departing from the scope of the invention. In particular, the technical features mentioned in the respective embodiments may be combined in any manner as long as there is no structural conflict. The present invention is not limited to the specific embodiments disclosed herein, but encompasses all technical solutions falling within the scope of the claims.
Claims (10)
1. The dehydrogenation reaction heating system comprises a heater arranged in a dehydrogenation reaction zone, wherein the dehydrogenation reaction zone is a reaction zone consisting of a single section, two sections and more dehydrogenation reactors connected in series, and at least one heater using fuel as a heat source is arranged in the reaction zone;
the heater comprises a fuel chamber, an exhaust gas chamber, an air chamber and a heat exchange tube box, wherein at least one heat exchange tube is arranged in the heat exchange tube box, the heat exchange tube is of a sleeve structure, the sleeve structure comprises a fuel tube, a combustion zone and an air preheating zone from inside to outside,
wherein the fuel tube communicates with the fuel chamber; fuel enters from a fuel feed port provided in the fuel chamber, and enters the fuel pipe through the fuel chamber;
the air preheating zone is communicated with the air chamber, air enters from an air feed port arranged in the air chamber, and enters the air preheating zone through the air chamber;
the combustion zone is connected with the exhaust gas chamber; the fuel and air are combusted in the combustion zone, and combusted exhaust gas is discharged from the exhaust gas chamber and from a discharge port provided in the exhaust gas chamber.
2. The dehydrogenation reaction heating system of claim 1, wherein the fuel chamber has a space of 10% -150% of the fuel pipe space.
3. The dehydrogenation reaction heating system of claim 1 or 2, wherein the exhaust gas chamber has a space of 10% -200% of the fuel pipe.
4. The dehydrogenation reaction heating system of claim 1, wherein the air chamber has a space of 10% -200% of the space of the air preheating zone.
5. The dehydrogenation reaction heating system of claim 1, wherein a single stage baffle distributor is provided in each of the air chamber and the fuel chamber, the distributor being perpendicular to the feed direction.
6. A heating method for a hydrocarbon dehydrogenation reaction feedstock comprising using the dehydrogenation reaction heating system as set forth in any one of claims 1 to 5, wherein
The heater is arranged in the dehydrogenation reaction zone, and the dehydrogenation reaction raw material or the mixture of the dehydrogenation reaction raw material and the diluent is heated to the temperature required by the reaction and then is introduced into the dehydrogenation reaction zone;
the fuel enters the fuel pipe from the fuel feed inlet through the fuel chamber, the air enters the air preheating zone from the air feed inlet through the air chamber, the fuel and the air burn in the combustion zone, and the burnt waste gas is discharged from the waste gas chamber through the waste gas discharge outlet;
the combustion zone height Wen Guanbi is not in direct contact with the dehydrogenation reaction raw material, but is heated by air in the air preheating zone with lower temperature, and the temperature of the heat exchange tube wall of the air preheating zone in contact with the dehydrogenation reaction raw material is lower than 750 ℃.
7. The process of claim 6, wherein the dehydrogenation feed is an alkane and/or an arene.
8. The method of claim 7, wherein the alkane is selected from at least one of propane, n-butane, and isobutane; the aromatic hydrocarbon is selected from at least one of ethylbenzene, diethylbenzene and methyl ethylbenzene.
9. The method of any of claims 6-8, wherein the heater has a dehydrogenation feed inlet temperature of 20-500 ℃ and a dehydrogenation feed outlet temperature of 300-800 ℃; the inlet pressure of the dehydrogenation reaction raw material is 20-300 kPaA, and the outlet pressure of the dehydrogenation reaction raw material is 10-300 kPaA.
10. The method of any of claims 6-8, wherein the air preheating zone has a tube wall temperature of 300-750 ℃ in contact with the hydrocarbon feedstock.
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| US4067682A (en) * | 1975-08-01 | 1978-01-10 | Nichols Engineering & Research Corporation | Oil burner system |
| RU1789829C (en) * | 1990-05-11 | 1993-01-23 | Дзержинский Филиал Государственного Института По Проектированию Газоочистных Сооружений | Noncombustible waste thermal detoxication reactor |
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| CN106090943A (en) * | 2016-08-03 | 2016-11-09 | 王海升 | A kind of atomising burner of disposal of commercial danger waste liquid |
| WO2019155357A1 (en) * | 2018-02-07 | 2019-08-15 | Tenova S.P.A. | Industrial recuperative burner for industrial furnaces |
| CN210718071U (en) * | 2019-10-31 | 2020-06-09 | 江苏巨益机电有限公司 | Gas type air heater for utilizing waste heat |
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| Publication number | Priority date | Publication date | Assignee | Title |
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| US4067682A (en) * | 1975-08-01 | 1978-01-10 | Nichols Engineering & Research Corporation | Oil burner system |
| RU1789829C (en) * | 1990-05-11 | 1993-01-23 | Дзержинский Филиал Государственного Института По Проектированию Газоочистных Сооружений | Noncombustible waste thermal detoxication reactor |
| DE10304489A1 (en) * | 2002-04-11 | 2003-10-30 | Das Duennschicht Anlagen Sys | Device for cleaning exhaust gases with fluorine-containing compounds in a combustion reactor with low nitrogen oxide emissions |
| DE102004011955A1 (en) * | 2004-03-11 | 2005-09-29 | Werner Deppe | Waste gas and solvent vapor thermal recuperative cleaning system for drying equipment for rotary offset printer has burner at bottom of vertical tube with pipes carrying gas down toward flame |
| CN106090943A (en) * | 2016-08-03 | 2016-11-09 | 王海升 | A kind of atomising burner of disposal of commercial danger waste liquid |
| WO2019155357A1 (en) * | 2018-02-07 | 2019-08-15 | Tenova S.P.A. | Industrial recuperative burner for industrial furnaces |
| CN210718071U (en) * | 2019-10-31 | 2020-06-09 | 江苏巨益机电有限公司 | Gas type air heater for utilizing waste heat |
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