WO2022077458A1 - 一种焦调控反应器、由含氧化合物制备低碳烯烃的装置和应用 - Google Patents

一种焦调控反应器、由含氧化合物制备低碳烯烃的装置和应用 Download PDF

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WO2022077458A1
WO2022077458A1 PCT/CN2020/121567 CN2020121567W WO2022077458A1 WO 2022077458 A1 WO2022077458 A1 WO 2022077458A1 CN 2020121567 W CN2020121567 W CN 2020121567W WO 2022077458 A1 WO2022077458 A1 WO 2022077458A1
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Prior art keywords
gas
reactor
catalyst
coke
regenerator
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PCT/CN2020/121567
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English (en)
French (fr)
Inventor
张涛
叶茂
张今令
徐庶亮
唐海龙
王贤高
张骋
贾金明
王静
李华
李承功
刘中民
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中国科学院大连化学物理研究所
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Priority to JP2022573393A priority Critical patent/JP7449414B2/ja
Priority to EP20957236.1A priority patent/EP4082658A4/en
Priority to US17/801,811 priority patent/US20230085715A1/en
Priority to KR1020227043988A priority patent/KR20230012556A/ko
Priority to PCT/CN2020/121567 priority patent/WO2022077458A1/zh
Publication of WO2022077458A1 publication Critical patent/WO2022077458A1/zh

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    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J29/00—Catalysts comprising molecular sieves
    • B01J29/90—Regeneration or reactivation
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J29/00—Catalysts comprising molecular sieves
    • B01J29/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites
    • B01J29/06—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof
    • B01J29/70—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof of types characterised by their specific structure not provided for in groups B01J29/08 - B01J29/65
    • B01J29/7015—CHA-type, e.g. Chabazite, LZ-218
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
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    • B01J29/00—Catalysts comprising molecular sieves
    • B01J29/82—Phosphates
    • B01J29/84—Aluminophosphates containing other elements, e.g. metals, boron
    • B01J29/85—Silicoaluminophosphates [SAPO compounds]
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J38/00—Regeneration or reactivation of catalysts, in general
    • B01J38/04—Gas or vapour treating; Treating by using liquids vaporisable upon contacting spent catalyst
    • B01J38/06—Gas or vapour treating; Treating by using liquids vaporisable upon contacting spent catalyst using steam
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    • B01J38/00—Regeneration or reactivation of catalysts, in general
    • B01J38/04—Gas or vapour treating; Treating by using liquids vaporisable upon contacting spent catalyst
    • B01J38/12—Treating with free oxygen-containing gas
    • B—PERFORMING OPERATIONS; TRANSPORTING
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    • B01J38/00—Regeneration or reactivation of catalysts, in general
    • B01J38/04—Gas or vapour treating; Treating by using liquids vaporisable upon contacting spent catalyst
    • B01J38/12—Treating with free oxygen-containing gas
    • B01J38/14—Treating with free oxygen-containing gas with control of oxygen content in oxidation gas
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    • B01J8/26—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles according to "fluidised-bed" technique with two or more fluidised beds, e.g. reactor and regeneration installations
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    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
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    • B01J8/24—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles according to "fluidised-bed" technique
    • B01J8/26—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles according to "fluidised-bed" technique with two or more fluidised beds, e.g. reactor and regeneration installations
    • B01J8/28—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles according to "fluidised-bed" technique with two or more fluidised beds, e.g. reactor and regeneration installations the one above the other
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C1/00—Preparation of hydrocarbons from one or more compounds, none of them being a hydrocarbon
    • C07C1/20—Preparation of hydrocarbons from one or more compounds, none of them being a hydrocarbon starting from organic compounds containing only oxygen atoms as heteroatoms
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C11/00—Aliphatic unsaturated hydrocarbons
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    • B—PERFORMING OPERATIONS; TRANSPORTING
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    • B01J2208/00654—Controlling the process by measures relating to the particulate material
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    • B01J2208/00743—Feeding or discharging of solids
    • B01J2208/00761—Discharging
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
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    • B01J2208/0084—Stationary elements inside the bed, e.g. baffles
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    • C—CHEMISTRY; METALLURGY
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    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C2529/00—Catalysts comprising molecular sieves
    • C07C2529/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites, pillared clays
    • C07C2529/06—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof
    • C07C2529/70—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof of types characterised by their specific structure not provided for in groups C07C2529/08 - C07C2529/65
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
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    • C07C2529/82—Phosphates
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    • C07C2529/85—Silicoaluminophosphates (SAPO compounds)
    • 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
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    • Y02P20/52—Improvements relating to the production of bulk chemicals using catalysts, e.g. selective catalysts
    • 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
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    • Y02P30/40—Ethylene production

Definitions

  • the application relates to a coke control reactor, a device and application for preparing low-carbon olefins from oxygen-containing compounds, and belongs to the field of chemical catalytic devices.
  • MTO Methanol-to-olefins technology
  • DMTO DMTO technology of the Dalian Institute of Chemical Physics of the Chinese Academy of Sciences
  • MTO technology of the UOP Company of the United States In 2010, Shenhua Baotou MTO plant using DMTO technology was completed and put into operation. This is the world's first industrial application of MTO technology. By the end of 2019, 14 sets of DMTO industrial plants had been put into operation, with a total capacity of about 8 million tons of low-carbon olefins per year. .
  • DMTO technology has been further developed, and a new generation of DMTO catalysts with better performance have gradually begun to be industrially applied, creating higher benefits for DMTO factories.
  • the new generation of DMTO catalysts has higher methanol handling capacity and lower olefin selectivity.
  • Existing DMTO industrial devices are difficult to fully utilize the advantages of the new generation of DMTO catalysts. Therefore, it is necessary to develop a DMTO device and production method that can meet the needs of a new generation of DMTO catalysts with high methanol processing capacity and high low-carbon olefin selectivity.
  • a coke control reactor which can control the conversion and generation of coke species in the catalyst, and on the one hand convert the residual inactive macromolecular coke species in the regenerated catalyst into Small molecular coke species, on the other hand, the raw materials of the riser reactor and the bed reactor can also enter the catalyst to generate highly active small molecular coke species, and the small molecular coke species are polymethylbenzene and polymethylnaphthalene. Mainly, the selectivity of ethylene can be improved.
  • a coke control reactor includes a riser reactor and a bed reactor;
  • the bed reactor includes a bed reactor shell, and the bed reactor shell is surrounded from bottom to top Synthesis reaction zone I, transition zone and gas-solid separation zone I;
  • the inner and lower part of the reaction zone I is provided with a bed reactor distributor;
  • the outside of the reaction zone I is also provided with a coke control catalyst conveying pipe;
  • the riser pipe The bottom of the upper section of the reactor penetrates the bed reactor and is axially inserted into the bed reactor; the outlet end of the riser reactor is located in the transition zone.
  • the coke control catalyst conveying pipe is used for conveying the coke control catalyst to the next-stage reactor, such as a methanol conversion reactor.
  • the reaction zone I is provided with at least one perforated plate, and a plurality of the perforated plates are arranged in sequence on the outer periphery of the riser reactor in the axial direction; the outlet end of the riser reactor is located in the Above the perforated plate; the bed reactor distributor is located below the perforated plate.
  • the perforated plate is placed horizontally with the outer periphery of the perforated plate abutting against the inner wall of the bed reactor, so that the stream passes through the pores in the perforated plate.
  • the aperture ratio of the porous plate is 1-30%.
  • the gas-solid separation zone I is provided with a bed reactor gas-solid separator and a bed reactor gas collection chamber; the gas outlet of the bed reactor gas-solid separator is connected to the bed reactor gas-solid separator.
  • the gas collection chamber is communicated; the catalyst outlet of the gas-solid separator of the bed reactor is located above the perforated plate; the gas collection chamber of the bed reactor is communicated with the coke control product gas conveying pipe located outside the bed reactor .
  • the bed reactor gas-solid separator is a bed reactor gas-solid cyclone.
  • the inner top of the bed reactor is provided with a bed reactor gas collection chamber; the catalyst outlet of the gas-solid cyclone separator of the bed reactor is located below the outlet end of the riser reactor.
  • the bed reactor distributor is used for feeding the raw materials of the bed reactor;
  • the raw materials of the bed reactor include 0-20wt% hydrogen, 0-50wt% methane, 0-50wt% ethane, 0- 20wt% ethylene, 0-50wt% propane, 0-20wt% propylene, 0-90wt% butane, 0-90wt% butene, 0-90wt% pentane, 0-90wt% pentene, 0-90wt% hexane , 0-90wt% hexene, 0-50wt% methanol, 0-50wt% ethanol and 0-50wt% water, and the total content of methanol, ethanol and water ⁇ 10wt%.
  • the riser reactor is used for feeding catalyst and riser reaction raw materials;
  • the riser reaction raw materials include 0-20wt% hydrogen, 0-50wt% methane, 0-50wt% ethane, 0-20wt% Ethylene, 0-50wt% propane, 0-20wt% propylene, 0-90wt% butane, 0-90wt% butene, 0-90wt% pentane, 0-90wt% pentene, 0-90wt% hexane, 0 -90wt% hexene, 0-50wt% methanol, 0-50wt% ethanol and 0-50wt% water, and the total content of methanol, ethanol and water ⁇ 10wt%.
  • a device for preparing light olefins from oxygenates comprising a methanol conversion reactor and the coke control reactor described in any one of the above.
  • the light olefins mentioned in this application refer to ethylene and propylene.
  • the methanol conversion reactor includes a methanol conversion reactor shell and a conveying pipe;
  • the methanol conversion reactor shell includes a lower shell and an upper shell;
  • the lower shell encloses the synthesis reaction zone II, so
  • the inner and lower part of the reaction zone II is provided with a methanol conversion reactor distributor;
  • the conveying pipe is located above the reaction zone II in the axial direction, and one end of the conveying pipe is closed, and the other end is communicated with the reaction zone II
  • the upper casing is arranged on the outer circumference of the conveying pipe; the upper casing and the pipe wall of the conveying pipe are enclosed to form a cavity; the cavity is divided into a waiting agent area and a gas-solid separation from bottom to top Zone II;
  • the ready-to-generate agent zone is provided with a ready-to-generate zone gas distributor.
  • the upper casing is provided on the outer periphery of the conveying pipe in a wrapping form.
  • the methanol conversion reactor distributor is used for feeding the raw material containing oxygenates; the gas distributor in the zone to be generated is used for feeding the fluidizing gas in the zone to be generated.
  • the gas-solid separation zone II is provided with the first gas-solid separation device of the methanol conversion reactor; the upper part of the conveying pipe is connected to the inlet of the first gas-solid separation device of the methanol conversion reactor; the methanol
  • the catalyst outlet of the first gas-solid separation device of the reforming reactor is located in the to-be-generating agent zone; the gas outlet of the first gas-solid separation device of the methanol reforming reactor is communicated with the gas collecting chamber of the methanol reforming reactor; the methanol reforming reaction
  • the gas collecting chamber of the device is connected with the product gas delivery pipe.
  • the gas-solid separation zone II is further provided with a second gas-solid separation device of the methanol conversion reactor; the air inlet of the second gas-solid separation device of the methanol conversion reactor is located in the gas-solid separation zone II; the The catalyst outlet of the second gas-solid separation device of the methanol reforming reactor is located in the to-be-generating agent zone; the gas outlet of the second gas-solid separation device of the methanol reforming reactor is communicated with the gas collecting chamber of the methanol reforming reactor.
  • the gas distributor in the waiting area is located below the first gas-solid separation device of the methanol conversion reactor and the second gas-solid separation device of the methanol conversion reactor; the waiting area is also provided with a methanol conversion reactor. Heater.
  • the outside of the ready-to-generate agent zone is further provided with a ready-to-generate agent circulation pipe and a ready-to-generate inclined pipe;
  • the ready-to-generate agent circulation pipe is used to connect the ready-to-generate agent zone and the reaction zone II;
  • the The ready-to-be-grown inclined pipe is used to output the un-grown catalyst.
  • the circulating pipe of the unborn agent is used to transport the unborn catalyst in the unborn zone to the reaction zone II.
  • the unborn agent circulation pipe is provided with a unborn agent circulation slide valve.
  • the gas-solid separation zone II is communicated with the gas collection chamber of the bed reactor through a coke-regulated product gas delivery pipe; the reaction zone II is communicated with the reaction zone I through a coke-regulated catalyst delivery pipe.
  • a coke control catalyst slide valve is also provided on the coke control catalyst conveying pipe.
  • the device further comprises a regenerator; the regenerator is connected with the inclined pipe to be produced, and is used for transporting the catalyst to be produced into the regenerator; the regenerator is connected to the riser reactor and is used for The regenerated catalyst is transported into the coke control reactor; the inner bottom of the regenerator is provided with a regenerator distributor.
  • the regenerator distributor is used to pass the regeneration gas.
  • a regenerator stripper is also provided at the bottom of the regenerator; the upper section of the regenerator stripper is arranged inside the regenerator, and the inlet of the upper section of the regenerator stripper is located in the upper section of the regenerator stripper. above the regenerator distributor; the lower section of the regenerator stripper is arranged outside the regenerator, and the outlet of the lower section of the regenerator stripper is connected to the riser reactor; the regenerator There is also a regenerator heat extractor in the stripper.
  • the regenerator is connected with the inclined pipe to be produced through the conveying pipe of the regenerating agent and the stripper of the methanol conversion reactor; the regenerator is connected to the riser through the stripper of the regenerator and the inclined pipe of regeneration. Inlet connection of the reactor.
  • a spool valve to be grown is provided between the conveying pipe of the unregenerate agent and the stripper of the methanol conversion reactor; the inlet of the spool valve to be produced is connected to the bottom of the stripper of the methanol conversion reactor through a pipeline, and the outlet of the spool valve to be produced is connected to the bottom of the stripper of the methanol conversion reactor through a pipeline. It is connected to the inlet of the delivery pipe for the generation agent through a pipeline.
  • a regeneration slide valve is arranged between the regenerator stripper and the regeneration inclined pipe.
  • the inlet of the regeneration slide valve is connected to the bottom of the regenerator stripper through a pipeline, and the outlet of the regeneration slide valve is connected to the inlet of the regeneration inclined pipe through a pipeline.
  • the regenerator is further provided with a regenerator gas-solid separation device and a regenerator gas collection chamber; the catalyst outlet of the regenerator gas-solid separation device is located above the regenerator distributor; the regenerator The gas outlet of the gas-solid separation device is connected with the regenerator gas collection chamber; the regenerator gas collection chamber is connected with a flue gas conveying pipe located outside the regenerator.
  • the raw materials of the riser reactor and the catalyst are passed from the riser reactor into the transition zone, and the raw materials of the bed reactor are passed into the reaction zone I
  • the catalyst contacts and reacts with the raw materials of the riser reactor and the bed reactor to generate a coke control catalyst and a coke control product gas;
  • the catalyst is a DMTO catalyst;
  • the coke control catalyst is a modified DMTO catalyst.
  • the active component of the catalyst is SAPO-34 molecular sieve.
  • the catalyst entering the riser reactor may be a new catalyst or a regenerated catalyst, preferably a regenerated catalyst, so that coke regulation and catalyst regeneration can be simultaneously achieved online.
  • the catalyst is a regenerated catalyst; the coke content in the regenerated catalyst is ⁇ 3wt%.
  • the coke content in the coke control catalyst is 4-9 wt %.
  • the interquartile range of the coke content distribution in the coke control catalyst is less than 1 wt%.
  • the coke content of the coke control catalyst is 4-9 wt % through the setting of the coke control reactor and the selection of the coke control process. Since the catalyst is granular, the coke content of the catalyst refers to each The average coke content of each catalyst particle, but the coke content in each catalyst particle is actually different.
  • the interquartile difference of the coke content distribution in the coke control catalyst can be controlled within the range of less than 1 wt%, so that the overall coke content distribution of the catalyst is narrow, thereby improving the activity of the catalyst and the selectivity of light olefins.
  • the coke species in the coke control catalyst include polymethylbenzene and polymethylnaphthalene; the content of the mass of polymethylbenzene and polymethylnaphthalene in the total mass of the coke is ⁇ 70wt%; molecular weight>184 The content of the mass of coke species in the total mass of coke is ⁇ 25wt%; wherein, the total mass of coke refers to the total mass of coke species.
  • the type of coke species and the content of coke species are also very important, and are also one of the purposes of regulation in this application.
  • the effect that the content of polymethylbenzene and polymethylnaphthalene in the total mass of coke is ⁇ 70 wt% is achieved, the activity of the catalyst is improved, and the low carbon Olefin selectivity.
  • the process operating conditions of the riser reactor are: the gas superficial linear velocity is 3-10m/s, the reaction temperature is 400-700°C, the reaction pressure is 100-500kPa, and the bed density is 10-150kg/m 3 .
  • the process operating conditions of the reaction zone I of the bed reactor are: the gas superficial linear velocity is 0.1-1.0m/s, the reaction temperature is 300-650°C, the reaction pressure is 100-500kPa, and the bed density is 150-800kg/m 3 .
  • a method for preparing light olefins from an oxygen-containing compound comprising the method for on-line modification of a DMTO catalyst described in any one of the above.
  • the method further comprises: passing the coke control product gas into the gas-solid separation zone of the methanol conversion reactor; passing the coke control catalyst into the reaction zone II of the methanol conversion reactor.
  • the oxygenate-containing feedstock is contacted with the coke control catalyst to react to generate a stream A containing light olefins and the catalyst to be produced.
  • the stream A is divided into a gas-phase stream B and a solid-phase stream C; the gas-phase stream B enters the gas-collecting chamber of the methanol conversion reactor; The solid phase stream C enters the unborn agent zone; wherein, the gas phase stream B contains light olefins, and the solid phase stream C contains the unborn catalyst.
  • the fluidized gas in the area of the agent to be generated into the area of the agent to be generated; the fluidized gas in the area of the agent to be generated and the coke control product gas are mixed to carry a part of the catalyst to be generated to form a stream D;
  • Solid separation, after separation, a gas-phase stream E and a solid-phase stream F are obtained;
  • the gas-phase stream E enters the gas collection chamber of the methanol conversion reactor;
  • the solid-phase stream F enters the to-be-generating agent zone; wherein, the gas-phase stream E is a Mixed gas of fluidizing gas and coke control product gas in the raw agent zone;
  • the solid phase stream F is the as-grown catalyst.
  • the gas-phase stream B and the gas-phase stream E are mixed in the gas collection chamber of the methanol conversion reactor to form a product gas, and the product gas enters a downstream section through a product gas conveying pipe.
  • a part of the to-be-grown catalyst in the to-be-grown agent zone is returned to the bottom of the reaction zone II through the to-be-grown agent circulation pipe; the other part of the to-be-grown catalyst is discharged through the to-be-grown inclined pipe.
  • gas-solid separation is performed on the stream G; the separated flue gas enters the regenerator gas collection chamber, and then enters the downstream flue gas treatment system through the flue gas conveying pipe; the separated regenerated catalyst is stripped, Take heat, and then enter the coke control reactor.
  • the separated regenerated catalyst enters the riser reactor for stripping and heat extraction, and is passed into the bed reactor from the riser reactor.
  • the oxygenates include methanol and/or dimethyl ether.
  • the coke content in the as-grown catalyst is 9-13 wt%.
  • the fluidizing gas in the to-be-generating agent zone includes nitrogen and/or water vapor.
  • the regeneration gas includes 0-100wt% air, 0-50wt% oxygen, 0-50wt% nitrogen and 0-50wt% water vapor, and the air, oxygen, nitrogen and water vapor are not 0 at the same time.
  • the process operating conditions of the reaction zone II of the methanol conversion reactor are: the gas superficial linear velocity is 0.5-7.0m/s, the reaction temperature is 350-550°C, the reaction pressure is 100-500kPa, and the bed density is 100-500kg/m 3 .
  • the process operating conditions of the ismeerant zone of the methanol conversion reactor are: the gas superficial linear velocity is 0.1-1.0 m/s, the reaction temperature is 350-550°C, the reaction pressure is 100-500kPa, and the bed density is 200-800kg/m 3 .
  • the process operating conditions of the regenerator are: the gas superficial linear velocity is 0.5-2.0m/s, the regeneration temperature is 600-750°C, the regeneration pressure is 100-500kPa, and the bed density is 150-700kg/ m3 .
  • coke content refers to the mass ratio of coke species to coke control catalyst
  • C 4 -C 6 all represent the number of carbon atoms contained in the group, for example, C 4 -C 6 hydrocarbons represent hydrocarbons with 4-6 carbon atoms.
  • the production unit time consumption is expressed, and the mass of dimethyl ether in the oxygen-containing compound is equivalently converted into methanol mass based on the mass of element C, and the unit of production unit consumption is ton methanol/ton light olefin.
  • the production unit consumption is 2.50-2.60 tons of methanol/ton of light olefins.
  • the riser reactor is similar to the plug flow reactor. Therefore, by using the riser reactor to control the coke of the catalyst, a narrower coke content distribution can be obtained.
  • the residence time of the catalyst in the riser reactor is short, generally 1-20s. Therefore, it is difficult to greatly increase the coke content in the catalyst by only using the riser reactor to treat the regenerated catalyst.
  • the coke control reactor in this application includes a riser reactor and a bed reactor. On the one hand, the advantages of the riser reactor are used to obtain a catalyst with a narrower coke content distribution, and on the other hand, the bed reactor is used. Further increase the coke content in the catalyst and improve the selectivity of light olefins.
  • the main feature of the bed reactor in the present application is that the perforated plate is used to suppress the back-mixing of the catalyst between the beds and improve the uniformity of coke distribution in the catalyst.
  • the catalyst first enters the upper layer of the bed reactor from the riser reactor, gradually flows to the lower layer, and then enters the reaction zone II of the methanol conversion reactor from the lower layer.
  • the coke control reactor in the present application can control the transformation and generation of coke species in the catalyst, on the one hand, convert the residual inactive macromolecular coke species in the regenerated catalyst into small molecular coke species, on the other hand, improve the
  • the raw materials of the tube reactor and the bed reactor can also enter the catalyst to generate highly active small molecular coke species, and the small molecular coke species are mainly polymethylbenzene and polymethylnaphthalene, which can improve the selectivity of ethylene.
  • the method for on-line modification of DMTO catalyst by coke regulation and control reaction in the present application can obtain the coke regulation catalyst with high coke content, narrow coke content distribution, and the main components of coke species are polymethylbenzene and polymethylnaphthalene,
  • the regenerated catalyst with low selectivity of light olefins is converted into a coke control catalyst with high selectivity of light olefins.
  • the regenerated catalyst in the present application can also be directly used in the process of preparing low-carbon olefins from oxygenates without the coke control process, and the low-carbon olefin selectivity in the obtained product gas is 80 when the coke control process is not performed. -83wt%.
  • the regenerated catalyst in the present application is processed by the coke control process and then used in the process of preparing light olefins from oxygenated compounds, and the selectivity of light olefins in the obtained product gas is 92-96 wt%.
  • the methanol conversion reactor in this application adopts a composite fluidized bed reactor comprising a fast fluidized bed area and a bubbling fluidized bed area; the fast fluidized bed area is a reaction area, which can obtain higher
  • the methanol flux increases the methanol processing capacity per unit volume of the equipment, and the methanol mass space velocity can reach 5-20h -1 .
  • the bubbling fluidized bed area is the area of the waiting agent, which is used to take heat, reduce the temperature of the waiting catalyst, and reduce the temperature of the catalyst to be formed.
  • the low-temperature waiting catalyst is transported to the reaction zone, the bed density of the reaction zone is increased, and the bed temperature of the reaction zone is controlled.
  • the gas superficial linear velocity is 0.5-7.0m/s
  • the corresponding bed density is 500- 100kg/m 3 .
  • the methanol conversion reactor in this application adopts the structure that the first gas-solid separation equipment of the methanol conversion reactor is directly connected to the conveying pipe, and realizes the rapid separation of the gas containing low-carbon olefins and the catalyst to be produced in the stream A, It is avoided that the lower olefins are further reacted under the action of the catalyst to be generated to generate hydrocarbon by-products with larger molecular weights.
  • FIG. 1 is a schematic structural diagram of an oxygenate-to-light olefin (DMTO) device according to an embodiment of the present application.
  • DMTO oxygenate-to-light olefin
  • a key characteristic of DMTO catalysts is that the selectivity to lower olefins of the methanol conversion process increases as the coke content of the catalyst increases.
  • the light olefins mentioned in this application refer to ethylene and propylene.
  • the applicant's research has found that the main factors affecting the activity of DMTO catalysts and the selectivity of light olefins are the coke content, coke content distribution and coke species in the catalyst.
  • the coke species in the catalyst include polymethyl aromatic hydrocarbons and polymethyl naphthenes, among which polymethylbenzene and polymethylnaphthalene can promote the formation of ethylene. Therefore, controlling the coke content, coke content distribution and coke species in the catalyst is the key to control the activity of DMTO catalyst and improve the selectivity of light olefins.
  • the present application provides a method for on-line modification of the DMTO catalyst by coke control reaction, the steps comprising:
  • a coke conditioning feedstock comprising hydrogen, methane, ethane, ethylene, propane, propylene, butane, butene, pentane, pentene, hexane, hexene, methanol, ethanol and water to a coke conditioning reactor ;
  • coke control material and the regenerated catalyst contact and react in the coke control reactor, the coke control material is coked on the regenerated catalyst, and the coked catalyst is called a coke control catalyst, and the coke content in the coke control catalyst is 4- 9wt%, the interquartile difference of the coke content distribution is less than 1wt%, the coke species contains polymethylbenzene and polymethylnaphthalene, and the content of the mass of polymethylbenzene and polymethylnaphthalene in the total mass of coke is ⁇ 70wt% %, the content of the mass of coke species with molecular weight>184 in the total mass of coke is ⁇ 25wt%;
  • the regenerated catalyst is a DMTO catalyst with a coke content of ⁇ 3wt%, and the active component of the DMTO catalyst is SAPO-34 molecular sieve.
  • the composition of the coke control raw material is 0-20wt% hydrogen, 0-50wt% methane, 0-50wt% ethane, 0-20wt% ethylene, 0-50wt% propane, 0-20wt% propylene, 0-90wt% butane alkane, 0-90wt% butene, 0-90wt% pentane, 0-90wt% pentene, 0-90wt% hexane, 0-90wt% hexene, 0-50wt% methanol, 0-50wt% ethanol and 0 -50wt% water, and the total content of methanol, ethanol and water is ⁇ 10wt%.
  • the reaction temperature of the coke control reaction is 300-700°C.
  • both the bed reactor raw material and the riser reaction raw material are coke control raw materials.
  • a method for preparing light olefins from oxygenated compounds comprising the above-mentioned method for on-line modification of a DMTO catalyst by a coke control reaction and a device therefor.
  • the apparatus comprises a coke control reactor (1), a methanol conversion reactor (2) and a regenerator (3).
  • a coke control reactor (1) for on-line modified DMTO catalyst comprises a riser reactor (1-1) and a bed reactor (1-2), the bed
  • the reactor (1-2) comprises: a bed reactor shell (1-3), a bed reactor distributor (1-4), a perforated plate (1-5), a bed reactor gas-solid cyclone separator (1-6), bed reactor gas collection chamber (1-7), coke control product gas delivery pipe (1-8), coke control catalyst delivery pipe (1-9) and coke control catalyst slide valve (1- 10);
  • the bed reactor (1-2) is divided into a reaction zone, a transition zone and a gas-solid separation zone from bottom to top.
  • the riser reactor (1-1) passes through the bed reactor shell (1-3), is partially located below the bed reactor (1-2), and partially located in the bed reactor (1-2) .
  • the bed reactor distributor (1-4) is located at the bottom of the reaction zone, and n pieces of perforated plates (1-5) are arranged in the reaction zone, 0 ⁇ n ⁇ 9;
  • the bed reactor gas-solid cyclone (1-6) is located in the gas-solid separation zone, the inlet of the bed reactor gas-solid cyclone (1-6) is located in the gas-solid separation zone, and the bed reactor gas-solid separation zone
  • the catalyst outlet of the cyclone separator (1-6) is located in the reaction zone, and the gas outlet of the gas-solid cyclone separator (1-6) of the bed reactor is connected to the gas collecting chamber (1-7) of the bed reactor;
  • the bed reactor gas collection chamber (1-7) is located at the top of the bed reactor (1-2); the inlet of the coke control product gas conveying pipe (1-8) is connected to the bed reactor gas collection chamber (1-8). 1-7), the outlet of coke regulating and controlling product gas conveying pipe (1-8) is connected to the top of methanol conversion reactor (2);
  • a coke control catalyst spool valve (1-10) is arranged in the coke control catalyst delivery pipe (1-9); the inlet of the coke control catalyst delivery pipe (1-9) is connected to the lower part of the reaction zone, and the coke control catalyst delivery pipe (1-9) is connected to the lower part of the reaction zone.
  • the outlet of 1-9) is connected to the lower part of the methanol conversion reactor (2).
  • the aperture ratio of the perforated plate is 1-30%.
  • a methanol conversion reactor (2) for methanol conversion to prepare low-carbon olefins comprising: a methanol conversion reactor shell (2-1), a methanol conversion reactor distributor (2) -2), the conveying pipe (2-3), the first gas-solid separation device (2-4) of the methanol conversion reactor, the gas collecting chamber (2-5) of the methanol conversion reactor, the gas distributor (2-5) in the area to be generated -6), methanol conversion reactor heat extractor (2-7), methanol conversion reactor second gas-solid separation equipment (2-8), product gas conveying pipe (2-9), waiting for generating agent circulation pipe (2 -10), the recycle slide valve for the unborn agent (2-11), the inclined pipe for the unborn (2-12), the methanol conversion reactor stripper (2-13), the unborn spool valve (2-14) and the unborn spool valve (2-14) Raw material delivery tube (2-15);
  • the lower part of the methanol conversion reactor (2) is the reaction zone, the middle part is the generating agent zone, and the upper part is the gas-solid separation zone.
  • the methanol conversion reactor distributor (2-2) is located at the bottom of the reaction zone of the methanol conversion reactor (2), the conveying pipe (2-3) is located in the central area of the middle and upper part of the methanol conversion reactor (2), and the conveying pipe ( The bottom end of 2-3) is connected to the top of the reaction zone, and the top of the conveying pipe (2-3) is connected to the inlet of the first gas-solid separation device (2-4) of the methanol conversion reactor, and the first gas-solid separation of the methanol conversion reactor.
  • the solid separation device (2-4) is located in the gas-solid separation zone of the methanol conversion reactor (2), and the gas outlet of the first gas-solid separation device (2-4) of the methanol conversion reactor is connected to the methanol conversion reactor gas collection chamber ( 2-5), the catalyst outlet of the first gas-solid separation device (2-4) of the methanol conversion
  • the gas distributor (2-6) in the waiting agent area is located at the bottom of the ungrowing agent area, the methanol conversion reactor heat extractor (2-7) is located in the waiting agent area, and the second gas-solid separation device of the methanol conversion reactor (2-8) is located in the gas-solid separation zone of methanol conversion reactor (2), and the inlet of the second gas-solid separation device (2-8) of methanol conversion reactor is positioned at the gas-solid separation zone of methanol conversion reactor (2),
  • the gas outlet of the methanol conversion reactor second gas-solid separation device (2-8) is connected to the methanol conversion reactor gas collection chamber (2-5), and the catalyst of the methanol conversion reactor second gas-solid separation device (2-8)
  • the outlet is located in the area to be generated, the methanol conversion reactor gas collection chamber (2-5) is located at the top of the methanol conversion reactor (2), and the product gas delivery pipe (2-9) is connected to the methanol conversion reactor gas collection chamber (2). -5) at the top;
  • the inlet of the ungenerating agent circulating pipe (2-10) is connected to the ungenerating agent zone, and the outlet of the ungenerating agent circulating pipe (2-10) is connected to the bottom of the reaction zone of the methanol conversion reactor (2).
  • a circulating slide valve (2-11) for the unborn agent is arranged in the agent circulation pipe (2-10), and the outlet of the coke control catalyst delivery pipe (1-9) is connected to the bottom of the reaction zone of the methanol conversion reactor (2).
  • the inlet of the inclined pipe (2-12) is connected to the area of the waiting agent, and the outlet of the inclined pipe (2-12) is connected to the upper part of the stripper (2-13) of the methanol conversion reactor.
  • the stripper (2-13) is placed outside the methanol conversion reactor shell (2-1), and the inlet of the to-be-grown slide valve (2-14) is connected to the methanol conversion reactor stripper (2-13) through a pipeline
  • the outlet of the spool valve (2-14) to be regenerated is connected to the inlet of the regenerated agent delivery pipe (2-15) through a pipeline, and the outlet of the regenerated agent delivery pipe (2-15) is connected to the regenerator (3) the middle of.
  • the first gas-solid separation device (2-4) of the methanol conversion reactor adopts one or more groups of gas-solid cyclones, and each group of gas-solid cyclones includes a first-stage gas-solid cyclone separation and a second-stage gas-solid cyclone.
  • the second gas-solid separation device (2-8) of the methanol conversion reactor adopts one or more sets of gas-solid cyclone separators, each set of gas-solid cyclone separators includes a first-stage gas-solid cyclone separator and a second-stage gas-solid cyclone.
  • the methanol conversion reactor (2) belongs to a fluidized bed reactor.
  • regenerator (3) for regenerating a catalyst, the regenerator (3) comprising: a regenerator housing (3-1), a regenerator distributor (3-2), and a regenerator gas-solid separation device (3) -3), regenerator gas collection chamber (3-4), flue gas delivery pipe (3-5), regenerator stripper (3-6), regenerator heat extractor (3-7), regeneration slide valve (3-8) and regeneration inclined pipe (3-9);
  • the regenerator distributor (3-2) is located at the bottom of the regenerator (3), the regenerator gas-solid separation device (3-3) is located at the upper part of the regenerator (3), and the regenerator gas-solid separation device (3-3) ) is located at the upper part of the regenerator (3), the gas outlet of the regenerator gas-solid separation device (3-3) is connected to the regenerator gas collection chamber (3-4), and the regenerator gas-solid separation device (3-3)
  • the catalyst outlet is located at the lower part of the regenerator (3)
  • the regenerator gas collection chamber (3-4) is located at the top of the regenerator (3)
  • the flue gas delivery pipe (3-5) is connected to the regenerator gas collection chamber (3- 4) at the top;
  • the regenerator stripper (3-6) is located outside the regenerator housing (3-1), and the inlet pipe of the regenerator stripper (3-6) penetrates the regenerator housing (3-1), The opening is above the regenerator distributor (3-2), the regenerator heat collector (3-7) is located in the regenerator stripper (3-6), and the inlet of the regeneration slide valve (3-8) is through the pipeline Connected to the bottom of the regenerator stripper (3-6), the outlet of the regeneration slide valve (3-8) is connected to the inlet of the regeneration inclined pipe (3-9) through the pipeline, and the outlet of the regeneration inclined pipe (3-9) Connected to the bottom of the riser reactor (1-1).
  • the regenerator gas-solid separation device (3-3) adopts one or more groups of gas-solid cyclones, each group of gas-solid cyclones comprises a first-stage gas-solid cyclone and a first-stage gas-solid cyclone Secondary gas-solid cyclone separator.
  • the regenerator (3) belongs to a fluidized bed reactor.
  • the inner diameter of the transition zone of the coke control reactor gradually increases from bottom to top; in the methanol conversion reactor, the inner diameter of the connection between the reaction zone II and the conveying pipe gradually decreases from bottom to top; The inner diameter of the junction of the zones increases gradually from bottom to top
  • the inlet of the coke control catalyst conveying pipe is located above the distributor of the bed reactor; the outlet of the coke control catalyst conveying pipe is located above the distributor of the methanol conversion reactor; Above the gas distributor, the outlet of the circulating pipe for the generation agent is located above the distributor of the methanol conversion reactor; the inclined pipe to be generated is located above the gas distributor in the area for the generation agent.
  • a methanol-to-olefin method comprising a method for on-line modification of a DMTO catalyst by a coke control reaction, comprising the following steps:
  • logistics A logistics A enters methanol conversion reactor first gas-solid separation equipment (2-4) through conveying pipe (2-3), after gas-solid separation, it is divided into gas-phase logistics B and solid-phase logistics C, and gas-phase logistics B is The gas containing low-carbon olefins, the solid phase stream C is the catalyst to be generated, the gas phase stream B enters the gas collection chamber (2-5) of the methanol conversion reactor, and the solid phase stream C enters the area to be generated agent; fluidize the area to be generated agent The gas is passed from the gas distributor (2-6) of the unborn agent zone into the unborn agent zone, and contacts with the unborn catalyst.
  • the fluidized gas in the unborn agent zone and the coke-regulated product gas are mixed to carry part of the unborn catalyst to form a stream D.
  • D enters the second gas-solid separation device (2-8) of the methanol conversion reactor, and after the gas-solid separation, it is divided into gas-phase stream E and solid-phase stream F, and gas-phase stream E is the fluidized gas in the waiting agent zone and the coke-regulated product gas
  • the mixed gas, the solid phase stream F is the catalyst to be generated, the gas phase stream E enters the gas collection chamber (2-5) of the methanol conversion reactor, and the solid phase stream F enters the zone to be generated agent;
  • the product gas is formed by mixing in the gas collecting chamber (2-5) of the reactor, and the product gas enters the downstream section through the product gas conveying pipe (2-9); a part of the catalyst to be generated in the unborn agent area passes through the unborn agent circulation pipe (2-9).
  • the flue gas After gas-solid separation, it is divided into flue gas and regenerated catalyst, and the flue gas enters the regenerator gas collection chamber (3-4), then enter the downstream flue gas treatment system through the flue gas delivery pipe (3-5), the regenerated catalyst returns to the bottom of the regenerator (3), and the regenerated catalyst in the regenerator (3) enters the regenerator for stripping After stripping and taking heat, the reactor (3-6) enters the coke control reactor (1) through the regeneration slide valve (3-8) and the regeneration inclined pipe (3-9).
  • the method described herein is carried out using the above-described apparatus comprising a coke control reactor (1), a methanol conversion reactor (2) and a regenerator (3).
  • the composition of the raw material of the riser reactor in the method is 0-20wt% hydrogen, 0-50wt% methane, 0-50wt% ethane, 0-20wt% ethylene, 0-50wt% Propane, 0-20wt% propylene, 0-90wt% butane, 0-90wt% butene, 0-90wt% pentane, 0-90wt% pentene, 0-90wt% hexane, 0-90wt% hexene, 0-50wt% methanol, 0-50wt% ethanol and 0-50wt% water, and the total content of methanol, ethanol and water ⁇ 10wt%.
  • the composition of the bed reactor feedstock in the method is 0-20wt% hydrogen, 0-50wt% methane, 0-50wt% ethane, 0-20wt% ethylene, 0-50wt% Propane, 0-20wt% propylene, 0-90wt% butane, 0-90wt% butene, 0-90wt% pentane, 0-90wt% pentene, 0-90wt% hexane, 0-90wt% hexene, 0-50wt% methanol, 0-50wt% ethanol and 0-50wt% water, and the total content of methanol, ethanol and water ⁇ 10wt%.
  • the oxygenate in the method is one of methanol or dimethyl ether or a mixture of methanol and dimethyl ether.
  • the fluidizing gas in the to-be-generating agent zone in the method is one of nitrogen and water vapor or a mixture of nitrogen and water vapor.
  • the regeneration gas in the method is 0-100wt% air, 0-50wt% oxygen, 0-50wt% nitrogen and 0-50wt% water vapor.
  • the active component of the catalyst is SAPO-34 molecular sieve.
  • the coke content in the regenerated catalyst is ⁇ 3wt%.
  • the coke content in the coke control catalyst is 4-9 wt%
  • the interquartile difference of the coke content distribution is less than 1 wt%
  • the coke species include polymethylbenzene and polymethylnaphthalene
  • the content of the mass of benzene and polymethylnaphthalene in the total mass of the coke is ⁇ 70 wt%
  • the content of the mass of the coke species with molecular weight>184 in the total mass of the coke is ⁇ 25 wt%.
  • the coke content in the to-be-grown catalyst is 9-13 wt %, and further preferably, the coke content in the un-grown catalyst is 10-12 wt %.
  • the process operating conditions of the riser reactor (1-1) are: the gas superficial linear velocity is 3-10m/s, the reaction temperature is 400-700°C, and the reaction pressure is 100-500kPa , the bed density is 10-150kg/m 3 .
  • the process operating conditions of the reaction zone of the bed reactor (1-2) are: the gas superficial linear velocity is 0.1-1.0 m/s, the reaction temperature is 300-650°C, and the reaction pressure is It is 100-500kPa, and the bed density is 150-800kg/m 3 .
  • the process operating conditions of the reaction zone of the methanol conversion reactor (2) are: the gas superficial linear velocity is 0.5-7.0m/s, the reaction temperature is 350-550°C, and the reaction pressure is 100 -500kPa, the bed density is 100-500kg/m 3 .
  • the process operating conditions of the to-be-generating agent zone of the methanol conversion reactor (2) are: the gas superficial linear velocity is 0.1-1.0 m/s, the reaction temperature is 350-550°C, and the reaction pressure is is 100-500kPa, and the bed density is 200-800kg/m 3 .
  • the process operating conditions of the regenerator (3) are: the gas superficial linear velocity is 0.5-2.0m/s, the regeneration temperature is 600-750°C, the regeneration pressure is 100-500kPa, the bed layer The density is 150-700kg/m 3 .
  • the composition of the product gas is 40-55wt% ethylene, 37-53wt% propylene, ⁇ 4wt% C4 - C6 hydrocarbons and ⁇ 4wt% other components, and the other components are methane, Ethane, propane, hydrogen, CO and CO2 , etc., and the total selectivity of ethylene and propylene in the product gas is 92-96 wt%.
  • the apparatus shown in FIG. 1 is adopted, and the perforated plate is not included in the bed reactor.
  • the raw material of the riser reactor is a mixture of 6wt% butane, 81wt% butene, 2wt% methanol and 11wt% water;
  • the bed reactor raw material is 6wt% butane, 81wt% butene, 2wt% methanol and 11wt% water;
  • the oxygenate is methanol;
  • the fluidizing gas in the regeneration zone is nitrogen;
  • the regeneration gas is air;
  • the active component in the catalyst is SAPO-34 molecular sieve;
  • the coke content in the regenerated catalyst is about 3wt% ;
  • the coke content in the coke control catalyst is about 6wt%, of which the mass of polymethylbenzene and polymethylnaphthalene is about 82wt% in the total mass of coke, and the mass of coke species with molecular weight>184 is in the total mass of coke
  • the content in the coke control catalyst is about 5wt%;
  • the mass space velocity of oxygenates in the methanol conversion reactor is about 20 h -1 ;
  • the composition of the product gas is 55 wt% ethylene, 37 wt% propylene, 4 wt% C 4 -C 6 hydrocarbons and 4 wt % other
  • the other components are methane, ethane, propane, hydrogen, CO and CO 2 , etc.;
  • the production unit consumption is 2.60 tons of methanol/ton of light olefins.
  • This embodiment adopts the device shown in FIG. 1 , and the bed reactor includes 4 perforated plates, and the opening rate of the perforated plates is 30%.
  • the riser reactor feedstock is a mixture of 22wt% methane, 24wt% ethane, 3wt% ethylene, 28wt% propane, 4wt% propylene, 7wt% hydrogen and 12wt% water;
  • the bed reactor feedstock is 22wt% Mixture of methane, 24wt% ethane, 3wt% ethylene, 28wt% propane, 4wt% propylene, 7wt% hydrogen and 12wt% water; oxygenates were 82wt% methanol and 18wt% dimethyl ether; is water vapor;
  • the regeneration gas is 50wt% air and 50wt% water vapor;
  • the active component in the catalyst is SAPO-34 molecular sieve;
  • the coke content in the regenerated catalyst is about 1wt%;
  • the coke content in the coke control catalyst is about 4wt% , wherein the content of the mass of polymethylbenzene and polymethylnaphthalene in the total mass of coke
  • the mass space velocity of oxygenates in the methanol conversion reactor is about 5h -1 ;
  • the composition of the product gas is 40wt% ethylene, 53wt% propylene, 4wt% C4 - C6 hydrocarbons and 3wt% other
  • the other components are methane, ethane, propane, hydrogen, CO and CO 2 , etc.;
  • the production unit consumption is 2.58 tons of methanol/ton of light olefins.
  • the device shown in FIG. 1 is used, and the bed reactor includes 4 perforated plates, and the opening rate of the perforated plates is 1%.
  • the riser reactor feedstock is 1wt% propane, 1wt% propylene, 3wt% butane, 51wt% butene, 3wt% pentane, 22wt% pentene, 1wt% hexane, 7wt% hexene, 2wt% Mixture of % methanol and 9 wt % water; bed reactor feed is 1 wt % propane, 1 wt % propylene, 3 wt % butane, 51 wt % butene, 3 wt % pentane, 22 wt % pentene, 1 wt % hexane, 7 wt % A mixture of hexene, 2wt% methanol and 9wt% water; the oxygenate is dimethyl ether; the fluidizing gas in the regeneration zone is 5wt% nitrogen and 95wt% water vapor; the regeneration gas is 50wt% air and 50wt% oxygen; catalyst The active component
  • the content in the mass is about 79wt%, and the mass of the coke species with molecular weight>184 is about 10wt% in the total mass of the coke; the interquartile difference of the coke content distribution in the coke control catalyst is about 0.5wt%;
  • the coke content in the catalyst is about 11wt%;
  • the process operating conditions of the riser reactor (1-1) are: the gas superficial linear velocity is 6.0m/s, the reaction temperature is 600°C, the reaction pressure is 300kPa, the bed density is The process operating conditions of the reaction zone of the bed reactor (1-2) are: the gas superficial linear velocity is about 0.4m/s, the reaction temperature is about 550°C, the reaction pressure is about 300kPa, the bed The layer density is about 500kg/m 3 ;
  • the process operating conditions of the reaction zone of the methanol conversion reactor (2) are: the gas superficial linear velocity is about 3.0m/s, the reaction temperature is about 450°C, and the reaction pressure is about 300kPa,
  • the mass space velocity of oxygenates in the methanol conversion reactor is about 15h -1 ;
  • the composition of the product gas is 50wt% ethylene, 45wt% propylene, 3wt% C4 - C6 hydrocarbons and 2wt% other
  • the other components are methane, ethane, propane, hydrogen, CO and CO 2 , etc.;
  • the production unit consumption is 2.53 tons of methanol/ton of light olefins.
  • This embodiment adopts the device shown in FIG. 1 , and the bed reactor includes 9 perforated plates, and the opening ratio of the perforated plates is 5%.
  • the raw material of the riser reactor is a mixture of 5wt% butane, 72wt% butene, 8wt% methanol and 15wt% water;
  • the bed reactor raw material is 5wt% butane, 72wt% butene, 8wt% methanol and 15wt% water;
  • the oxygenate is methanol;
  • the fluidizing gas in the regeneration zone is 73wt% nitrogen and 27wt% water vapor;
  • the regeneration gas is 50wt% air and 50wt% nitrogen;
  • the active component in the catalyst is SAPO- 34 molecular sieve;
  • the coke content in the regenerated catalyst is about 2wt%;
  • the coke content in the coke control catalyst is about 9wt%, wherein the content of polymethylbenzene and polymethylnaphthalene in the total coke mass is about 71wt% , the mass of coke species with molecular weight>184 is about 25wt% in the total mass of coke; the inter
  • the mass space velocity of oxygenates in the methanol conversion reactor is about 11 h -1 ;
  • the composition of the product gas is 52 wt% ethylene, 44 wt% propylene, 2 wt% C 4 -C 6 hydrocarbons and 2 wt % other
  • the other components are methane, ethane, propane, hydrogen, CO and CO 2 , etc.;
  • the production unit consumption is 2.50 tons of methanol/ton of light olefins.
  • This example is a comparative case, and the difference from Example 4 is that the coke control reaction is not used to modify the DMTO catalyst on-line.
  • the raw material introduced into the riser reactor and the bed reactor is nitrogen, which is an inert gas and will not increase
  • the properties of the regenerated catalyst are changed in the tube reactor and the bed reactor, ie, the catalyst equivalent to entering the reaction zone II of the methanol conversion reactor is the regenerated catalyst.
  • the composition of the product gas is 46wt% ethylene, 37wt% propylene, 12wt% C4-C6 hydrocarbons and 5wt% other components, the other components are methane, ethane, propane, hydrogen, CO and CO2 etc.; the production unit consumption is 2.90 tons of methanol/ton of light olefins.
  • This comparative case shows that the online modification of DMTO catalyst through coke control reaction can greatly improve the performance of the catalyst and reduce the production unit consumption.

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Abstract

一种焦调控反应器、由含氧化合物制备低碳烯烃的装置和应用。该焦调控反应器包括提升管反应器和床层反应器；床层反应器包括床层反应器壳体，床层反应器壳体由下至上围合成反应区Ⅰ、过渡区和气固分离区Ⅰ；反应区Ⅰ设有床层反应器分布器；反应区Ⅰ的外部还设有焦调控催化剂输送管；提升管反应器的上段穿透床层反应器的底部沿轴向插设在床层反应器中；提升管反应器的出口端位于过渡区。焦调控反应器可以控制催化剂中的焦物种的转化和生成，一方面将再生催化剂中残留的非活性的大分子焦物种转化为小分子焦物种，另一方面，提升管反应器原料和床层反应器原料还可进入催化剂中生成高活性的小分子焦物种，提高乙烯的选择性。

Description

一种焦调控反应器、由含氧化合物制备低碳烯烃的装置和应用 技术领域
本申请涉及一种焦调控反应器、由含氧化合物制备低碳烯烃的装置和应用，属于化工催化装置领域。
背景技术
甲醇制烯烃技术(MTO)主要有中国科学院大连化学物理研究所的DMTO技术和美国UOP公司的MTO技术。2010年，采用DMTO技术的神华包头甲醇制烯烃工厂建成投产，此为MTO技术的全球首次工业化应用，截至2019年底，已有14套DMTO工业装置投产，低碳烯烃产能共计约800万吨/年。
最近几年，DMTO技术进一步发展，性能更加优良的新一代DMTO催化剂逐渐开始工业化应用，为DMTO工厂创造了更高的效益。新一代DMTO催化剂具有更高的甲醇处理能力和低碳烯烃选择性。现有的DMTO工业装置难以充分利用新一代DMTO催化剂的优势，因此，需要开发出一种可以适应高甲醇处理能力、高低碳烯烃选择性的新一代DMTO催化剂需求的DMTO装置及生产方法。
发明内容
根据本申请的一个方面，提供了一种焦调控反应器，该焦调控反应器可以控制催化剂中的焦物种的转化和生成，一方面将再生催化剂中残留的非活性的大分子焦物种转化为小分子焦物种，另一方面，提升管反应器原料和床层反应器原料还可进入催化剂中生成高活性的小分子焦物种，并且小分子焦物种以多甲基苯和多甲基萘为主，可以提高乙烯的选择性。
一种焦调控反应器，所述焦调控反应器包括提升管反应器和床层反应器；所述床层反应器包括床层反应器壳体，所述床层反应器壳体由下至上围合成反应区Ⅰ、过渡区和气固分离区Ⅰ；所述反应区Ⅰ的内下部设有床层反应器分布器；所述反应区Ⅰ的外部还设有焦调控催化剂输送管；所述提升管反应器的上段穿透床层反应器的底部沿轴向插设在床层反应器中；所述提升管反应器的出口端位于所述过渡区。
具体地，焦调控催化剂输送管用于将焦调控催化剂输送至下一级反应器中，例如甲醇转化反应器。
可选地，所述反应区Ⅰ内设有至少一个多孔板，多个所述多孔板沿轴向依次排列在所述提升管反应器的外周；所述提升管反应器的出口端位于所述多孔板的上方；所述床层反应器分布器位于所述多孔板的下方。
具体地，多孔板水平放置，且所述多孔板的外周与床层反应器的内壁抵顶，使得物流从多孔板中 的孔隙穿过。
可选地，所述多孔板的开孔率为1-30％。
可选地，所述气固分离区Ⅰ设有床层反应器气固分离器和床层反应器集气室；所述床层反应器气固分离器的气体出口与所述床层反应器集气室连通；所述床层反应器气固分离器的催化剂出口位于多孔板的上方；所述床层反应器集气室与位于所述床层反应器外部的焦调控产品气输送管连通。
具体地，床层反应器气固分离器为床层反应器气固旋风分离器。
具体地，床层反应器的内顶部设有床层反应器集气室；床层反应器气固旋风分离器的催化剂出口位于所述提升管反应器的出口端的下方。
可选地，所述床层反应器分布器用于通入床层反应器原料；所述床层反应器原料包括0-20wt％氢气、0-50wt％甲烷、0-50wt％乙烷、0-20wt％乙烯、0-50wt％丙烷、0-20wt％丙烯、0-90wt％丁烷、0-90wt％丁烯、0-90wt％戊烷、0-90wt％戊烯、0-90wt％己烷、0-90wt％己烯、0-50wt％甲醇、0-50wt％乙醇和0-50wt％水，并且，甲醇、乙醇和水的总含量≥10wt％。
可选地，所述提升管反应器用于通入催化剂和提升管反应原料；所述提升管反应原料包括0-20wt％氢气、0-50wt％甲烷、0-50wt％乙烷、0-20wt％乙烯、0-50wt％丙烷、0-20wt％丙烯、0-90wt％丁烷、0-90wt％丁烯、0-90wt％戊烷、0-90wt％戊烯、0-90wt％己烷、0-90wt％己烯、0-50wt％甲醇、0-50wt％乙醇和0-50wt％水，并且，甲醇、乙醇和水的总含量≥10wt％。
根据本申请的另一方面，还提供了一种由含氧化合物制备低碳烯烃的装置，所述装置包括甲醇转化反应器和上述任一项所述的焦调控反应器。
本申请中所述的低碳烯烃是指乙烯和丙烯。
可选地，所述甲醇转化反应器包括甲醇转化反应器壳体和输送管；所述甲醇转化反应器壳体包括下壳体和上壳体；所述下壳体围合成反应区Ⅱ，所述反应区Ⅱ的内下部设有甲醇转化反应器分布器；所述输送管沿轴向位于所述反应区Ⅱ的上方，且所述输送管的一端闭合，另一端与所述反应区Ⅱ连通；所述上壳体设置在所述输送管的外周；所述上壳体与所述输送管的管壁围合形成空腔；所述空腔自下至上分为待生剂区和气固分离区Ⅱ；所述待生剂区设有待生区气体分布器。
具体地，上壳体以包覆的形式设置在输送管的外周。甲醇转化反应器分布器用于通入含有含氧化合物的原料；待生区气体分布器用于通入待生剂区流化气体。
可选地，所述气固分离区Ⅱ设有甲醇转化反应器第一气固分离设备；所述输送管的上部与所述甲醇转化反应器第一气固分离设备的入口连接；所述甲醇转化反应器第一气固分离设备的催化剂出口位于所述待生剂区；所述甲醇转化反应器第一气固分离设备的气体出口与甲醇转化反应器集气室连通；所述甲醇转化反应器集气室与产品气输送管连接。
可选地，所述气固分离区Ⅱ还设有甲醇转化反应器第二气固分离设备；所述甲醇转化反应器第二气固分离设备的进气口位于气固分离区Ⅱ；所述甲醇转化反应器第二气固分离设备的催化剂出口位于所述待生剂区；所述甲醇转化反应器第二气固分离设备的气体出口与甲醇转化反应器集气室连通。
可选地，所述待生区气体分布器位于甲醇转化反应器第一气固分离设备和甲醇转化反应器第二气固分离设备的下方；所述待生剂区还设有甲醇转化反应器取热器。
可选地，所述待生剂区的外部还设有待生剂循环管和待生斜管；所述待生剂循环管用于将所述待生剂区和所述反应区Ⅱ连接；所述待生斜管用于将待生催化剂输出。
具体地，待生剂循环管用于将待生区部分的待生催化剂输送至反应区Ⅱ中。待生剂循环管上设有待生剂循环滑阀。
可选地，所述气固分离区Ⅱ通过焦调控产品气输送管与床层反应器集气室连通；所述反应区Ⅱ通过焦调控催化剂输送管与反应区Ⅰ连通。
具体地，焦调控催化剂输送管上还设有焦调控催化剂滑阀。
可选地，所述装置还包括再生器；所述再生器与待生斜管连接，用于将待生催化剂输送至所述再生器中；所述再生器与提升管反应器连接，用于将再生催化剂输送至所述焦调控反应器中；所述再生器的内底部设有再生器分布器。
具体地，再生器分布器用于通入再生气体。
可选地，所述再生器的底部还设有再生器汽提器；所述再生器汽提器的上段设置在所述再生器的内部，且所述再生器汽提器上段的入口位于所述再生器分布器的上方；所述再生器汽提器的下段设置在所述再生器的外部，且所述再生器汽提器下段的出口与所述提升管反应器连接；所述再生器汽提器中还设有再生器取热器。
可选地，所述再生器通过待生剂输送管和甲醇转化反应器汽提器与所述待生斜管连接；所述再生器通过再生器汽提器和再生斜管与所述提升管反应器的进口连接。
具体地，在待生剂输送管和甲醇转化反应器汽提器之间设有待生滑阀；待生滑阀入口经管道连接于甲醇转化反应器汽提器的底部，待生滑阀的出口经管道连接于待生剂输送管的入口。
在再生器汽提器和再生斜管之间设有再生滑阀，再生滑阀的入口经管道连接于再生器汽提器的底部，再生滑阀的出口经管道连接于再生斜管的入口。
可选地，所述再生器中还设有再生器气固分离设备和再生器集气室；所述再生器气固分离设备的催化剂出口位于所述再生器分布器的上方；所述再生器气固分离设备的气体出口与所述再生器集气室连接；所述再生器集气室与位于所述再生器外部的烟气输送管连接。
根据本申请的第三方面，还提供了一种在线改性DMTO催化剂的方法，将提升管反应器原料和 催化剂从提升管反应器通入过渡区，将床层反应器原料通入反应区Ⅰ；所述催化剂与提升管反应器原料和床层反应器原料接触，反应，生成焦调控催化剂和焦调控产品气；所述催化剂为DMTO催化剂；所述焦调控催化剂为改性DMTO催化剂。
可选地，所述催化剂的活性组分为SAPO-34分子筛。
本申请中，进入提升管反应器中的催化剂可以为新的催化剂，或者再生催化剂，优选的为再生催化剂，这样可以在线同时实现焦调控和催化剂再生。
可选地，所述催化剂为再生催化剂；所述再生催化剂中的焦含量≤3wt％。
可选地，所述焦调控催化剂中的焦含量为4-9wt％。
可选地，所述焦调控催化剂中的焦含量分布的四分位差小于1wt％。具体地，本申请中，通过焦调控反应器的设置以及焦调控工艺的选择，实现了焦调控催化剂中的焦含量为4-9wt％，由于催化剂为颗粒状，所以催化剂的焦含量是指每个催化剂颗粒焦含量的均值，但是每个催化剂颗粒中的焦含量实际上是不一样的。本申请中，可以将焦调控催化剂中的焦含量分布的四分位差控制在小于1wt％的范围内，使得催化剂整体焦含量分布窄，从而提高催化剂的活性、以及低碳烯烃选择性。
可选地，所述焦调控催化剂中的焦物种包含多甲基苯和多甲基萘；多甲基苯和多甲基萘的质量在焦总质量中的含量为≥70wt％；分子量＞184的焦物种的质量在焦总质量中的含量为≤25wt％；其中，所述焦总质量是指焦物种的总质量。
本申请中，焦物种的类型，以及焦物种的含量也非常重要，也是本申请调控的目的之一。本申请中，通过焦调控的设置以及焦调控工艺参数的选择，实现了多甲基苯和多甲基萘在焦总质量中的含量≥70wt％的效果，提高了催化剂的活性，以及低碳烯烃选择性。
可选地，提升管反应器的工艺操作条件为：气体表观线速度为3-10m/s，反应温度为400-700℃，反应压力为100-500kPa，床层密度为10-150kg/m 3。
可选地，床层反应器的反应区Ⅰ的工艺操作条件为：气体表观线速度为0.1-1.0m/s，反应温度为300-650℃，反应压力为100-500kPa，床层密度为150-800kg/m 3。
根据本申请的第四方面，还提供了一种含氧化合物制备低碳烯烃的方法，所述方法包括上述任一项所述的在线改性DMTO催化剂的方法。
可选地，所述方法还包括：将焦调控产品气通入甲醇转化反应器的气固分离区；将焦调控催化剂通入甲醇转化反应器的反应区Ⅱ。
可选地，在反应区Ⅱ中，将含有含氧化合物的原料与焦调控催化剂接触，反应，生成含有低碳烯烃和待生催化剂的物流A。
可选地，所述物流A在甲醇转化反应器的气固分离区Ⅱ进行气固分离后，分为气相物流B和固 相物流C；所述气相物流B进入甲醇转化反应器集气室；所述固相物流C进入待生剂区；其中，所述气相物流B含有低碳烯烃，所述固相物流C含待生催化剂。
可选地，将待生剂区流化气体通入待生剂区；所述待生剂区流化气体、焦调控产品气混合携带部分待生催化剂形成物流D；对所述物流D进行气固分离，分离后得到气相物流E和固相物流F；所述气相物流E进入甲醇转化反应器集气室；所述固相物流F进入待生剂区；其中，所述气相物流E是待生剂区流化气体和焦调控产品气的混合气体；
所述固相物流F是待生催化剂。
可选地，所述气相物流B和气相物流E在甲醇转化反应器集气室中混合形成产品气，所述产品气经由产品气输送管进入下游工段。
可选地，在待生剂区的一部分所述待生催化剂经过待生剂循环管返回反应区Ⅱ的底部；另一部分所述待生催化剂经由待生斜管排出。
可选地，将经由所述待生斜管排出的待生催化剂通入再生器中；将再生气体通入所述再生器中，与所述待生催化剂接触、反应，得到含有烟气和再生催化剂的物流G。
可选地，对所述物流G进行气固分离；分离后的烟气进入再生器集气室，再经由烟气输送管进入下游的烟气处理系统；对分离后的再生催化剂进行汽提、取热，之后进入焦调控反应器中。
具体地，分离后的再生催化剂进入提升管反应器进行汽提、取热，由所述提升管反应器通入所述床层反应器中。
可选地，所述含氧化合物包括甲醇和/或二甲醚。
可选地，所述待生催化剂中的焦含量为9-13wt％。
可选地，所述待生剂区流化气体包括氮气和/或水蒸气。
可选地，所述再生气体包括0-100wt％空气、0-50wt％氧气、0-50wt％氮气和0-50wt％水蒸气，所述空气、氧气、氮气和水蒸气不同时为0。
可选地，甲醇转化反应器的反应区Ⅱ的工艺操作条件为：气体表观线速度为0.5-7.0m/s，反应温度为350-550℃，反应压力为100-500kPa，床层密度为100-500kg/m 3。
可选地，甲醇转化反应器的待生剂区的工艺操作条件为：气体表观线速度为0.1-1.0m/s，反应温度为350-550℃，反应压力为100-500kPa，床层密度为200-800kg/m 3。
可选地，再生器的工艺操作条件为：气体表观线速度为0.5-2.0m/s，再生温度为600-750℃，再生压力为100-500kPa，床层密度为150-700kg/m 3。
本申请中，“焦含量”，是指焦物种与焦调控催化剂的质量比；
C 4-C 6中的下标均表示基团所包含的碳原子数，例如C 4-C 6烃类表示碳原子数为4-6的烃类。
本申请表述生产单耗时，将含氧化合物中的二甲醚质量依据C元素质量等同折算为甲醇质量计，生产单耗的单位为吨甲醇/吨低碳烯烃。
本申请所述的方法中，生产单耗为2.50-2.60吨甲醇/吨低碳烯烃。
提升管反应器近似于平推流反应器，因此，采用提升管反应器对催化剂进行焦调控处理，可以获得较窄的焦含量分布。催化剂在提升管反应器中的停留时间较短，一般为1-20s，因此，仅采用提升管反应器处理再生催化剂，难以大幅度地提高催化剂中的焦含量。本申请中的焦调控反应器包含一个提升管反应器和一个床层反应器，一方面利用提升管反应器的优势，获得具有较窄焦含量分布的催化剂，另一方面又利用床层反应器进一步提高催化剂中的焦含量，提高低碳烯烃的选择性。本申请中的床层反应器的主要特征还在于采用多孔板抑制催化剂在床层间的反混，提高催化剂中的焦分布的均匀性。催化剂先由提升管反应器进入床层反应器的上层，逐步向下层流动，然后由下层进入甲醇转化反应器的反应区II。
本申请能产生的有益效果包括：
(1)本申请中的焦调控反应器可以控制催化剂中的焦物种的转化和生成，一方面将再生催化剂中残留的非活性的大分子焦物种转化为小分子焦物种，另一方面，提升管反应器原料和床层反应器原料还可进入催化剂中生成高活性的小分子焦物种，并且小分子焦物种以多甲基苯和多甲基萘为主，可以提高乙烯的选择性。
(2)本申请中的通过焦调控反应在线改性DMTO催化剂的方法，可以获得焦含量高，焦含量分布窄，焦物种的主要成分是多甲基苯和多甲基萘的焦调控催化剂，将低碳烯烃选择性较低的再生催化剂转化为低碳烯烃选择性高的焦调控催化剂。
(3)本申请中的再生催化剂也可以不经过焦调控过程处理，直接用于含氧化合物制备低碳烯烃过程，不经过焦调控处理时，所得的产品气中的低碳烯烃选择性为80-83wt％。本申请中的再生催化剂经过焦调控过程处理后再用于含氧化合物制备低碳烯烃过程，所得的产品气中的低碳烯烃选择性为92-96wt％。
(4)本申请中的甲醇转化反应器采用了包含一个快速流化床区和一个鼓泡流化床区的复合流化床反应器；快速流化床区是反应区，可以获得较高的甲醇通量，提高设备单位体积的甲醇处理量，甲醇质量空速可以达到5-20h -1，鼓泡流化床区是待生剂区，用于取热、降低待生催化剂的温度，并向反应区输送低温的待生催化剂，提高反应区的床层密度、控制反应区的床层温度，当气体表观线速度为0.5-7.0m/s时，相对应的床层密度为500-100kg/m 3。
(5)本申请中的甲醇转化反应器采用了甲醇转化反应器第一气固分离设备直接连接于输送管的结构，实现了物流A中含有低碳烯烃的气体和待生催化剂的快速分离，避免了低碳烯烃在待生催化 剂的作用下进一步反应生成具有更大分子量的烃类副产品。
附图说明
图1为本申请一个实施方案的含氧化合物制低碳烯烃(DMTO)装置的结构示意图。
部件和附图标记列表：
1焦调控反应器；1-1提升管反应器；1-2床层反应器；
1-3床层反应器壳体；1-4床层反应器分布器；1-5多孔板；
1-6床层反应器气固旋风分离器；1-7床层反应器集气室；
1-8焦调控产品气输送管；1-9焦调控催化剂输送管；
1-10焦调控催化剂滑阀；
2甲醇转化反应器；2-1甲醇转化反应器壳体；
2-2甲醇转化反应器分布器；2-3输送管；
2-4甲醇转化反应器第一气固分离设备；
2-5甲醇转化反应器集气室；2-6待生剂区气体分布器；
2-7甲醇转化反应器取热器；
2-8甲醇转化反应器第二气固分离设备；
2-9产品气输送管；2-10待生剂循环管；
2-11待生剂循环滑阀；2-12待生斜管；
2-13甲醇转化反应器汽提器；
2-14待生滑阀；2-15待生剂输送管；
3-再生器；3-1-再生器壳体；3-2-再生器分布器；
3-3-再生器气固分离设备；3-4-再生器集气室；
3-5-烟气输送管；3-6-再生器汽提器；
3-7-再生器取热器；3-8-再生滑阀；3-9-再生斜管。
具体实施方式
下面结合实施例详述本申请，但本申请并不局限于这些实施例。
DMTO催化剂的一个主要特性是甲醇转化过程的低碳烯烃选择性随着催化剂的焦含量升高而升 高。本申请中所述的低碳烯烃是指乙烯和丙烯。
申请人研究发现，影响DMTO催化剂的活性和低碳烯烃选择性的主要因素是催化剂中的焦含量、焦含量分布和焦物种。催化剂的平均焦含量相同时，焦含量分布窄，则低碳烯烃选择性高、活性高。催化剂中的焦物种包含多甲基芳烃和多甲基环烷烃等，其中，多甲基苯和多甲基萘能促进乙烯的生成。因此，控制催化剂中的焦含量、焦含量分布以及焦物种是控制DMTO催化剂活性、提高低碳烯烃选择性的关键。
为了提高DMTO催化剂的性能，本申请提供了一种通过焦调控反应在线改性DMTO催化剂的方法，其步骤包含：
a)将再生催化剂输送至焦调控反应器；
b)将包含氢气、甲烷、乙烷、乙烯、丙烷、丙烯、丁烷、丁烯、戊烷、戊烯、己烷、己烯、甲醇、乙醇和水的焦调控原料输送至焦调控反应器；
c)焦调控原料和再生催化剂在焦调控反应器中接触并发生反应，焦调控原料在再生催化剂上结焦，结焦后的催化剂被称之为焦调控催化剂，焦调控催化剂中的焦含量为4-9wt％，焦含量分布的四分位差小于1wt％，焦物种中包含多甲基苯和多甲基萘，多甲基苯和多甲基萘的质量在焦总质量中的含量为≥70wt％，分子量>184的焦物种的质量在焦总质量中的含量为≤25wt％；
d)将焦调控催化剂输送至甲醇转化反应器。
所述再生催化剂是焦含量≤3wt％的DMTO催化剂，所述DMTO催化剂的活性组分是SAPO-34分子筛。
所述焦调控原料的组成为0-20wt％氢气、0-50wt％甲烷、0-50wt％乙烷、0-20wt％乙烯、0-50wt％丙烷、0-20wt％丙烯、0-90wt％丁烷、0-90wt％丁烯、0-90wt％戊烷、0-90wt％戊烯、0-90wt％己烷、0-90wt％己烯、0-50wt％甲醇、0-50wt％乙醇和0-50wt％水，并且，甲醇、乙醇和水的总含量≥10wt％。
所述焦调控反应的反应温度为300-700℃。
本申请中，床层反应器原料和提升管反应原料均为焦调控原料。
根据本申请的一个方面，提供了一种包含上述通过焦调控反应在线改性DMTO催化剂的方法的含氧化合物制备低碳烯烃的方法及其使用的装置。所述装置包含焦调控反应器(1)、甲醇转化反应器(2)和再生器(3)。
一种在线改性DMTO催化剂的焦调控反应器(1)，所述焦调控反应器(1)包含提升管反应器(1-1)和床层反应器(1-2)，所述床层反应器(1-2)包含：床层反应器壳体(1-3)，床层反应器分布器(1-4)，多孔板(1-5)，床层反应器气固旋风分离器(1-6)，床层反应器集气室(1-7)，焦调控产品气输送管(1-8)，焦调控催化剂输送管(1-9)和焦调控催化剂滑阀(1-10)；
所述床层反应器(1-2)由下至上分为反应区、过渡区和气固分离区。提升管反应器(1-1)穿过 床层反应器壳体(1-3)，部分位于床层反应器(1-2)的下方，部分位于床层反应器(1-2)之中。床层反应器分布器(1-4)位于反应区的底部，反应区内设置n块多孔板(1-5)，0≤n≤9；
所述床层反应器气固旋风分离器(1-6)位于气固分离区，床层反应器气固旋风分离器(1-6)的入口位于气固分离区，床层反应器气固旋风分离器(1-6)的催化剂出口位于反应区，床层反应器气固旋风分离器(1-6)的气体出口连接于床层反应器集气室(1-7)；
所述床层反应器集气室(1-7)位于床层反应器(1-2)的顶部；焦调控产品气输送管(1-8)的入口连接于床层反应器集气室(1-7)，焦调控产品气输送管(1-8)的出口连接于甲醇转化反应器(2)的上部；
所述焦调控催化剂输送管(1-9)中设置焦调控催化剂滑阀(1-10)；焦调控催化剂输送管(1-9)的入口连接于反应区的下部，焦调控催化剂输送管(1-9)的出口连接于甲醇转化反应器(2)的下部。
在一个优选实施方式中，多孔板的开孔率为1-30％。
一种用于甲醇转化制备低碳烯烃的甲醇转化反应器(2)，所述甲醇转化反应器(2)包含：甲醇转化反应器壳体(2-1)，甲醇转化反应器分布器(2-2)，输送管(2-3)，甲醇转化反应器第一气固分离设备(2-4)，甲醇转化反应器集气室(2-5)，待生剂区气体分布器(2-6)，甲醇转化反应器取热器(2-7)，甲醇转化反应器第二气固分离设备(2-8)，产品气输送管(2-9)，待生剂循环管(2-10)，待生剂循环滑阀(2-11)，待生斜管(2-12)，甲醇转化反应器汽提器(2-13)，待生滑阀(2-14)和待生剂输送管(2-15)；
所述甲醇转化反应器(2)的下部是反应区，中部是待生剂区，上部是气固分离区。甲醇转化反应器分布器(2-2)位于甲醇转化反应器(2)的反应区的底部，输送管(2-3)位于甲醇转化反应器(2)中部和上部的中心区域，输送管(2-3)的底端连接于反应区的顶端，输送管(2-3)的上部连接于甲醇转化反应器第一气固分离设备(2-4)的入口，甲醇转化反应器第一气固分离设备(2-4)位于甲醇转化反应器(2)的气固分离区，甲醇转化反应器第一气固分离设备(2-4)的气体出口连接于甲醇转化反应器集气室(2-5)，甲醇转化反应器第一气固分离设备(2-4)的催化剂出口位于待生剂区；
所述待生剂区气体分布器(2-6)位于待生剂区的底部，甲醇转化反应器取热器(2-7)位于待生剂区，甲醇转化反应器第二气固分离设备(2-8)位于甲醇转化反应器(2)的气固分离区，甲醇转化反应器第二气固分离设备(2-8)的入口位于甲醇转化反应器(2)的气固分离区，甲醇转化反应器第二气固分离设备(2-8)的气体出口连接于甲醇转化反应器集气室(2-5)，甲醇转化反应器第二气固分离设备(2-8)的催化剂出口位于待生剂区，甲醇转化反应器集气室(2-5)位于甲醇转化反应器(2)的顶部，产品气输送管(2-9)连接于甲醇转化反应器集气室(2-5)的顶部；
所述待生剂循环管(2-10)的入口连接于待生剂区，待生剂循环管(2-10)的出口连接于甲醇转 化反应器(2)的反应区的底部，待生剂循环管(2-10)中设置待生剂循环滑阀(2-11)，焦调控催化剂输送管(1-9)的出口连接于甲醇转化反应器(2)的反应区的底部，待生斜管(2-12)的入口连接于待生剂区，待生斜管(2-12)的出口连接于甲醇转化反应器汽提器(2-13)的上部，甲醇转化反应器汽提器(2-13)置于甲醇转化反应器壳体(2-1)之外，待生滑阀(2-14)的入口经管道连接于甲醇转化反应器汽提器(2-13)的底部，待生滑阀(2-14)的出口经管道连接于待生剂输送管(2-15)的入口，待生剂输送管(2-15)的出口连接于再生器(3)的中部。
在一个优选实施方式中，甲醇转化反应器第一气固分离设备(2-4)采用一组或多组气固旋风分离器，每组气固旋风分离器包含一个第一级气固旋风分离器和一个第二级气固旋风分离器。
在一个优选实施方式中，甲醇转化反应器第二气固分离设备(2-8)采用一组或多组气固旋风分离器，每组气固旋风分离器包含一个第一级气固旋风分离器和一个第二级气固旋风分离器。
所述甲醇转化反应器(2)属于流化床反应器。
一种用于再生催化剂的再生器(3)，所述再生器(3)包含：再生器壳体(3-1)，再生器分布器(3-2)，再生器气固分离设备(3-3)，再生器集气室(3-4)，烟气输送管(3-5)，再生器汽提器(3-6)，再生器取热器(3-7)，再生滑阀(3-8)和再生斜管(3-9)；
所述再生器分布器(3-2)位于再生器(3)的底部，再生器气固分离设备(3-3)位于再生器(3)的上部，再生器气固分离设备(3-3)的入口位于再生器(3)的上部，再生器气固分离设备(3-3)的气体出口连接于再生器集气室(3-4)，再生器气固分离设备(3-3)的催化剂出口位于再生器(3)的下部，再生器集气室(3-4)位于再生器(3)的顶部，烟气输送管(3-5)连接于再生器集气室(3-4)的顶部；
所述再生器汽提器(3-6)位于再生器壳体(3-1)之外，再生器汽提器(3-6)的入口管穿透再生器壳体(3-1)，开口于再生器分布器(3-2)的上方，再生器取热器(3-7)位于再生器汽提器(3-6)之中，再生滑阀(3-8)的入口经管道连接于再生器汽提器(3-6)的底部，再生滑阀(3-8)的出口经管道连接于再生斜管(3-9)的入口，再生斜管(3-9)的出口连接于提升管反应器(1-1)的底部。
在一个优选实施方式中，再生器气固分离设备(3-3)采用一组或多组气固旋风分离器，每组气固旋风分离器包含一个第一级气固旋风分离器和一个第二级气固旋风分离器。
所述再生器(3)属于流化床反应器。
本申请中，焦调控反应器的过渡区的内径由下至上逐渐增大；甲醇转化反应器中，反应区Ⅱ与输送管连接处的内径由下至上逐渐减小；反应区Ⅱ与待生剂区的连接处的内径由下至上逐渐增大
可选地，焦调控催化剂输送管的进口位于床层反应器分布器的上方；焦调控催化剂输送管的出口位于甲醇转化反应器分布器的上方；待生剂循环管的进口位于待生剂区气体分布器的上方，待生剂循环管的出口位于甲醇转化反应器分布器的上方；待生斜管位于待生剂区气体分布器的上方。
根据本申请的另一个方面，还提供了一种包含通过焦调控反应在线改性DMTO催化剂的方法的甲醇制烯烃方法，包括以下步骤：
a.将提升管反应器原料从提升管反应器(1-1)的底部通入，其携带由再生斜管(3-9)而来的再生催化剂向上进入床层反应器(1-2)的过渡区；将床层反应器原料从床层反应器分布器(1-4)通入床层反应器(1-2)的反应区；在提升管反应器(1-1)和床层反应器(1-2)中，再生催化剂和提升管反应器原料以及床层反应器原料接触，发生化学反应，生成焦调控催化剂和焦调控产品气；焦调控产品气携带部分焦调控催化剂进入床层反应器气固旋风分离器(1-6)，气固分离后，焦调控产品气经由床层反应器集气室(1-7)和焦调控产品气输送管(1-8)进入甲醇转化反应器(2)的气固分离区；焦调控催化剂经由焦调控催化剂输送管(1-9)、焦调控催化剂滑阀(1-10)进入甲醇转化反应器(2)的反应区；
b.将含有含氧化合物的原料从甲醇转化反应器分布器(2-2)通入甲醇转化反应器(2)的反应区，与焦调控催化剂接触，生成含有低碳烯烃和待生催化剂的物流A，物流A经过输送管(2-3)进入甲醇转化反应器第一气固分离设备(2-4)，气固分离后，分为气相物流B和固相物流C，气相物流B是含有低碳烯烃的气体，固相物流C是待生催化剂，气相物流B进入甲醇转化反应器集气室(2-5)，固相物流C进入待生剂区；将待生剂区流化气体从待生剂区气体分布器(2-6)通入待生剂区，和待生催化剂接触，待生剂区流化气体和焦调控产品气混合携带部分待生催化剂形成物流D，物流D进入甲醇转化反应器第二气固分离设备(2-8)，气固分离后，分为气相物流E和固相物流F，气相物流E是待生剂区流化气体和焦调控产品气的混合气体，固相物流F是待生催化剂，气相物流E进入甲醇转化反应器集气室(2-5)，固相物流F进入待生剂区；气相物流B和气相物流E在甲醇转化反应器集气室(2-5)中混合形成产品气，产品气经由产品气输送管(2-9)进入下游工段；待生剂区的一部分待生催化剂经过待生剂循环管(2-10)和待生剂循环滑阀(2-11)返回甲醇转化反应器(2)的反应区的底部，另一部分待生催化剂经由待生斜管(2-12)进入甲醇转化反应器汽提器(2-13)，汽提之后，待生催化剂再经由待生滑阀(2-14)和待生剂输送管(2-15)进入再生器(3)的中部；
c.将再生气体从再生器分布器(3-2)通入再生器(3)的底部，在再生器中，再生气体和待生催化剂接触，发生化学反应，待生催化剂中的部分焦被燃烧消除，生成含有烟气和再生催化剂的物流G，物流G进入再生器气固分离设备(3-3)，气固分离后，分为烟气和再生催化剂，烟气进入再生器集气室(3-4)，再经由烟气输送管(3-5)进入下游的烟气处理系统，再生催化剂返回再生器(3)的底部，再生器(3)中的再生催化剂进入再生器汽提器(3-6)，汽提、取热之后，再经由再生滑阀(3-8)和再生斜管(3-9)进入焦调控反应器(1)。
在一个优选实施方式中，本申请所述的方法使用上述包含焦调控反应器(1)、甲醇转化反应器(2) 和再生器(3)的装置进行。
在一个优选实施方式中，所述的方法中的提升管反应器原料的组成为0-20wt％氢气、0-50wt％甲烷、0-50wt％乙烷、0-20wt％乙烯、0-50wt％丙烷、0-20wt％丙烯、0-90wt％丁烷、0-90wt％丁烯、0-90wt％戊烷、0-90wt％戊烯、0-90wt％己烷、0-90wt％己烯、0-50wt％甲醇、0-50wt％乙醇和0-50wt％水，并且，甲醇、乙醇和水的总含量≥10wt％。
在一个优选实施方式中，所述的方法中的床层反应器原料的组成为0-20wt％氢气、0-50wt％甲烷、0-50wt％乙烷、0-20wt％乙烯、0-50wt％丙烷、0-20wt％丙烯、0-90wt％丁烷、0-90wt％丁烯、0-90wt％戊烷、0-90wt％戊烯、0-90wt％己烷、0-90wt％己烯、0-50wt％甲醇、0-50wt％乙醇和0-50wt％水，并且，甲醇、乙醇和水的总含量≥10wt％。
在一个优选实施方式中，所述的方法中的含氧化合物是甲醇或二甲醚中的一种或甲醇和二甲醚的混合物。
在一个优选实施方式中，所述的方法中的待生剂区流化气体是氮气和水蒸气中的一种或氮气和水蒸气的混合物。
在一个优选实施方式中，所述的方法中的再生气体是0-100wt％空气、0-50wt％氧气、0-50wt％氮气和0-50wt％水蒸气。
在一个优选实施方式中，所述催化剂的活性组分是SAPO-34分子筛。
在一个优选实施方式中，所述再生催化剂中的焦含量≤3wt％。
在一个优选实施方式中，所述焦调控催化剂中的焦含量为4-9wt％，焦含量分布的四分位差小于1wt％，焦物种中包含多甲基苯和多甲基萘，多甲基苯和多甲基萘的质量在焦总质量中的含量为≥70wt％，分子量>184的焦物种的质量在焦总质量中的含量为≤25wt％。
在一个优选实施方式中，所述待生催化剂中的焦含量为9-13wt％，进一步优选地，待生催化剂中的焦含量为10-12wt％。
在一个优选实施方式中，所述提升管反应器(1-1)的工艺操作条件为：气体表观线速度为3-10m/s，反应温度为400-700℃，反应压力为100-500kPa，床层密度为10-150kg/m 3。
在一个优选实施方式中，所述床层反应器(1-2)的反应区的工艺操作条件为：气体表观线速度为0.1-1.0m/s，反应温度为300-650℃，反应压力为100-500kPa，床层密度为150-800kg/m 3。
在一个优选实施方式中，所述甲醇转化反应器(2)的反应区的工艺操作条件为：气体表观线速度为0.5-7.0m/s，反应温度为350-550℃，反应压力为100-500kPa，床层密度为100-500kg/m 3。
在一个优选实施方式中，所述甲醇转化反应器(2)的待生剂区的工艺操作条件为：气体表观线速度为0.1-1.0m/s，反应温度为350-550℃，反应压力为100-500kPa，床层密度为200-800kg/m 3。
在一个优选实施方式中，所述再生器(3)的工艺操作条件为：气体表观线速度为0.5-2.0m/s，再 生温度为600-750℃，再生压力为100-500kPa，床层密度为150-700kg/m 3。
本申请所述的方法中，产品气的组成为40-55wt％乙烯，37-53wt％丙烯，≤4wt％C 4-C 6烃类和≤4wt％的其他组分，其他组分是甲烷、乙烷、丙烷、氢气、CO和CO 2等，并且乙烯和丙烯在产品气中的总选择性为92-96wt％。
实施例1
本实施方案采用图1所示的装置，床层反应器中不含多孔板。
本实施方案中，提升管反应器原料是6wt％丁烷、81wt％丁烯、2wt％甲醇和11wt％水的混合物；床层反应器原料是6wt％丁烷、81wt％丁烯、2wt％甲醇和11wt％水的混合物；含氧化合物是甲醇；待生剂区流化气体是氮气；再生气体是空气；催化剂中的活性组分是SAPO-34分子筛；再生催化剂中的焦含量约为3wt％；焦调控催化剂中的焦含量约为6wt％，其中，多甲基苯和多甲基萘的质量在焦总质量中的含量约为82wt％，分子量>184的焦物种的质量在焦总质量中的含量约为5wt％；焦调控催化剂中的焦含量分布的四分位差约为0.9wt％；待生催化剂中的焦含量约为12wt％；提升管反应器(1-1)的工艺操作条件为：气体表观线速度为10.0m/s，反应温度为700℃，反应压力为100kPa，床层密度为10kg/m 3；床层反应器(1-2)的反应区的工艺操作条件为：气体表观线速度约为0.3m/s，反应温度约为650℃，反应压力约为100kPa，床层密度约为600kg/m 3；甲醇转化反应器(2)的反应区的工艺操作条件为：气体表观线速度约为7.0m/s，反应温度约为550℃，反应压力约为100kPa，床层密度约为100kg/m 3；甲醇转化反应器(2)的待生剂区的工艺操作条件为：气体表观线速度约为1.0m/s，反应温度约为550℃，反应压力约为100kPa，床层密度约为200kg/m 3；再生器(3)的工艺操作条件为：气体表观线速度约为0.5m/s，再生温度约为750℃，再生压力约为100kPa，床层密度约为700kg/m 3。
本实施方案中，甲醇转化反应器的含氧化合物的质量空速约为20h -1；产品气的组成为55wt％乙烯，37wt％丙烯，4wt％C 4-C 6烃类和4wt％的其他组分，其他组分是甲烷、乙烷、丙烷、氢气、CO和CO 2等；生产单耗为2.60吨甲醇/吨低碳烯烃。
实施例2
本实施方案采用图1所示的装置，床层反应器中包含4块多孔板，多孔板的开孔率为30％。
本实施方案中，提升管反应器原料是22wt％甲烷、24wt％乙烷、3wt％乙烯、28wt％丙烷、4wt％丙烯、7wt％氢气和12wt％水的混合物；床层反应器原料是22wt％甲烷、24wt％乙烷、3wt％乙烯、28wt％丙烷、4wt％丙烯、7wt％氢气和12wt％水的混合物；含氧化合物是82wt％甲醇和18wt％二甲醚；待生剂区流化气体是水蒸气；再生气体是50wt％空气和50wt％水蒸气；催化剂中的活性组分是SAPO-34分子筛；再生催化剂中的焦含量约为1wt％；焦调控催化剂中的焦含量约为4wt％，其中，多甲基苯 和多甲基萘的质量在焦总质量中的含量约为76wt％，分子量>184的焦物种的质量在焦总质量中的含量约为15wt％；焦调控催化剂中的焦含量分布的四分位差约为0.5wt％；待生催化剂中的焦含量约为9wt％；提升管反应器(1-1)的工艺操作条件为：气体表观线速度为3.0m/s，反应温度为400℃，反应压力为500kPa，床层密度为150kg/m 3；床层反应器(1-2)的反应区的工艺操作条件为：气体表观线速度约为0.1m/s，反应温度约为300℃，反应压力约为500kPa，床层密度约为800kg/m 3；甲醇转化反应器(2)的反应区的工艺操作条件为：气体表观线速度约为0.5m/s，反应温度约为350℃，反应压力约为500kPa，床层密度约为500kg/m 3；甲醇转化反应器(2)的待生剂区的工艺操作条件为：气体表观线速度约为0.1m/s，反应温度约为350℃，反应压力约为500kPa，床层密度约为800kg/m 3；再生器(3)的工艺操作条件为：气体表观线速度约为2.0m/s，再生温度约为600℃，再生压力约为500kPa，床层密度约为150kg/m 3。
本实施方案中，甲醇转化反应器的含氧化合物的质量空速约为5h -1；产品气的组成为40wt％乙烯，53wt％丙烯，4wt％C 4-C 6烃类和3wt％的其他组分，其他组分是甲烷、乙烷、丙烷、氢气、CO和CO 2等；生产单耗为2.58吨甲醇/吨低碳烯烃。
实施例3
本实施方案采用图1所示的装置，床层反应器中包含4块多孔板，多孔板的开孔率为1％。
本实施方案中，提升管反应器原料是1wt％丙烷、1wt％丙烯、3wt％丁烷、51wt％丁烯、3wt％戊烷、22wt％戊烯、1wt％己烷、7wt％己烯、2wt％甲醇和9wt％水的混合物；床层反应器原料是1wt％丙烷、1wt％丙烯、3wt％丁烷、51wt％丁烯、3wt％戊烷、22wt％戊烯、1wt％己烷、7wt％己烯、2wt％甲醇和9wt％水的混合物；含氧化合物是二甲醚；待生剂区流化气体是5wt％氮气和95wt％水蒸气；再生气体是50wt％空气和50wt％氧气；催化剂中的活性组分是SAPO-34分子筛；再生催化剂中的焦含量约为2wt％；焦调控催化剂中的焦含量约为5wt％，其中，多甲基苯和多甲基萘的质量在焦总质量中的含量约为79wt％，分子量>184的焦物种的质量在焦总质量中的含量约为10wt％；焦调控催化剂中的焦含量分布的四分位差约为0.5wt％；待生催化剂中的焦含量约为11wt％；提升管反应器(1-1)的工艺操作条件为：气体表观线速度为6.0m/s，反应温度为600℃，反应压力为300kPa，床层密度为80kg/m 3；床层反应器(1-2)的反应区的工艺操作条件为：气体表观线速度约为0.4m/s，反应温度约为550℃，反应压力约为300kPa，床层密度约为500kg/m 3；甲醇转化反应器(2)的反应区的工艺操作条件为：气体表观线速度约为3.0m/s，反应温度约为450℃，反应压力约为300kPa，床层密度约为230kg/m 3；甲醇转化反应器(2)的待生剂区的工艺操作条件为：气体表观线速度约为0.2m/s，反应温度约为450℃，反应压力约为300kPa，床层密度约为600kg/m 3；再生器(3)的工艺操作条件为：气体表观线速度约为1.0m/s，再生温度约为680℃，再生压力约为300kPa，床层密度约为360kg/m 3。
本实施方案中，甲醇转化反应器的含氧化合物的质量空速约为15h -1；产品气的组成为50wt％乙烯，45wt％丙烯，3wt％C 4-C 6烃类和2wt％的其他组分，其他组分是甲烷、乙烷、丙烷、氢气、CO和CO 2等；生产单耗为2.53吨甲醇/吨低碳烯烃。
实施例4
本实施方案采用图1所示的装置，床层反应器中包含9块多孔板，多孔板的开孔率为5％。
本实施方案中，提升管反应器原料是5wt％丁烷、72wt％丁烯、8wt％甲醇和15wt％水的混合物；床层反应器原料是5wt％丁烷、72wt％丁烯、8wt％甲醇和15wt％水的混合物；含氧化合物是甲醇；待生剂区流化气体是73wt％氮气和27wt％水蒸气；再生气体是50wt％空气和50wt％氮气；催化剂中的活性组分是SAPO-34分子筛；再生催化剂中的焦含量约为2wt％；焦调控催化剂中的焦含量约为9wt％，其中，多甲基苯和多甲基萘的质量在焦总质量中的含量约为71wt％，分子量>184的焦物种的质量在焦总质量中的含量约为25wt％；焦调控催化剂中的焦含量分布的四分位差约为0.2wt％；待生催化剂中的焦含量约为13wt％；提升管反应器(1-1)的工艺操作条件为：气体表观线速度为4.0m/s，反应温度为550℃，反应压力为200kPa，床层密度为120kg/m 3；床层反应器(1-2)的反应区的工艺操作条件为：气体表观线速度约为1.0m/s，反应温度约为500℃，反应压力约为200kPa，床层密度约为150kg/m 3；甲醇转化反应器(2)的反应区的工艺操作条件为：气体表观线速度约为4.0m/s，反应温度约为500℃，反应压力约为200kPa，床层密度约为160kg/m 3；甲醇转化反应器(2)的待生剂区的工艺操作条件为：气体表观线速度约为0.5m/s，反应温度约为500℃，反应压力约为200kPa，床层密度约为300kg/m 3；再生器(3)的工艺操作条件为：气体表观线速度约为1.5m/s，再生温度约为700℃，再生压力约为200kPa，床层密度约为280kg/m 3。
本实施方案中，甲醇转化反应器的含氧化合物的质量空速约为11h -1；产品气的组成为52wt％乙烯，44wt％丙烯，2wt％C 4-C 6烃类和2wt％的其他组分，其他组分是甲烷、乙烷、丙烷、氢气、CO和CO 2等；生产单耗为2.50吨甲醇/吨低碳烯烃。
实施例5
本实施例是对比案例，和实施例4的差异在于不采用焦调控反应在线改性DMTO催化剂，提升管反应器和床层反应器通入的原料是氮气，氮气是惰性气体，不会在提升管反应器和床层反应器中改变再生催化剂的性质，即，相当于进入甲醇转化反应器的反应区II的催化剂是再生催化剂。
本实施方案中，产品气的组成为46wt％乙烯，37wt％丙烯，12wt％C4-C6烃类和5wt％的其他组分，其他组分是甲烷、乙烷、丙烷、氢气、CO和CO 2等；生产单耗为2.90吨甲醇/吨低碳烯烃。
本对比案例说明了通过焦调控反应在线改性DMTO催化剂可以大幅度的提升催化剂的性能，降 低生产单耗。
以上所述，仅是本申请的几个实施例，并非对本申请做任何形式的限制，虽然本申请以较佳实施例揭示如上，然而并非用以限制本申请，任何熟悉本专业的技术人员，在不脱离本申请技术方案的范围内，利用上述揭示的技术内容做出些许的变动或修饰均等同于等效实施案例，均属于技术方案范围内。

Claims (41)

  1. 一种焦调控反应器，其特征在于，所述焦调控反应器包括提升管反应器和床层反应器；
    所述床层反应器包括床层反应器壳体，所述床层反应器壳体由下至上围合成反应区Ⅰ、过渡区和气固分离区Ⅰ；
    所述反应区Ⅰ的内下部设有床层反应器分布器；
    所述反应区Ⅰ的外部还设有焦调控催化剂输送管；
    所述提升管反应器的上段穿透所述床层反应器的底部沿轴向插设在所述床层反应器中；
    所述提升管反应器的出口端位于所述过渡区。
  2. 根据权利要求1所述的焦调控反应器，其特征在于，所述反应区Ⅰ内设有至少一个多孔板，多个所述多孔板沿轴向依次排列在所述提升管反应器的外周；
    所述提升管反应器的出口端位于所述多孔板的上方；
    所述床层反应器分布器位于所述多孔板的下方。
  3. 根据权利要求2所述的焦调控反应器，其特征在于，所述多孔板的开孔率为1-30％。
  4. 根据权利要求1所述的焦调控反应器，其特征在于，所述气固分离区Ⅰ设有床层反应器气固分离器和床层反应器集气室；
    所述床层反应器气固分离器的气体出口与所述床层反应器集气室连通；
    所述床层反应器气固分离器的催化剂出口位于多孔板的上方；
    所述床层反应器集气室与位于所述床层反应器外部的焦调控产品气输送管连通。
  5. 根据权利要求1所述的焦调控反应器，其特征在于，所述床层反应器分布器用于通入床层反应器原料；
    所述床层反应器原料包括0-20wt％氢气、0-50wt％甲烷、0-50wt％乙烷、0-20wt％乙烯、0-50wt％丙烷、0-20wt％丙烯、0-90wt％丁烷、0-90wt％丁烯、0-90wt％戊烷、0-90wt％戊烯、0-90wt％己烷、0-90wt％己烯、0-50wt％甲醇、0-50wt％乙醇和0-50wt％水，并且，甲醇、乙醇和水的总含量≥10wt％。
  6. 根据权利要求1所述的焦调控反应器，其特征在于，所述提升管反应器用于通入催化剂和提升管反应原料；
    所述提升管反应原料包括0-20wt％氢气、0-50wt％甲烷、0-50wt％乙烷、0-20wt％乙烯、0-50wt％丙烷、0-20wt％丙烯、0-90wt％丁烷、0-90wt％丁烯、0-90wt％戊烷、0-90wt％戊烯、0-90wt％己烷、0-90wt％己烯、0-50wt％甲醇、0-50wt％乙醇和0-50wt％水，并且，甲醇、乙醇和水的总含量≥10wt％。
  7. 一种由含氧化合物制备低碳烯烃的装置，其特征在于，所述装置包括甲醇转化反应器和权利要求1至6任一项所述的焦调控反应器。
  8. 根据权利要求7所述的装置，其特征在于，所述甲醇转化反应器包括甲醇转化反应器壳体和输送管；
    所述甲醇转化反应器壳体包括下壳体和上壳体；
    所述下壳体围合成反应区Ⅱ，所述反应区Ⅱ的内下部设有甲醇转化反应器分布器；
    所述输送管位于所述反应区Ⅱ的上方，且所述输送管的一端闭合，另一端与所述反应区Ⅱ连通；
    所述上壳体设置在所述输送管的外周；
    所述上壳体与所述输送管的管壁围合形成空腔；
    所述空腔自下至上分为待生剂区和气固分离区Ⅱ；
    所述待生剂区设有待生区气体分布器。
  9. 根据权利要求8所述的装置，其特征在于，所述气固分离区Ⅱ设有甲醇转化反应器第一气固分离设备；
    所述输送管的上部与所述甲醇转化反应器第一气固分离设备的入口连接；
    所述甲醇转化反应器第一气固分离设备的催化剂出口位于所述待生剂区；
    所述甲醇转化反应器第一气固分离设备的气体出口与甲醇转化反应器集气室连通；
    所述甲醇转化反应器集气室与产品气输送管连接。
  10. 根据权利要求8所述的装置，其特征在于，所述气固分离区Ⅱ还设有甲醇转化反应器第二气固分离设备；
    所述甲醇转化反应器第二气固分离设备的进气口位于气固分离区Ⅱ；
    所述甲醇转化反应器第二气固分离设备的催化剂出口位于所述待生剂区；
    所述甲醇转化反应器第二气固分离设备的气体出口与甲醇转化反应器集气室连通。
  11. 根据权利要求8所述的装置，其特征在于，所述待生区气体分布器位于甲醇转化反应器第一气固分离设备和甲醇转化反应器第二气固分离设备的下方；
    所述待生剂区还设有甲醇转化反应器取热器。
  12. 根据权利要求8所述的装置，其特征在于，所述待生剂区的外部还设有待生剂循环管和待生斜管；
    所述待生剂循环管用于将所述待生剂区和所述反应区Ⅱ连接；
    所述待生斜管用于将待生催化剂输出。
  13. 根据权利要求8所述的装置，其特征在于，所述气固分离区Ⅱ通过焦调控产品气输送管与床层反应器集气室连通；
    所述反应区Ⅱ通过焦调控催化剂输送管与反应区Ⅰ连通。
  14. 根据权利要求7所述的装置，其特征在于，所述装置还包括再生器；
    所述再生器与待生斜管连接，用于将待生催化剂输送至所述再生器中；
    所述再生器与提升管反应器连接，用于将再生催化剂输送至所述焦调控反应器中；
    所述再生器的内底部设有再生器分布器。
  15. 根据权利要求14所述的装置，其特征在于，所述再生器的底部还设有再生器汽提器；
    所述再生器汽提器的上段设置在所述再生器的内部，且所述再生器汽提器上段的入口位于所述再生器分布器的上方；
    所述再生器汽提器的下段设置在所述再生器的外部，且所述再生器汽提器下段的出口与所述提升管反应器连接；
    所述再生器汽提器中还设有再生器取热器。
  16. 根据权利要求15所述的装置，其特征在于，所述再生器通过待生剂输送管和甲醇转化反应器汽提器与所述待生斜管连接；
    所述再生器通过再生器汽提器和再生斜管与所述提升管反应器的进口连接。
  17. 根据权利要求14所述的装置，其特征在于，所述再生器中还设有再生器气固分离设备和再生器集气室；
    所述再生器气固分离设备的催化剂出口位于所述再生器分布器的上方；
    所述再生器气固分离设备的气体出口与所述再生器集气室连接；
    所述再生器集气室与位于所述再生器外部的烟气输送管连接。
  18. 一种在线改性DMTO催化剂的方法，其特征在于，将提升管反应器原料和催化剂从提升管反应器通入过渡区，将床层反应器原料通入反应区Ⅰ；
    所述催化剂与提升管反应器原料和床层反应器原料接触，反应，生成焦调控催化剂和焦调控产品气；
    所述催化剂为DMTO催化剂；
    所述焦调控催化剂为改性DMTO催化剂。
  19. 根据权利要求18所述的在线改性DMTO催化剂的方法，其特征在于，所述催化剂的活性组分为SAPO-34分子筛。
  20. 根据权利要求18所述的在线改性DMTO催化剂的方法，其特征在于，所述催化剂为再生催化剂；
    所述再生催化剂中的焦含量≤3wt％。
  21. 根据权利要求18所述的在线改性DMTO催化剂的方法，其特征在于，所述焦调控催化剂中的焦含量为4-9wt％；
    所述焦调控催化剂中的焦含量分布的四分位差小于1wt％。
  22. 根据权利要求18所述的在线改性DMTO催化剂的方法，其特征在于，所述焦调控催化剂中的焦物种包含多甲基苯和多甲基萘；
    多甲基苯和多甲基萘的质量在焦总质量中的含量为≥70wt％；
    分子量＞184的焦物种的质量在焦总质量中的含量为≤25wt％；
    其中，所述焦总质量是指焦物种的总质量。
  23. 根据权利要求18所述的在线改性DMTO催化剂的方法，其特征在于，提升管反应器的工艺操作条件为：气体表观线速度为3-10m/s，反应温度为400-700℃，反应压力为100-500kPa，床层密度为10-150kg/m 3。
  24. 根据权利要求18所述的在线改性DMTO催化剂的方法，其特征在于，床层反应器的反应区Ⅰ的工艺操作条件为：气体表观线速度为0.1-1.0m/s，反应温度为300-650℃，反应压力为100-500kPa，床层密度为150-800kg/m 3。
  25. 一种含氧化合物制备低碳烯烃的方法，其特征在于，所述方法包括权利要求18至24任一项所述的在线改性DMTO催化剂的方法。
  26. 根据权利要求25所述的方法，其特征在于，所述方法还包括：将焦调控产品气通入甲醇转化反应器的气固分离区。
  27. 根据权利要求25所述的方法，其特征在于，所述方法还包括：将焦调控催化剂通入甲醇转化反应器的反应区Ⅱ。
  28. 根据权利要求27所述的方法，其特征在于，在反应区Ⅱ中，将含有含氧化合物的原料与焦调控催化剂接触，反应，生成含有低碳烯烃和待生催化剂的物流A。
  29. 根据权利要求28所述的方法，其特征在于，所述物流A在甲醇转化反应器的气固分离区Ⅱ进行气固分离后，分为气相物流B和固相物流C；
    所述气相物流B进入甲醇转化反应器集气室；
    所述固相物流C进入待生剂区；
    其中，所述气相物流B含有低碳烯烃，所述固相物流C含待生催化剂。
  30. 根据权利要求29所述的方法，其特征在于，将待生剂区流化气体通入待生剂区；
    所述待生剂区流化气体、焦调控产品气混合携带部分待生催化剂形成物流D；
    对所述物流D进行气固分离，分离后得到气相物流E和固相物流F；
    所述气相物流E进入甲醇转化反应器集气室；
    所述固相物流F进入待生剂区；
    其中，所述气相物流E是待生剂区流化气体和焦调控产品气的混合气体；
    所述固相物流F是待生催化剂。
  31. 根据权利要求30所述的方法，其特征在于，所述气相物流B和气相物流E在甲醇转化反应器集气室中混合形成产品气，所述产品气经由产品气输送管进入下游工段。
  32. 根据权利要求30所述的方法，其特征在于，在待生剂区的一部分所述待生催化剂经过待生剂循环管返回反应区Ⅱ的底部；
    另一部分所述待生催化剂经由待生斜管排出。
  33. 根据权利要求32所述的方法，其特征在于，将经由所述待生斜管排出的待生催化剂通入再生器中；
    将再生气体通入所述再生器中，与所述待生催化剂接触、反应，得到含有烟气和再生催化剂的物流G。
  34. 根据权利要求33所述的方法，其特征在于，对所述物流G进行气固分离；
    分离后的烟气进入再生器集气室，再经由烟气输送管进入下游的烟气处理系统；
    对分离后的再生催化剂进行汽提、取热，之后进入焦调控反应器中。
  35. 根据权利要求28所述的方法，其特征在于，所述含氧化合物包括甲醇和/或二甲醚。
  36. 根据权利要求28所述的方法，其特征在于，所述待生催化剂中的焦含量为9-13wt％。
  37. 根据权利要求30所述的方法，其特征在于，所述待生剂区流化气体包括氮气和/或水蒸气。
  38. 根据权利要求33所述的方法，其特征在于，所述再生气体包括0-100wt％空气、0-50wt％氧气、0-50wt％氮气和0-50wt％水蒸气；
    所述空气、氧气、氮气和水蒸气不同时为0。
  39. 根据权利要求25所述的方法，其特征在于，甲醇转化反应器的反应区Ⅱ的工艺操作条件为：气体表观线速度为0.5-7.0m/s，反应温度为350-550℃，反应压力为100-500kPa，床层密度为100-500kg/m 3。
  40. 根据权利要求25所述的方法，其特征在于，甲醇转化反应器的待生剂区的工艺操作条件为：气体表观线速度为0.1-1.0m/s，反应温度为350-550℃，反应压力为100-500kPa，床层密度为200-800kg/m 3。
  41. 根据权利要求25所述的方法，其特征在于，再生器的工艺操作条件为：气体表观线速度为0.5-2.0m/s，再生温度为600-750℃，再生压力为100-500kPa，床层密度为150-700kg/m 3。
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