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WO2018196364A1 - 甲醇和/或二甲醚与苯制对二甲苯联产低碳烯烃的装置及方法 - Google Patents

甲醇和/或二甲醚与苯制对二甲苯联产低碳烯烃的装置及方法 Download PDF

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Publication number
WO2018196364A1
WO2018196364A1 PCT/CN2017/112814 CN2017112814W WO2018196364A1 WO 2018196364 A1 WO2018196364 A1 WO 2018196364A1 CN 2017112814 W CN2017112814 W CN 2017112814W WO 2018196364 A1 WO2018196364 A1 WO 2018196364A1
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WIPO (PCT)
Prior art keywords
reactor
fluidized bed
regenerator
gas
catalyst
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PCT/CN2017/112814
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English (en)
French (fr)
Inventor
叶茂
张涛
刘中民
张今令
唐海龙
贾金明
何长青
王贤高
张骋
李华
赵银峰
李承功
Original Assignee
中国科学院大连化学物理研究所
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Application filed by 中国科学院大连化学物理研究所 filed Critical 中国科学院大连化学物理研究所
Priority to KR1020197034521A priority Critical patent/KR102309237B1/ko
Priority to EP17907026.3A priority patent/EP3616784A4/en
Priority to RU2019133631A priority patent/RU2742576C1/ru
Priority to SG11201910004V priority patent/SG11201910004VA/en
Priority to JP2019555948A priority patent/JP7046977B2/ja
Priority to US16/608,438 priority patent/US11084765B2/en
Publication of WO2018196364A1 publication Critical patent/WO2018196364A1/zh

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Definitions

  • the present invention relates to a device for producing low-carbon olefins produced by co-p-xylene (PX) and a production method thereof, and is particularly suitable for fluidization of co-production of low-carbon olefins by alkylation of methanol and/or dimethyl ether with benzene to para-xylene.
  • Bed device and production method belong to the field of chemistry and chemical industry.
  • Paraxylene is one of the basic organic raw materials in the petrochemical industry. It has a wide range of applications in chemical fiber, synthetic resins, pesticides, pharmaceuticals, and polymer materials.
  • p-xylene production mainly uses toluene, C 9 aromatic hydrocarbons and mixed xylene as raw materials, and is obtained by disproportionation, isomerization, adsorption separation or cryogenic separation. Since the p-xylene content in the product is controlled by thermodynamics, p-xylene only accounts for ⁇ 24% of the C 8 mixed aromatics, and the material circulation processing amount is large during the process, and the equipment is large and the operation cost is high.
  • Low-carbon olefins namely ethylene, propylene and butene
  • Ethylene and propylene are mainly produced from naphtha, depending on the petroleum route.
  • the non-oil route to produce ethylene and propylene has received more and more attention, especially the methanol conversion to low-carbon olefin (MTO) process route, which is to achieve oil substitution strategy, reduce and alleviate China's demand for oil and dependence. An important way.
  • MTO low-carbon olefin
  • Alkylation of toluene with methanol to produce para-xylene and alkylation of benzyl alcohol to toluene and xylene are new ways to increase aromatics production.
  • the selective alkylation of toluene methanol can produce highly selective para-xylene products, but the toluene required for the process is also an intermediate raw material for the production of para-xylene by the aromatics unit, which is in short supply in the market.
  • benzene is a by-product of the aromatics unit, it is estimated that an annual output of 800,000 tons of PX aromatics unit can produce about 300,000 tons of benzene. Therefore, the alkylation of benzene with methanol to produce toluene and xylene will increase the production of aromatics. An effective way.
  • a conventional toluene alkylation process involves mixing methanol and/or dimethyl ether with benzene upstream of the reactor and then feeding the mixture together into the reactor.
  • the reactor type consists of a fixed bed and a fluidized bed.
  • the phased injection of reactants has been employed in various fixed bed and fluidized bed processes.
  • Methanol is the raw material for the alkylation of benzene and / or methyl methoxide and / or dimethyl ether with benzene, and is also the raw material for MTO reaction, but the MTO reaction rate is much higher than benzene and / or methyl methoxide and / or dimethyl ether and benzene The rate of alkylation reaction.
  • One of the characteristics of the MTO reaction is that the reaction rate is much higher than the alkylation reaction of methanol and/or dimethyl ether with benzene.
  • Another important feature is that after the catalyst is carbonized, the methanol conversion rate decreases and the selectivity of the low carbon olefin increases. Therefore, controlling the carbonation of the catalyst is an effective way to improve the selectivity of low-carbon olefins in the MTO reaction.
  • the technical field needs to coordinate and optimize the competition between the alkylation reaction and the MTO reaction from the two aspects of catalyst design and reactor design, so as to synergistically improve the conversion of benzene and the yield of p-xylene. And low carbon olefin yield.
  • a fast fluidized bed reactor for co-production of a low carbon olefin with methanol and/or dimethyl ether and benzene para-xylene
  • the fast fluidized bed reactor solved or improved
  • the reaction, coordination, optimization of the competition between the alkylation reaction and the MTO reaction so as to synergistically improve the conversion of benzene, the yield of p-xylene and the selectivity of low-carbon olefins.
  • a fast fluidized bed reactor for co-production of low carbon olefins with methanol and/or dimethyl ether and benzene para-xylene, comprising a first reactor feed distributor and a plurality of second reactor feed distributors
  • the first reactor feed distributor and the plurality of second reactor feed distributors are arranged in sequence along the gas flow direction in the fast fluidized bed reactor.
  • the second reactor feed distributor is from 2 to 10.
  • the fast fluidized bed reactor comprises a first reactor gas-solids separator and a second reactor gas-solids separator, the first reactor gas-solids separator being placed in a dilute phase zone or a reactor shell Externally, the second reactor gas-solid separator is placed in a dilute phase zone or outside the reactor housing;
  • the first reactor gas-solids separator is provided with a regenerated catalyst inlet, the first reactor gas-solid separator a catalyst outlet is placed at the bottom of the reaction zone, and a gas outlet of the first reactor gas-solid separator is placed in the dilute phase zone;
  • An inlet of the second reactor gas-solids separator is placed in the dilute phase zone, a catalyst outlet of the second reactor gas-solids separator is placed in the reaction zone, and the second reactor gas-solid separator a gas outlet connected to the product gas outlet of the fast fluidized bed reactor;
  • the reaction zone is located in a lower portion of the fast fluidized bed reactor, and the dilute phase zone is located in an upper portion of the fast fluidized bed reactor.
  • the first reactor gas-solid separator and the second reactor gas-solid separator are cyclones.
  • the fast fluidized bed reactor comprises a reactor heat takeer placed inside or outside the fast fluidized bed reactor housing.
  • the reactor heat extractor is disposed between the plurality of reactor feed distributors.
  • the fast fluidized bed reactor comprises a reactor stripper that passes from the outside to the inside of the reactor housing at the bottom of the fast fluidized bed reactor and is open to the fast fluidized bed In the reaction zone of the reactor, the bottom of the reactor stripper is provided with a reactor stripping gas inlet and a catalyst outlet to be produced.
  • the reactor stripper has a higher level of opening in the interior of the reactor housing than the first reactor feed distributor.
  • the lower olefin includes at least one of ethylene, propylene, and butene.
  • the methanol in the feed may be replaced in whole or in part by dimethyl ether. If the feed contains dimethyl ether, the dimethyl ether may be converted into methanol having the same number of carbon atoms in the calculation.
  • methanol and/or dimethyl ether means including three cases: only methanol; or only dimethyl ether; or both methanol and dimethyl ether.
  • methanol and / or dimethyl ether and benzene including three cases: methanol and benzene; or dimethyl ether and benzene; or methanol, dimethyl ether and benzene.
  • an apparatus for co-producing a lower olefin of methanol and/or dimethyl ether with benzene para-xylene solving or improving methanol and/or dimethyl ether
  • the competition between MTO reaction and alkylation reaction in the production of p-xylene to produce low-carbon olefins, the synergistic effect of MTO reaction and alkylation reaction, coordination and optimization of alkylation by controlling mass transfer and reaction The competition between the reaction and the MTO reaction synergizes to increase benzene conversion, p-xylene yield and low olefin selectivity.
  • the apparatus comprises at least one of the fast fluidized bed reactors of the above aspect and a fluidized bed regenerator for regenerating the catalyst.
  • the fluidized bed regenerator is a turbulent fluidized bed regenerator comprising a regenerator housing, a regenerator gas solids separator, a regenerator heat extractor, and a regenerator stripper
  • the lower part of the fluidized bed regenerator is the regeneration zone
  • the upper part of the fluidized bed regenerator is the regenerator dilute phase zone
  • the regenerator feed distributor is placed at the bottom of the regeneration zone
  • the regenerator heat extractor is placed in the regeneration zone for regeneration.
  • the gas-solid separator is placed outside the dilute phase zone or the regenerator housing;
  • the inlet of the regenerator gas-solids separator is placed in a regenerator dilute phase zone, the catalyst outlet of the regenerator gas-solids separator is placed in a regeneration zone, and the regenerator stripper is open at the bottom of the regenerator housing.
  • the fluidized bed regenerator comprises a regenerator housing, a regenerator feed distributor, a regenerator gas-solid separator, a regenerator heat extractor, a flue gas outlet, and a regenerator stripper;
  • the lower part of the fluidized bed regenerator is a regeneration zone, and the upper part of the fluidized bed regenerator is a dilute phase zone;
  • the regenerator feed distributor is placed at the bottom of the regeneration zone, the regenerator heat extractor is placed in the regeneration zone, and the regenerator gas-solids separator is placed outside the dilute phase zone or the regenerator housing, and the inlet of the regenerator gas-solids separator is placed In the dilute phase region, the catalyst outlet of the regenerator gas-solid separator is placed in the regeneration zone, the gas outlet of the regenerator gas-solid separator is connected to the flue gas outlet, and the regenerator stripper is open at the bottom of the regenerator housing;
  • the outlet of the catalyst to be produced in the reactor stripper is connected to the inlet of the inclined tube to be produced, and the slip tube to be produced is provided with a slide valve to be produced, and the outlet of the inclined tube is connected to the inlet of the riser tube to be raised.
  • the bottom of the tube is provided with a rising gas inlet, and the outlet of the rising riser is connected to the dilute phase region of the fluidized bed regenerator;
  • the bottom of the regenerator stripper is provided with a regenerator stripping gas inlet, the bottom of the regenerator stripper is connected to the inlet of the regenerative inclined tube, and the regenerative inclined tube is provided with a regenerative sliding valve, and the outlet of the regenerative inclined tube is connected to
  • the inlet of the regeneration riser is provided with a regeneration lift gas inlet at the bottom of the regeneration riser, and the outlet of the regeneration riser is connected to the regeneration catalyst inlet of the first reactor gas-solid separator, the first reactor gas-solid separator is placed The dilute phase zone of the fluidized bed reactor or the exterior of the reactor housing.
  • a process for the co-production of a lower olefin with methanol and/or dimethyl ether and benzene para-xylene which method solves or improves methanol and/or dimethyl ether
  • the competition between the reaction and the MTO reaction synergizes to increase benzene conversion, p-xylene yield and low olefin selectivity.
  • the method for co-producing a lower olefin with methanol and/or dimethyl ether and benzene para-xylene employs at least one of the fast fluidized bed reactors of the above aspect.
  • the feedstock A containing methanol and/or dimethyl ether and benzene is fed from the first reactor feed distributor to the reaction zone of the fast fluidized bed reactor to feed the feedstock B containing methanol and/or dimethyl ether.
  • the reaction zone fed to the fast fluidized bed reactor by a plurality of second reactor feed distributors, respectively, is contacted with a catalyst to produce a stream C comprising a para-xylene and a low-carbon olefin product and a spent catalyst.
  • the stream C is separated to obtain para-xylene, low-carbon olefins, C 5+ chain hydrocarbons, aromatic hydrocarbon by-products, and unconverted methanol, dimethyl ether and benzene;
  • the unconverted methanol and dimethyl ether are fed to the reaction zone of the fast fluidized bed reactor from a plurality of second reactor feed distributors, the aromatic hydrocarbon by-products and the unconverted benzene are fed from the first reactor.
  • the reaction zone fed to the fast fluidized bed reactor is contacted with a catalyst.
  • the catalyst to be produced is regenerated by a fluidized bed regenerator and then reacted in a fast fluidized bed reactor. The bottom of the area.
  • the method comprises the following steps:
  • Stream B containing methanol and/or dimethyl ether is fed from 2 to 10 second reactor feed distributors to the reaction zone of the fast fluidized bed reactor to contact the catalyst to form p-xylene and a stream C of the low carbon olefin product and a catalyst to be produced; the 2 to 10 second reactor feed distributors are arranged in sequence above the first reactor feed distributor;
  • step C (3) separating step C from stream (c) to obtain unconverted methanol and dimethyl ether stream C-1, unconverted benzene and aromatics by-product stream C-2; and stream C-1 from 2 to 10, respectively.
  • the second reactor feed distributor is fed into the reaction zone of the fast fluidized bed reactor to be in contact with the catalyst; the stream C-2 is fed from the first reactor feed distributor to the reaction zone and catalyst of the fast fluidized bed reactor contact;
  • the aromatic by-product comprises toluene, o-xylene, m-xylene, ethylbenzene and C 9+ arene;
  • step (2) obtains the catalyst to be regenerated by the fluidized bed regenerator, and the regenerated catalyst passes through the first reactor gas-solid separator, and after gas-solid separation, enters the bottom of the reaction zone in the fast fluidized bed reactor.
  • the first reactor feed distributor is fed to the mixture of the fast fluidized bed reactor, the ratio of the molecular moles of aromatics to the carbon moles of methanol and/or dimethyl ether being greater than 0.5.
  • the first reactor feed distributor is fed to the mixture of the fast fluidized bed reactor in a ratio of the molecular moles of aromatic hydrocarbons to the carbon moles of methanol and/or dimethyl ether of from 0.5 to 5.
  • the molar ratio of all oxygenates in the mixture entering the fast fluidized bed reactor from the plurality of second reactor feed distributors to methanol entering the first reactor feed distributor is greater than one.
  • the number of moles of molecules refers to the number of moles of molecules in the substance
  • the number of moles of carbon refers to the number of moles of carbon atoms in the substance.
  • the molar ratio of all oxygenates entering from the plurality of second reactor feed distributors to methanol entering from the first reactor feed distributor is from 1 to 20.
  • the catalyst to be produced passes through the reactor stripper, the inclined tube to be produced, the standby slide valve and the standby riser into the dilute phase region of the fluidized bed regenerator;
  • the regeneration medium is passed into the regeneration zone of the fluidized bed regenerator, and the charcoal reacts with the catalyst to be generated to generate flue gas containing CO and CO 2 and a regenerated catalyst, and the flue gas is discharged after being removed by the regenerator gas-solid separator;
  • the regenerated catalyst enters the inlet of the first reactor gas-solid separator through the regenerator stripper, the regeneration inclined tube, the regeneration slide valve and the regeneration riser. After the gas-solid separation, the regenerated catalyst enters the bottom of the reaction zone in the fast fluidized bed reactor. ;
  • the reactor stripping gas enters the reactor stripper from the reactor stripping gas inlet and is in countercurrent contact with the catalyst to be produced, and then enters the fast fluidized bed reactor; the rising gas to be produced enters the waiting riser from the inlet of the rising gas to be produced and The catalyst to be contacted is in downstream contact and then enters the dilute phase region of the fluidized bed regenerator;
  • regenerator stripping gas is fed back into the regenerator stripper and the regenerated catalyst by the regenerator stripping gas inlet, After entering the fluidized bed regenerator; the regeneration lift gas enters the regeneration riser and the regenerated catalyst in downstream flow from the regeneration lift gas inlet, and then enters the inlet of the first reactor gas-solid separator, the first reactor gas-solid separator Placed in the dilute phase zone of the fluidized bed reactor or outside the reactor housing.
  • the regenerated catalyst has a carbon content of ⁇ 0.5 wt%.
  • the regeneration medium is at least one of air, oxygen-depleted air or water vapor;
  • the reactor stripping gas, the regenerator stripping gas, the raw lifting gas, and the regeneration lifting gas are water vapor or nitrogen.
  • the reaction conditions of the reaction zone of the fast fluidized bed reactor are: an apparent linear velocity of the gas of 1.0 m/s to 8.0 m/s, a reaction temperature of 350 to 600 ° C, and a reaction pressure of 0.1 Mpa to 1.0. MPa, the bed density is from 50 kg/m 3 to 500 kg/m 3 .
  • the reaction condition of the fluidized bed regenerator regeneration zone is: an apparent apparent linear velocity of gas of 0.1 m/s to 2 m/s, a regeneration temperature of 500 ° C to 750 ° C, and a regeneration pressure of 0.1 Mpa to 1.0 MPa, and a bed.
  • the layer density is from 200 kg/m 3 to 1200 kg/m 3 .
  • the catalyst in the fast fluidized bed reactor, has a lower dense phase region and an upper rare phase region in a fluidized state.
  • the dense phase zone is the reaction zone of the fast fluidized bed reactor.
  • the present application provides a fast fluidized bed reactor for co-production of low carbon olefins with methanol and/or dimethyl ether and benzene to produce para-xylene
  • the fast fluidized bed reactor comprising: a reactor housing 2 n reactor feed distributors (3-1 to 3-n), reactor gas-solid separator 4, reactor gas-solid separator 5, reactor heat extractor (6), product gas outlet 7 and reactor Stripper 8, wherein the lower part of the fast fluidized bed reactor 1 is a reaction zone, the upper part of the fast fluidized bed reactor 1 is a dilute phase zone, and n reactor feed distributors (3-1 to 3-n) Placed in the reaction zone from bottom to top, the reactor heat extractor (6) is placed outside the reaction zone or reactor housing 2, and the reactor gas-solids separator 4 and the reactor gas-solids separator 5 are placed in the dilute phase zone or Outside the reactor housing 2, the reactor gas-solid separator 4 is provided with a regenerated catalyst inlet, the catalyst outlet of the reactor gas-solid separator 4 is placed at the
  • the inlet of the reactor gas-solids separator 5 is placed in the dilute phase zone, and the catalyst outlet of the reactor gas-solids separator 5 is placed in the reaction zone, and the reactor gas solids are divided.
  • the gas outlet of the vessel 5 is connected to the product gas outlet 7, and the reactor stripper 8 passes through the reactor shell from the outside to the inside of the fast fluidized bed reactor and opens into the reaction zone of the fast fluidized bed reactor 1
  • the bottom of the reactor stripper 8 is provided with a reactor stripping gas inlet 9, and the bottom of the reactor stripper is provided with a catalyst outlet to be produced.
  • the n reactor feed distributors (3-1 to 3-n) of the fast fluidized bed reactor 1 are placed in the reaction zone from bottom to top, 3 ⁇ n ⁇ 11, said n being The total number of reactor feed distributors.
  • the level of the opening of the reactor stripper 8 inside the reactor housing 2 is higher than the first reactor feed distributor to avoid direct entry of fresh catalyst into the reactor stripper.
  • the reactor gas-solids separator 4 and the reactor gas-solids separator 5 are cyclones.
  • the present application further provides a co-production of light olefins for methanol and/or dimethyl ether and benzene to produce para-xylene.
  • the apparatus comprises the above described fast fluidized bed reactor 1 and a fluidized bed regenerator 14 for regenerating the catalyst.
  • the fluidized bed regenerator 14 is a turbulent fluidized bed regenerator.
  • the fluidized bed regenerator 14 includes a regenerator housing 15, a regenerator feed distributor 16, a regenerator gas solids separator 17, a regenerator heat extractor 18, a flue gas outlet 19, and a regenerator.
  • the stripper 20 wherein the lower portion of the fluidized bed regenerator 14 is a regeneration zone, the upper portion of the fluidized bed regenerator 14 is a dilute phase zone, and the regenerator feed distributor 16 is placed at the bottom of the regeneration zone, the regenerator heats up 18 is placed in the regeneration zone, the regenerator gas-solids separator 17 is placed outside the dilute phase zone or the regenerator housing 15, the inlet of the regenerator gas-solids separator 17 is placed in the dilute phase zone, and the catalyst of the regenerator gas-solids separator 17 The outlet is placed in the regeneration zone, the gas outlet of the regenerator gas-solid separator 17 is connected to the flue gas outlet 19, and the inlet of the regenerator stripper 20 is connected to the bottom of the regenerator housing 15;
  • the catalyst outlet of the reactor stripper 8 is connected to the inlet of the inclined tube 10, and the inlet inclined tube 10 is provided with a slide valve 11 to be produced, and the outlet of the inclined tube 10 is connected to the inlet of the riser tube 12 to be produced.
  • the bottom of the standby riser 12 is provided with a riser gas inlet 13 to be connected, the outlet of the riser riser 12 is connected to the dilute phase zone of the fluidized bed regenerator 14; and the bottom of the regenerator stripper 20 is provided with a regenerator vapor
  • the gas extraction inlet 21, the bottom of the regenerator stripper 20 is connected to the inlet of the regeneration inclined tube 22, and the regeneration inclined tube 22 is provided with a regeneration spool 23, and the outlet of the regeneration inclined tube 22 is connected to the inlet of the regeneration riser 24 to be regenerated.
  • the bottom of the riser 24 is provided with a regeneration lift gas inlet 25, and the outlet of the regeneration riser 24 is connected to the inlet of the reactor gas-solid separator 4.
  • the present application provides a method for co-production of a lower olefin with methanol and/or dimethyl ether and benzene to produce para-xylene, comprising:
  • the stream containing para-xylene and low-carbon olefin products flowing out of the fast fluidized bed reactor 1 is sent to a product separation system to obtain para-xylene, ethylene, propylene, butene, C 5+ chain hydrocarbons, and aromatic hydrocarbon by-products.
  • unconverted methanol, dimethyl ether and benzene, the aromatic by-products comprising toluene, o-xylene, m-xylene, ethylbenzene and C9 + aromatics, unconverted methanol and dimethyl ether from the reactor feed distributor 3 -2 to 3-n are fed to the reaction zone of the fast fluidized bed reactor 1, and the aromatic by-products and unconverted benzene are fed from the reactor feed distributor 3-1 to the reaction zone of the fast fluidized bed reactor 1, Contact with a catalyst to convert to a product;
  • the catalyst to be produced is regenerated by the fluidized bed regenerator 14, and the regenerated catalyst is separated into the bottom of the reaction zone in the fast fluidized bed reactor 1 after gas-solid separation of the reactor gas-solids separator 4.
  • the process described herein is carried out using the apparatus described above for the co-production of lower olefins of methanol and/or dimethyl ether with benzene to produce para-xylene.
  • the catalyst to be produced passes through the reactor stripper 8, the inclined tube 10, the slide valve 11 to be produced, and the riser tube 12 to enter the dilute phase region of the fluidized bed regenerator 14;
  • the regeneration medium passes from the regenerator feed distributor 16 to the regeneration zone of the fluidized bed regenerator 14 and reacts with the catalyst to be activated to generate flue gas containing CO, CO 2 and regenerated catalyst, and the flue gas passes through the regenerator gas.
  • the solid separator 17 is discharged after dust removal;
  • the regenerated catalyst enters the inlet of the reactor gas-solids separator 4 through the regenerator stripper 20, the regeneration inclined tube 22, the regeneration slide valve 23 and the regeneration riser 24, and after the gas-solid separation, the regenerated catalyst enters the fast fluidized bed reactor 1 The bottom of the reaction zone;
  • the reactor stripping gas enters the reactor stripper 8 from the reactor stripping gas inlet 8 and is in countercurrent contact with the catalyst to be produced, and then enters the fast fluidized bed reactor 1; the rising gas to be produced is fed by the rising gas inlet 13; The raw riser 12 is in downstream contact with the catalyst to be produced, and then enters the dilute phase region of the fluidized bed regenerator 14;
  • regenerator stripping gas is fed from the regenerator stripping gas inlet 21 into the regenerator stripper 20 in countercurrent contact with the regenerated catalyst, and then into the fluidized bed regenerator 14; the regenerating propellant gas is passed from the regenerating ascending gas inlet 25 into the regeneration riser 24 and The regenerated catalyst is contacted in a downstream flow and then enters the inlet of the reactor gas-solids separator 4.
  • the method for co-producing low-carbon olefins for methanol and/or dimethyl ether and benzene to produce para-xylene as described in the present application is entered by the reactor feed distributor 3-1 at the bottom of the fast fluidized bed reactor.
  • the mass ratio of the aromatic hydrocarbon to the methanol is more than 0.5, further preferably, more than 1.
  • the oxygenates introduced by the reactor feed distributors 3-2 to 3-n and The reactor feed distributor 3-1 enters a mass ratio of methanol of greater than 1, more preferably greater than 5.
  • the catalyst comprises a HZSM-5 molecular sieve having both the function of alkylation of methanol and/or dimethyl ether with benzene, methanol to alkene and methanol aromatization.
  • the catalyst comprises a HZSM-11 molecular sieve having both the function of alkylation of methanol and/or dimethyl ether with benzene, methanol to alkene and methanol aromatization.
  • the regenerated catalyst has a carbon content of ⁇ 0.5 wt.%, and more preferably, the regenerated catalyst has a carbon content of ⁇ 0.1 wt.%.
  • the reaction conditions of the reaction zone of the fast fluidized bed reactor are: an apparent linear velocity of the gas of 1.0 m/s to 8.0 m/s, a reaction temperature of 350 to 600 ° C, and a reaction pressure of 0.1.
  • Mpa ⁇ 1.0 MPa, bed density is 50 kg / m 3 ⁇ 500 kg / m 3 .
  • the reaction conditions of the fluidized bed regenerator regeneration zone are: the apparent apparent linear velocity of the gas is 0.1 m/s to 2 m/s, the regeneration temperature is 500 ° C to 750 ° C, and the regeneration pressure is 0.1 Mpa. 1.0 MPa, the bed density is 200 kg/m 3 to 1200 kg/m 3 .
  • the regeneration medium is any one or a mixture of any one of air, oxygen-depleted air or water vapor; the reactor stripping gas, the regenerator stripping gas, and the rising gas to be produced.
  • the regeneration lift gas is water vapor or nitrogen.
  • the method for co-producing low-carbon olefins for methanol and/or dimethyl ether and benzene to produce para-xylene as described in the present application the conversion of benzene is higher than 40%, the conversion of methanol is higher than 70%, and the selectivity of para-xylene More than 90%, based on aromatics
  • the yield of xylene in the single pass is higher than 33%, and the selectivity of the low carbon olefin (ethylene + propylene + butene) in the chain hydrocarbon is more than 70%, and a good technical effect is obtained.
  • the main feature of the fast fluidized bed reactor in the present application is that the aromatic hydrocarbon feedstock is introduced from the lowermost reactor feed distributor, the oxygenates are separately introduced by the n reactor feed distributors, and the highly active regenerated catalyst is directly introduced.
  • the aromatic feedstock comprises fresh benzene, unconverted benzene, and aromatic by-products, the oxygenate comprising fresh methanol, unconverted methanol, and dimethyl ether.
  • the catalyst activity is high, which is favorable for the alkylation reaction of benzene and the isomerization reaction of aromatic hydrocarbon by-products, methyl transfer reaction, etc.;
  • the method of multi-stage feeding of oxygenates Only a small portion of the oxygenates are fed from the bottom of the reactor.
  • the low concentration of oxygen in the bottom region and the high concentration of aromatics impair the adsorption competition of oxygenates with fast diffusion rate in the molecular sieve pores.
  • the advantage is that most of the aromatic hydrocarbons are adsorbed in the catalyst in the bottom region; thirdly, most of the oxygenates are fed from the middle to the upper part, and the oxygenate conversion reaction mainly occurs in the upper middle part of the reaction zone, avoiding the bottom zone
  • the highly active regenerated catalyst rapidly reduces the activity due to carbon formation during the MTO reaction; fourth, the higher carbon content of the catalyst in the upper middle region of the reaction zone is beneficial to improve the selectivity of the low carbon olefin in the MTO reaction; Fifth, the multi-stage feed of oxygenates, the concentration distribution of oxygenates in the reaction zone is relatively uniform, providing sufficient alkylation. Reactant adsorbed on the catalyst in aromatic alkylation reaction and contacting alkylation reaction occurs rapidly, improve the conversion of benzene and paraxylene yield.
  • the fast fluidized bed reactor in the present application can coordinate and optimize the competition between methanol and/or dimethyl ether and benzene alkylation reaction and MTO reaction, so as to synergistically increase benzene conversion rate, Para-xylene yield and low carbon olefin yield.
  • the present invention coordinates and optimizes the competition between the alkylation reaction and the MTO reaction by controlling the concentration of methanol and/or dimethyl ether relative to benzene to increase the yield of p-xylene and Low-carbon olefin selectivity to ensure that neither the MTO reaction rapidly consumes most of the methanol and/or dimethyl ether to inhibit the alkylation reaction, nor does it occur due to the high levels of methanol and/or dimethyl ether.
  • MTO reaction occurs in a large amount, and the amount of benzene adsorbed in the catalyst per unit time is low, which is disadvantageous for the alkylation reaction.
  • the fluidized bed reactor and apparatus provided by the present application in a common feed system with a large difference in raw material reaction rates, distributes feeds in different regions through different raw material streams, thereby realizing mass transfer control, thereby coordinating and optimizing
  • the cofeed system increases the reaction yield.
  • FIG. 1 is a schematic diagram of an apparatus for co-production of low carbon olefins from methanol and/or dimethyl ether with benzene to produce p-xylene according to one embodiment of the present application.
  • 1-Fast fluidized bed reactor 2-reactor housing; 3-reactor feed distributor (3-1 ⁇ 3-n); 4-reactor gas-solid separator; 5-reactor gas-solid separation 6-reactor heat extractor; 7-product gas outlet; 8-reactor stripper; 9-reactor stripping gas inlet; 10-waiting inclined tube; 11-waiting slide valve; Raw riser; 13-waiting lift gas inlet; 14-fluidized bed regenerator; 15-regenerator shell; 16-regenerator feed distributor; 17-regenerator gas-solid separator; 18-regenerator Heater; 19-flue gas outlet; 20-regenerator stripper; 21-regenerator stripping gas inlet; 22-regeneration inclined pipe; 23-regeneration slide valve; 24-regeneration riser; 25-regeneration lift gas inlet .
  • FIG. 1 a schematic diagram of a device for co-production of a low-carbon olefin with methanol and/or dimethyl ether and benzene to produce para-xylene is shown in FIG. 1 , the device comprising: a fast fluidized bed reactor 1 It comprises a reactor housing 2, n reactor feed distributors 3-1 to 3-n (the distributor between 3-1 and 3-n in Figure 1 takes 3-i as an example), reactor gas a solid separator 4, a reactor gas-solid separator 5, a reactor heat extractor 6, a product gas outlet 7 and a reactor stripper 8, wherein the lower portion of the fast fluidized bed reactor 1 is a reaction zone, a fast fluidized bed
  • the upper part of the reactor 1 is a dilute phase zone, and the n reactor feed distributors 3-1 to 3-n are arranged in the reaction zone from bottom to top, 2 ⁇ n ⁇ 11, and the reactor heat extractor 6 is placed in the reaction zone.
  • the reactor gas-solid separator 4 and the reactor gas-solid separator 5 are placed outside the dilute phase zone or the reactor housing 2, and the inlet of the reactor gas-solid separator 4 is connected to the regeneration riser 24.
  • the catalyst outlet of the reactor gas-solids separator 4 is placed at the bottom of the reaction zone, and the gas outlet of the reactor gas-solids separator 4 is placed in the dilute phase zone, and the reactor gas-solid separator 5 is introduced. Placed in the dilute phase zone, the catalyst outlet of the reactor gas-solids separator 5 is placed in the reaction zone, the gas outlet of the reactor gas-solids separator 5 is connected to the product gas outlet 7, and the inlet of the reactor stripper 8 is rapidly fluidized.
  • the reaction zone of bed reactor 1 has a higher level than the first reactor feed distributor.
  • the apparatus comprises: a fluidized bed regenerator 14 comprising a regenerator housing 15, a regenerator feed distributor 16, a regenerator gas-solid separator 17, a regenerator heat extractor 18, and a flue gas.
  • the regenerator heat extractor 18 is placed in the regeneration zone, the regenerator gas-solid separator 17 is placed outside the dilute phase zone or the regenerator housing 15, and the inlet of the regenerator gas-solids separator 17 is placed in the dilute phase zone, and the regenerator gas is solidified.
  • the catalyst outlet of the separator 17 is placed in the regeneration zone, the gas outlet of the regenerator gas-solid separator 17 is connected to the flue gas outlet 19, and the inlet of the regenerator stripper 20 is connected to the bottom of the regenerator housing 15.
  • the bottom of the reactor stripper 8 is provided with a reactor stripping gas inlet 9, a reactor stripper 8
  • the bottom of the inclined tube 10 is connected to the inlet of the inclined tube 10, and the outlet tube 10 is provided with a slide valve 11 to be produced.
  • the outlet of the inclined tube 10 is connected to the inlet of the riser tube 12, and the bottom of the riser tube 12 is provided.
  • the bottom of the regenerator stripper 20 is provided with a regenerator stripping gas inlet 21, and the bottom of the regenerator stripper 20 is connected to the inlet of the regenerative inclined tube 22, and the regenerative inclined tube 22 is provided with regenerative sliding
  • the valve 23 the outlet of the regeneration ramp 22 is connected to the inlet of the regeneration riser 24, the bottom of the regeneration riser 24 is provided with a regeneration lift gas inlet 25, and the outlet of the regeneration riser 24 is connected to the inlet of the reactor gas-solid separator 4.
  • the fluidized bed regenerator 14 may be a turbulent fluidized bed regenerator; the reactor gas-solids separator 4, the reactor gas-solid separator 5, and the regenerator gas-solid separator 17 may be Cyclone separator.
  • the method for co-producing low-carbon olefins for methanol and/or dimethyl ether and benzene to produce para-xylene includes:
  • the regenerated catalyst enters the inlet of the reactor gas-solids separator 4 through the regenerator stripper 20, the regeneration inclined tube 22, the regeneration slide valve 23 and the regeneration riser 24, and after the gas-solid separation, the regenerated catalyst enters the fast fluidized bed reactor.
  • the reactor stripping gas is fed from the reactor stripping gas inlet 9 into the reactor stripper 8 in countercurrent contact with the catalyst to be produced, and then enters the fast fluidized bed reactor 1; the rising gas is raised by the rising gas inlet 13 Entering the standby riser 12 and the catalyst to be produced in downstream contact, and then entering the dilute phase region of the fluidized bed regenerator 14;
  • regenerator stripping gas is fed from the regenerator stripping gas inlet 21 into the regenerator stripper 20 in countercurrent contact with the regenerated catalyst and then into the fluidized bed regenerator 14; the regenerative ascending gas is passed from the regenerating ascending gas inlet 25 to the regenerative riser 24 is in downstream contact with the regenerated catalyst and then enters the inlet of the reactor gas-solids separator 4.
  • the apparatus shown in Fig. 1 is employed, but the reactor fluid-solid separator 4 is not included in the fast fluidized bed reactor 1, and the regeneration riser 24 is directly connected to the dilute phase zone of the fast fluidized bed reactor 1.
  • the fast fluidized bed reactor 1 contains one reactor feed distributor 3-1.
  • the reaction conditions of the reaction zone of the fast fluidized bed reactor 1 are: the apparent linear velocity of the gas is about 1.0 m/s, the reaction temperature is about 500 ° C, the reaction pressure is about 0.15 MPa, and the bed density is about 350 kg/m 3 .
  • the reaction conditions of the regeneration zone of the fluidized bed regenerator 14 are: the apparent apparent linear velocity of the gas is about 1.0 m/s, the regeneration temperature is about 650 ° C, the regeneration pressure is about 0.15 MPa, and the bed density is about 350 kg/m 3 .
  • the catalyst contains HZSM-5 molecular sieve, and the regenerated catalyst has a carbon content of about 0.15 wt.%.
  • the regeneration medium is air; the reactor stripping gas, the regenerator stripping gas, the waiting lift gas and the regeneration lift gas are water vapor.
  • the mass ratio of aromatic hydrocarbon to methanol was 0.5.
  • the experimental results are: benzene conversion rate of 17%, methanol conversion rate of 97%, para-xylene selectivity of 99%, aromatics-based para-xylene mass per pass yield of 13%, low-carbon olefins (ethylene + propylene + butyl)
  • the selectivity of the alkene in the chain hydrocarbon is 65%.
  • the fast fluidized bed reactor 1 contains three reactor feed distributors 3-1 to 3-3, and the reactor gas-solid separator 4 is placed outside the reactor housing 2.
  • the reaction conditions of the reaction zone of the fast fluidized bed reactor 1 are: the apparent linear velocity of the gas is about 1.0 m/s, the reaction temperature is about 500 ° C, the reaction pressure is about 0.15 MPa, and the bed density is about 350 kg/m 3 .
  • the reaction conditions of the regeneration zone of the fluidized bed regenerator 14 are: the apparent apparent linear velocity of the gas is about 1.0 m/s, the regeneration temperature is about 650 ° C, the regeneration pressure is about 0.15 MPa, and the bed density is about 350 kg/m 3 .
  • the catalyst contains HZSM-5 molecular sieve, and the regenerated catalyst has a carbon content of about 0.15 wt.%.
  • the regeneration medium is air; the reactor stripping gas, the regenerator stripping gas, the waiting lift gas and the regeneration lift gas are water vapor.
  • the mass ratio of aromatic hydrocarbon to methanol was 2 in the mixture entering the lower reactor feed distributor 3-1 of the fast fluidized bed reactor.
  • the amount ratio of the oxygen-containing compound entering from the reactor feed distributors 3-2 to 3-3 and the methanol entering from the reactor feed distributor 3-1 was 3.
  • the experimental results are: benzene conversion rate of 40%, methanol conversion rate of 94%, p-xylene selectivity of 97%, aromatics-based para-xylene mass per pass yield of 33%, low-carbon olefins (ethylene + propylene + butyl)
  • the selectivity of the alkene in the chain hydrocarbon is 72%.
  • the regenerated catalyst enters the bottom of the fast fluidized bed reactor, while the regenerated catalyst of Example 1 enters the dilute phase zone of the fast fluidized bed reactor;
  • Example 2 Comparing the present example with Example 1, it is known that the catalyst is first exposed to a high concentration of aromatic hydrocarbon raw material, which greatly improves the conversion of benzene, the yield of p-xylene and the selectivity of lower olefins.
  • the fast fluidized bed reactor 1 contains six reactor feed distributors 3-1 to 3-6, and the reactor gas-solid separator 4 is placed inside the reactor housing 2.
  • the reaction conditions of the reaction zone of the fast fluidized bed reactor 1 are: the apparent linear velocity of the gas is about 6.0 m/s, the reaction temperature is about 570 ° C, the reaction pressure is about 0.7 MPa, and the bed density is about 60 kg/m 3 .
  • the reaction conditions of the regeneration zone of the fluidized bed regenerator 14 are: the apparent apparent linear velocity of the gas is about 1.7 m/s, the regeneration temperature is about 600 ° C, the regeneration pressure is about 0.7 MPa, and the bed density is about 220 kg/m 3 .
  • the catalyst contains HZSM-11 molecular sieve, and the regenerated catalyst has a carbon content of about 0.1 wt.%.
  • the regeneration medium is air; the reactor stripping gas, the regenerator stripping gas, the waiting lift gas and the regeneration lift gas are water vapor.
  • the mass ratio of aromatic hydrocarbon to methanol was 4 in the mixture entering the lower reactor feed distributor 3-1 of the fast fluidized bed reactor.
  • the mass ratio of the oxygen-containing compound entering from the reactor feed distributors 3-2 to 3-6 to the methanol entering from the reactor feed distributor 3-1 was 20.
  • the experimental results are: benzene conversion rate of 44%, methanol conversion rate of 76%, para-xylene selectivity of 90%, aromatics-based para-xylene mass per pass yield of 42%, low-carbon olefins (ethylene + propylene + butyl The selectivity of the alkene in the chain hydrocarbon is 73%.
  • the fast fluidized bed reactor 1 contains four reactor feed distributors 3-1 to 3-4, and the reactor gas-solids separator 4 is placed outside the reactor housing 2.
  • the reaction conditions of the reaction zone of the fast fluidized bed reactor 1 are: the apparent apparent linear velocity of the gas is about 3.0 m/s, the reaction temperature is about 420 ° C, the reaction pressure is about 0.3 MPa, and the bed density is about 180 kg/m 3 .
  • the reaction conditions of the regenerating zone of the fluidized bed regenerator 14 are: the apparent apparent linear velocity of the gas is about 1.2 m/s, the regeneration temperature is about 700 ° C, the regeneration pressure is about 0.3 MPa, and the bed density is about 330 kg/m 3 .
  • the catalyst contains HZSM-5 molecular sieve, and the regenerated catalyst has a carbon content of about 0.1 wt.%.
  • the regeneration medium is water vapor; the reactor stripping gas, the regenerator stripping gas, the waiting lift gas, and the regeneration lift gas are nitrogen.
  • the mass ratio of aromatic hydrocarbon to methanol was 3 in the mixture entering the lower reactor feed distributor 3-1 of the fast fluidized bed reactor.
  • the mass ratio of the oxygen-containing compound entering from the reactor feed distributors 3-1 to 3-4 to the methanol entering from the reactor feed distributor 3-1 was 10.
  • the experimental results are: benzene conversion rate of 42%, methanol conversion rate of 85%, para-xylene selectivity of 93%, aromatics-based para-xylene mass per pass yield of 39%, low-carbon olefins (ethylene + propylene + butyl)
  • the selectivity of the alkene in the chain hydrocarbon is 72%.

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Abstract

一种用于甲醇和/或二甲醚与苯制对二甲苯联产低碳烯烃的快速流化床反应器（1）、装置和方法，快速流化床反应器（1）包括第一反应器进料分布器（3-1）和多个第二反应器进料分布器（3-2~3-n），第一反应器进料分布器（3-1）和多个第二反应器进料分布器（3-2~3-n）沿快速流化床反应器（1）的反应区中气体流向依次布置。改善了甲醇和/或二甲醚与苯制对二甲苯联产低碳烯烃过程中MTO反应和烷基化反应之间的竞争问题，实现MTO反应和烷基化反应的协同作用，提高苯转化率、对二甲苯收率和低碳烯烃选择性。

Description

甲醇和/或二甲醚与苯制对二甲苯联产低碳烯烃的装置及方法 技术领域
本申请涉及一种生产对二甲苯(PX)联产低碳烯烃的装置及生产方法，尤其适用于甲醇和/或二甲醚与苯烷基化制对二甲苯联产低碳烯烃的流化床装置及生产方法，属于化学化工领域。
背景技术
对二甲苯(PX)是石化工业的基本有机原料之一，在化纤、合成树脂、农药、医药、高分子材料等众多领域有着广泛的用途。目前对二甲苯生产主要采用甲苯、C9芳烃及混合二甲苯为原料，通过歧化、异构化、吸附分离或深冷分离而制取。由于其产物中的对二甲苯含量由热力学控制，对二甲苯在C8混合芳烃中只占～24％，工艺过程中物料循环处理量大，设备庞大、操作费用高。特别是二甲苯中三个异构体的沸点相差很小，通过常规的蒸馏技术很难得到高纯度的对二甲苯，而必须采用昂贵的吸附分离工艺。近年来，国内外许多专利公开了生产对二甲苯的新路线，其中甲醇和/或二甲醚与苯烷基化技术是高选择性生产对二甲苯的新途径，已经受到了业界的高度重视和极大关注。
低碳烯烃，即乙烯、丙烯和丁烯，是两种基本的石油化工原料，其需求量日益增加。乙烯和丙烯主要以石脑油为原料进行生产，依赖于石油路线。近年来，非石油路线制取乙烯、丙烯越来越受到重视，特别是甲醇转化制取低碳烯烃(MTO)工艺路线，此路线是实现石油替代战略，减轻和缓解我国对石油的需求和依赖的重要途径。
甲苯甲醇烷基化制备对二甲苯以及苯甲醇烷基化制甲苯和二甲苯均是增产芳烃的新途径。其中甲苯甲醇择形烷基化可以制备高选择性对二甲苯产品，但该过程所需的甲苯亦是芳烃联合装置生产对二甲苯的中间原料，市场供应短缺。而苯是芳烃联合装置的副产品，据估算年产80万吨PX芳烃联合装置可副产苯约30万吨，因此以苯为原料与甲醇烷基化生产甲苯和二甲苯将是增产芳烃的一种有效途径。
传统的甲苯烷基化方法包含在反应器的上游将甲醇和/或二甲醚与苯混合，然后将混合物一起供入反应器。反应器种类包含固定床和流化床。为了提高甲苯的传化率，反应物分阶段注入已经在各种固定床和流化床工艺中得到采用。
MTO反应和烷基化反应之间的竞争是影响对苯转化率、对二甲苯收率和低碳烯烃选择性的主要因素。在同一个反应器中同时实现两个反应，工艺简单，但苯转化率较低；在不同的反应器中分别实现两个反应，工艺较为复杂，但苯转化率和对二甲苯收率较高。因此，甲醇和/或二甲醚与苯烷基化制备对二甲苯联产低碳烯烃工艺在工艺配置和反应器设计方面急需较大的突破，从而能够协调、优化烷基化反应和MTO反应之间的竞争，提高苯转化率、对二甲苯收率和低碳烯烃选择性。
上述提及的制备对二甲苯和低碳烯烃的新途径，其反应过程都是酸催化反应。基于ZSM-5分子筛催化剂的苯和/或甲甲醇和/或二甲醚与苯烷基化制对二甲苯反应过程中必然存在甲醇制烯烃反应。在这个反应过程中主要发生如下几个反应：
C6H5-CH3+CH3OH→C6H4-(CH3)2+H2O    (1)
n CH3OH→(CH2)n+n H2O n＝2,3     (2)
C6H6+CH3OH→C6H5-CH3+H2O          (3)
甲醇既是苯和/或甲甲醇和/或二甲醚与苯烷基化反应的原料，也是MTO反应的原料，但MTO反应速率远高于苯和/或甲甲醇和/或二甲醚与苯烷基化反应速率。MTO反应的特点之一是反应速率远高于甲醇和/或二甲醚与苯烷基化反应，另一个重要特点是催化剂结碳后，甲醇转化速率下降，低碳烯烃选择性升高。因此，控制催化剂的结碳，是提高MTO反应过程中低碳烯烃选择性的一个有效途径。
从以上分析可知，本技术领域需要从催化剂设计和反应器设计两个方面来协调、优化烷基化反应和MTO反应之间的竞争，使之协同作用、提高苯转化率、对二甲苯收率和低碳烯烃收率。
发明内容
根据本申请的一个方面，提供了一种用于甲醇和/或二甲醚与苯制对二甲苯联产低碳烯烃的快速流化床反应器，该快速流化床反应器，解决或改善了甲醇和/或二甲醚与苯生产对二甲苯联产低碳烯烃过程中MTO反应和烷基化反应之间的竞争问题，实现MTO反应和烷基化反应的协同作用，通过控制传质和反应，协调、优化烷基化反应和MTO反应的竞争，使之协同作用，提高苯转化率、对二甲苯收率和低碳烯烃选择性。
我们的实验研究表明，苯和/或甲甲醇和/或二甲醚与苯共同进料，原料中甲醇含量较低时，MTO反应快速消耗了大部分的甲醇(烷基化反应物)，抑制了苯和/或甲甲醇和/或二甲醚与苯烷基化反应，苯和/或甲苯转化率较低。原料中的甲醇含量远远过量时，甲醇与苯和/或甲苯在分子筛孔道内扩散速度的不同使得单位时间内苯和/或甲苯的吸附量偏低，也不利于甲醇和/或二甲醚与苯烷基化反应。因此，优化反应区内甲醇和苯和/或甲苯的浓度，是提高苯和/或甲苯转化率和对二甲苯收率的一个有效途径。
用于甲醇和/或二甲醚与苯制对二甲苯联产低碳烯烃的快速流化床反应器，包含第一反应器进料分布器和多个第二反应器进料分布器，所述第一反应器进料分布器和多个第二反应器进料分布器沿所述快速流化床反应器中气体流向依次布置。
优选地，所述第二反应器进料分布器为2至10个。
优选地，所述快速流化床反应器包括第一反应器气固分离器和第二反应器气固分离器，所述第一反应器气固分离器置于稀相区或反应器壳体外部，所述第二反应器气固分离器置于稀相区或反应器壳体外部；
所述第一反应器气固分离器设有再生催化剂入口，所述第一反应器气固分离器的 催化剂出口置于反应区的底部，所述第一反应器气固分离器的气体出口置于所述稀相区；
所述第二反应器气固分离器的入口置于所述稀相区，所述第二反应器气固分离器的催化剂出口置于所述反应区，所述第二反应器气固分离器的气体出口连接于快速流化床反应器的产品气出口；
所述反应区位于所述快速流化床反应器的下部，所述稀相区位于所述快速流化床反应器的上部。
优选地，所述第一反应器气固分离器和第二反应器气固分离器是旋风分离器。
优选地，所述快速流化床反应器包括反应器取热器，反应器取热器置于所述快速流化床反应器壳体内部或外部。
进一步优选地，所述反应器取热器设置于所述多个反应器进料分布器之间。
优选地，所述快速流化床反应器包括反应器汽提器，所述反应器汽提器在快速流化床反应器的底部由外向内穿过反应器壳体并且开口于快速流化床反应器的反应区内，所述反应器汽提器的底部设有反应器汽提气入口和待生催化剂出口。
进一步优选地，所述反应器汽提器在反应器壳体内部的开口的水平高度高于第一反应器进料分布器。
本申请中，低碳烯烃包括乙烯、丙烯、丁烯中的至少一种。
本申请中，进料中的甲醇可以全部或部分的被二甲醚代替，如进料中含有二甲醚，在计算中将二甲醚换算为相同碳原子数的甲醇即可。
本申请中，“甲醇和/或二甲醚”，是指包括三种情况：只有甲醇；或者只有二甲醚；或者甲醇和二甲醚都有。
例如“甲醇和/或二甲醚与苯”，包括三种情况：甲醇和苯；或者二甲醚和苯；或者甲醇、二甲醚和苯。
根据本申请的又一个方面，提供了一种用于甲醇和/或二甲醚与苯制对二甲苯联产低碳烯烃的装置，该装置，解决或改善了甲醇和/或二甲醚与苯生产对二甲苯联产低碳烯烃过程中MTO反应和烷基化反应之间的竞争问题，实现MTO反应和烷基化反应的协同作用，通过控制传质和反应，协调、优化烷基化反应和MTO反应的竞争，使之协同作用，提高苯转化率、对二甲苯收率和低碳烯烃选择性。
所述装置包含以上一个方面所述的快速流化床反应器中的至少一种和用于再生催化剂的流化床再生器。
优选地，所述流化床再生器是湍动流化床再生器，所述流化床再生器包含再生器壳体、再生器气固分离器、再生器取热器和再生器汽提器；流化床再生器的下部是再生区，流化床再生器的上部是再生器稀相区，再生器进料分布器置于再生区的底部，再生器取热器置于再生区，再生器气固分离器置于稀相区或再生器壳体外部；
所述再生器气固分离器的入口置于再生器稀相区，所述再生器气固分离器的催化剂出口置于再生区，再生器汽提器开口于再生器壳体的底部。
优选地，所述流化床再生器包含再生器壳体、再生器进料分布器、再生器气固分离器、再生器取热器、烟气出口和再生器汽提器；
所述流化床再生器的下部是再生区，流化床再生器的上部是稀相区；
再生器进料分布器置于再生区的底部，再生器取热器置于再生区，再生器气固分离器置于稀相区或再生器壳体外部，再生器气固分离器的入口置于稀相区，再生器气固分离器的催化剂出口置于再生区，再生器气固分离器的气体出口连接于烟气出口，再生器汽提器开口于再生器壳体的底部；
所述反应器汽提器的待生催化剂出口连接于待生斜管的入口，待生斜管中设有待生滑阀，待生斜管的出口连接于待生提升管的入口，待生提升管的底部设有待生提升气入口，待生提升管的出口连接于流化床再生器的稀相区；
所述再生器汽提器的底部设有再生器汽提气入口，再生器汽提器的底部连接于再生斜管的入口，再生斜管中设有再生滑阀，再生斜管的出口连接于再生提升管的入口，再生提升管的底部设有再生提升气入口，再生提升管的出口连接于第一反应器气固分离器的再生催化剂入口，所述第一反应器气固分离器置于流化床反应器的稀相区或反应器壳体外部。
根据本申请的又一个方面，提供了一种用于甲醇和/或二甲醚与苯制对二甲苯联产低碳烯烃的方法，该方法，解决或改善了甲醇和/或二甲醚与苯生产对二甲苯联产低碳烯烃过程中MTO反应和烷基化反应之间的竞争问题，实现MTO反应和烷基化反应的协同作用，通过控制传质和反应，协调、优化烷基化反应和MTO反应的竞争，使之协同作用，提高苯转化率、对二甲苯收率和低碳烯烃选择性。
所述甲醇和/或二甲醚与苯制对二甲苯联产低碳烯烃的方法，采用上一个方面所述的快速流化床反应器中的至少一种。
优选地，将含有甲醇和/或二甲醚与苯的原料A由第一反应器进料分布器送入快速流化床反应器的反应区，将含有甲醇和/或二甲醚的原料B分别由多个第二反应器进料分布器送入快速流化床反应器的反应区与催化剂接触，生成含有对二甲苯和低碳烯烃产品的物流C和待生催化剂。
优选地，将所述物流C分离获得对二甲苯、低碳烯烃、C5+链烃、芳烃副产物以及未转化的甲醇、二甲醚和苯；
其中，未转化的甲醇和二甲醚由多个第二反应器进料分布器送入快速流化床反应器的反应区，芳烃副产物和未转化的苯由第一反应器进料分布器送入快速流化床反应器的反应区与催化剂接触。
优选地，所述待生催化剂经流化床再生器再生后，进入快速流化床反应器中反应 区的底部。
优选地，所述的方法，包括如下步骤：
(1)将含有甲醇和/或二甲醚与苯的物流A由快速流化床反应器下方的第一反应器进料分布器送入快速流化床反应器的反应区与催化剂接触；
(2)将含有甲醇和/或二甲醚的物流B分别由2至10个第二反应器进料分布器送入快速流化床反应器的反应区与催化剂接触，生成含有对二甲苯和低碳烯烃产品的物流C和待生催化剂；所述2至10个第二反应器进料分布器依次布置在第一反应器进料分布器的上方；
(3)将步骤(2)得到物流C分离得到未转化的甲醇和二甲醚物流C-1、未转化的苯和芳烃副产物物流C-2；将物流C-1分别由2至10个第二反应器进料分布器送入快速流化床反应器的反应区与催化剂接触；将物流C-2由第一反应器进料分布器送入快速流化床反应器的反应区与催化剂接触；
所述芳烃副产物包含甲苯、邻二甲苯、间二甲苯、乙苯和C9+芳烃；
(4)将步骤(2)得到待生催化剂经流化床再生器再生，再生催化剂经第一反应器气固分离器，气固分离后，进入快速流化床反应器中反应区的底部。
优选地，由第一反应器进料分布器送入快速流化床反应器的混合物中，芳烃的分子摩尔数与甲醇和/或二甲醚的碳摩尔数之比大于0.5。
进一步优选地，由第一反应器进料分布器送入快速流化床反应器的混合物中，芳烃的分子摩尔数与甲醇和/或二甲醚的碳摩尔数之比为0.5～5。
优选地，由多个第二反应器进料分布器进入快速流化床反应器的混合物中的全部含氧化合物和由第一反应器进料分布器进入的甲醇的摩尔比大于1。
本申请中，分子摩尔数是指所述物质中分子数量的摩尔数，碳摩尔数是指所述物质中碳原子数量的摩尔数。
进一步优选地，由多个第二反应器进料分布器进入的全部含氧化合物和由第一反应器进料分布器进入的甲醇的摩尔比为1～20。
进一步优选地，待生催化剂经过反应器汽提器、待生斜管、待生滑阀和待生提升管进入流化床再生器的稀相区；
再生介质通入流化床再生器的再生区，和待生催化剂发生烧炭反应，生成含有CO、CO2的烟气和再生催化剂，烟气经过再生器气固分离器除尘后排放；
再生催化剂经过再生器汽提器、再生斜管、再生滑阀和再生提升管进入第一反应器气固分离器入口，气固分离后，再生催化剂进入快速流化床反应器中反应区的底部；
反应器汽提气由反应器汽提气入口进入反应器汽提器和待生催化剂逆流接触，然后进入快速流化床反应器；待生提升气由待生提升气入口进入待生提升管和待生催化剂顺流接触，然后进入流化床再生器的稀相区；
再生器汽提气由再生器汽提气入口进入再生器汽提器和再生催化剂逆流接触，然 后进入流化床再生器；再生提升气由再生提升气入口进入再生提升管和再生催化剂顺流接触，然后进入第一反应器气固分离器的入口，所述第一反应器气固分离器置于流化床反应器的稀相区或反应器壳体外部。
优选地，所述再生催化剂碳含量≤0.5wt％。
进一步优选地，所述再生介质为空气、贫氧空气或水蒸气中的至少一种；
和/或，所述反应器汽提气、再生器汽提气、待生提升气和再生提升气为水蒸气或氮气。
优选地，所述快速流化床反应器反应区的反应条件为：气体表观线速度为1.0m/s～8.0m/s，反应温度为350℃～600℃，反应压力为0.1Mpa～1.0MPa，床层密度为50kg/m3～500kg/m3。
优选地，所述流化床再生器再生区反应条件为：气体表观线速度为0.1m/s～2m/s，再生温度为500℃～750℃，再生压力为0.1Mpa～1.0MPa，床层密度为200kg/m3～1200kg/m3。
本申请中，所述快速流化床反应器中，催化剂在流态化状态下存在下部的密相区和上部的稀相区。其中密相区为所述快速流化床反应器的反应区。
本申请提供了一种用于甲醇和/或二甲醚与苯生产对二甲苯联产低碳烯烃的快速流化床反应器，所述快速流化床反应器包含：反应器壳体2、n个反应器进料分布器(3-1～3-n)、反应器气固分离器4、反应器气固分离器5、反应器取热器(6)、产品气出口7和反应器汽提器8，其中快速流化床反应器1的下部是反应区，快速流化床反应器1的上部是稀相区，n个反应器进料分布器(3-1～3-n)由下至上排列置于反应区，反应器取热器(6)置于反应区或反应器壳体2外部，反应器气固分离器4和反应器气固分离器5置于稀相区或反应器壳体2外部，反应器气固分离器4设有再生催化剂入口，反应器气固分离器4的催化剂出口置于反应区的底部，反应器气固分离器4的气体出口置于稀相区，反应器气固分离器5的入口置于稀相区，反应器气固分离器5的催化剂出口置于反应区，反应器气固分离器5的气体出口连接于产品气出口7，反应器汽提器8在快速流化床反应器的底部由外向内穿过反应器壳体并且开口于快速流化床反应器1的反应区内，反应器汽提器8的底部设有反应器汽提气入口9，并且反应器汽提器的底部设有待生催化剂出口。
在一个优选实施方式中，快速流化床反应器1的n个反应器进料分布器(3-1～3-n)由下至上置于反应区，3≤n≤11，所述n为反应器进料分布器的总数。
在一个优选实施方式中，反应器汽提器8在反应器壳体2内部的开口的水平高度高于第一反应器进料分布器，以避免新鲜催化剂直接进入反应器汽提器。
在一个优选实施方式中，反应器气固分离器4和反应器气固分离器5是旋风分离器。
本申请进一步提供了一种用于甲醇和/或二甲醚与苯生产对二甲苯联产低碳烯烃 的装置，所述装置包含上述的快速流化床反应器1和用于再生催化剂的流化床再生器14。
在一个优选实施方式中，流化床再生器14是湍动流化床再生器。
在一个优选实施方式中，流化床再生器14包含再生器壳体15、再生器进料分布器16、再生器气固分离器17、再生器取热器18、烟气出口19和再生器汽提器20，其中流化床再生器14的下部是再生区，流化床再生器14的上部是稀相区，再生器进料分布器16置于再生区的底部，再生器取热器18置于再生区，再生器气固分离器17置于稀相区或再生器壳体15外部，再生器气固分离器17的入口置于稀相区，再生器气固分离器17的催化剂出口置于再生区，再生器气固分离器17的气体出口连接于烟气出口19，再生器汽提器20的入口连接于再生器壳体15的底部；
反应器汽提器8的待生催化剂出口连接于待生斜管10的入口，待生斜管10中设有待生滑阀11，待生斜管10的出口连接于待生提升管12的入口，待生提升管12的底部设有待生提升气入口13，待生提升管12的出口连接于流化床再生器14的稀相区；并且再生器汽提器20的底部设有再生器汽提气入口21，再生器汽提器20的底部连接于再生斜管22的入口，再生斜管22中设有再生滑阀23，再生斜管22的出口连接于再生提升管24的入口，再生提升管24的底部设有再生提升气入口25，再生提升管24的出口连接于反应器气固分离器4的入口。
在另一方面，本申请提供了一种用于甲醇和/或二甲醚与苯生产对二甲苯联产低碳烯烃的方法，包括：
将含有甲醇和/或二甲醚与苯的原料由快速流化床反应器1最下方的反应器进料分布器3-1送入快速流化床反应器1的反应区，将甲醇由快速流化床反应器1中的反应器进料分布器3-2至3-n送入快速流化床反应器1的反应区，与催化剂接触，生成含有对二甲苯和低碳烯烃产品的物流和含碳的待生催化剂；
将由快速流化床反应器1流出的含有对二甲苯和低碳烯烃产品的物流送入产品分离系统，经分离获得对二甲苯、乙烯、丙烯、丁烯、C5+链烃、芳烃副产物以及未转化的甲醇、二甲醚和苯，芳烃副产物包含甲苯、邻二甲苯、间二甲苯、乙苯和C9+芳烃，未转化的甲醇和二甲醚由反应器进料分布器3-2至3-n送入快速流化床反应器1的反应区，芳烃副产物和未转化的苯由反应器进料分布器3-1送入快速流化床反应器1的反应区，与催化剂接触，转化为产物；
待生催化剂经流化床再生器14再生，再生催化剂经反应器气固分离器4的气固分离后，进入快速流化床反应器1中反应区的底部。
在一个优选实施方式中，本申请所述的方法使用上述用于甲醇和/或二甲醚与苯生产对二甲苯联产低碳烯烃的装置进行。
在一个优选实施方式中，待生催化剂经过反应器汽提器8、待生斜管10、待生滑阀11和待生提升管12进入流化床再生器14的稀相区；
再生介质从再生器进料分布器16通入流化床再生器14的再生区，和待生催化剂发生烧炭反应，生成含有CO、CO2的烟气和再生催化剂，烟气经过再生器气固分离器17除尘后排放；
再生催化剂经过再生器汽提器20、再生斜管22、再生滑阀23和再生提升管24进入反应器气固分离器4入口，气固分离后，再生催化剂进入快速流化床反应器1中反应区的底部；
反应器汽提气由反应器汽提气入口9进入反应器汽提器8和待生催化剂逆流接触，然后进入快速流化床反应器1；待生提升气由待生提升气入口13进入待生提升管12和待生催化剂顺流接触，然后进入流化床再生器14的稀相区；
再生器汽提气由再生器汽提气入口21进入再生器汽提器20和再生催化剂逆流接触，然后进入流化床再生器14；再生提升气由再生提升气入口25进入再生提升管24和再生催化剂顺流接触，然后进入反应器气固分离器4的入口。
本申请所述的用于甲醇和/或二甲醚与苯生产对二甲苯联产低碳烯烃的方法中，由快速流化床反应器最下方的反应器进料分布器3-1进入的混合物中，芳烃和甲醇的物质量比大于0.5，进一步优选地，大于1。
本申请所述的用于甲醇和/或二甲醚与苯生产对二甲苯联产低碳烯烃的方法中，由反应器进料分布器3-2至3-n进入的含氧化合物和由反应器进料分布器3-1进入的甲醇的物质量比大于1，进一步优选地，大于5。
在一个优选实施方式中，所述催化剂含有HZSM-5分子筛，该催化剂同时具有甲醇和/或二甲醚与苯烷基化、甲醇制烯烃和甲醇芳构化的功能。
在一个优选实施方式中，所述催化剂含有HZSM-11分子筛，该催化剂同时具有甲醇和/或二甲醚与苯烷基化、甲醇制烯烃和甲醇芳构化的功能。
在一个优选实施方式中，所述再生催化剂碳含量<0.5wt.％，进一步优选地，再生催化剂碳含量<0.1wt.％。
在一个优选实施方式中，所述快速流化床反应器反应区反应条件为：气体表观线速度为1.0m/s～8.0m/s，反应温度为350℃～600℃，反应压力为0.1Mpa～1.0MPa，床层密度为50kg/m3～500kg/m3。
在一个优选实施方式中，所述流化床再生器再生区反应条件为：气体表观线速度为0.1m/s～2m/s，再生温度为500℃～750℃，再生压力为0.1Mpa～1.0MPa，床层密度为200kg/m3～1200kg/m3。
在一个优选实施方式中，所述再生介质为空气、贫氧空气或水蒸气中的任意一种或任意几种的混合物；所述反应器汽提气、再生器汽提气、待生提升气和再生提升气为水蒸气或氮气。
本申请所述的用于甲醇和/或二甲醚与苯生产对二甲苯联产低碳烯烃的方法中，苯转化率高于40％，甲醇转化率高于70％，对二甲苯选择性高于90％，基于芳烃的对 二甲苯单程收率高于33％，低碳烯烃(乙烯+丙烯+丁烯)在链烃中的选择性大于70％，取得了较好的技术效果。
本申请中的快速流化床反应器的主要特征在于芳烃原料由最下方的反应器进料分布器进入，含氧化合物由n个反应器进料分布器分别进入，高活性的再生催化剂直接进入反应区底部。芳烃原料包含新鲜苯、未转化的苯以及芳烃副产物，含氧化合物包含新鲜甲醇、未转化的甲醇和二甲醚。其一，在反应区的下部，催化剂活性高，有利于苯的烷基化反应和芳烃副产物的异构化反应、甲基转移反应等；其二，采用含氧化合物多段进料的方式，仅有一小部分的含氧化合物由反应器的底部进料，底部区域含氧化合物浓度低、芳烃浓度高，削弱了在分子筛孔道内扩散速度快的含氧化合物对扩散速度慢的芳烃的吸附竞争优势，保证大部分的芳烃在底部区域内吸附于催化剂中；其三，大部分的含氧化合物由中上部进料，含氧化合物转化反应主要发生在反应区的中上部，避免了底部区域的高活性的再生催化剂由于MTO反应过程中产生的结碳而快速降低活性；其四，反应区的中上部区域中催化剂的结碳量较高，有利于提高MTO反应的低碳烯烃选择性；其五，采用含氧化合物多段进料的方式，反应区域内含氧化合物浓度分布较为均匀，提供了充分的烷基化反应物，吸附于催化剂中的芳烃和烷基化反应物接触后，快速发生烷基化反应，提高苯转化率和对二甲苯收率。
综上所述，本申请中的快速流化床反应器可以协调、优化甲醇和/或二甲醚与苯烷基化反应和MTO反应之间的竞争，使之协同作用、提高苯转化率、对二甲苯收率和低碳烯烃收率。
本申请从反应器设计和工艺配置的角度出发，通过控制甲醇和/或二甲醚相对于苯的浓度来协调、优化烷基化反应和MTO反应之间的竞争，提高对二甲苯收率和低碳烯烃选择性，以确保既不会出现MTO反应快速消耗大部分的甲醇和/或二甲醚从而抑制烷基化反应的情况，也不会发生因甲醇和/或二甲醚含量远远过量、MTO反应大量发生、单位时间催化剂内苯的吸附量偏低从而不利于烷基化反应的情况。
本申请能产生的有益效果包括：
(1)本申请所提供的流化床反应器和装置，在原料反应速率差距较大的共进料体系中，通过不同原料物流在不同区域分布进料，实现传质控制，进而协调、优化共进料体系，提高反应收率。
(2)本申请所提供的甲醇和/或二甲醚与苯生产对二甲苯联产低碳烯烃的方法，同时具有较高的苯转化率和对二甲苯选择性，苯转化率大于40％，对二甲苯选择性大于90％，基于芳烃的对二甲苯单程收率大于33％，取得了较好的技术效果。
附图说明
图1为本申请根据本申请一个实施方案的用于甲醇和/或二甲醚与苯生产对二甲苯联产低碳烯烃的装置示意图。
附图中的附图标记说明如下：
1-快速流化床反应器；2-反应器壳体；3-反应器进料分布器(3-1～3-n)；4-反应器气固分离器；5-反应器气固分离器；6-反应器取热器；7-产品气出口；8-反应器汽提器；9-反应器汽提气入口；10-待生斜管；11-待生滑阀；12-待生提升管；13-待生提升气入口；14-流化床再生器；15-再生器壳体；16-再生器进料分布器；17-再生器气固分离器；18-再生器取热器；19-烟气出口；20-再生器汽提器；21-再生器汽提气入口；22-再生斜管；23-再生滑阀；24-再生提升管；25-再生提升气入口。
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。
如无特别说明，本申请的实施例中的原料和催化剂均通过商业途径购买。
作为本申请的一种实施方式，用于甲醇和/或二甲醚与苯生产对二甲苯联产低碳烯烃的装置示意图如图1所示，该装置包含：快速流化床反应器1，其包含反应器壳体2、n个反应器进料分布器3-1至3-n(图1中3-1至3-n之间的分布器以3-i作为示例)、反应器气固分离器4、反应器气固分离器5、反应器取热器6、产品气出口7和反应器汽提器8，其中快速流化床反应器1的下部是反应区，快速流化床反应器1的上部是稀相区，n个反应器进料分布器3-1至3-n由下至上排列置于反应区，2≤n≤11,反应器取热器6置于反应区或反应器壳体2外部，反应器气固分离器4和反应器气固分离器5置于稀相区或反应器壳体2外部，反应器气固分离器4的入口连接于再生提升管24，反应器气固分离器4的催化剂出口置于反应区的底部，反应器气固分离器4的气体出口置于稀相区，反应器气固分离器5的入口置于稀相区，反应器气固分离器5的催化剂出口置于反应区，反应器气固分离器5的气体出口连接于产品气出口7，反应器汽提器8的入口在快速流化床反应器1的反应区内，其水平高度高于第一反应器进料分布器。
如图1所示，该装置包含：流化床再生器14，其包含再生器壳体15、再生器进料分布器16、再生器气固分离器17、再生器取热器18、烟气出口19和再生器汽提器20，其中流化床再生器14的下部是再生区，流化床再生器14的上部是稀相区，再生器进料分布器16置于再生区的底部，再生器取热器18置于再生区，再生器气固分离器17置于稀相区或再生器壳体15外部，再生器气固分离器17的入口置于稀相区，再生器气固分离器17的催化剂出口置于再生区，再生器气固分离器17的气体出口连接于烟气出口19，再生器汽提器20的入口连接于再生器壳体15的底部。
如图1所示，反应器汽提器8的底部设有反应器汽提气入口9，反应器汽提器8 的底部连接于待生斜管10的入口，待生斜管10中设有待生滑阀11，待生斜管10的出口连接于待生提升管12的入口，待生提升管12的底部设有待生提升气入口13，待生提升管12的出口连接于流化床再生器14的稀相区；
如图1所示，再生器汽提器20的底部设有再生器汽提气入口21，再生器汽提器20的底部连接于再生斜管22的入口，再生斜管22中设有再生滑阀23，再生斜管22的出口连接于再生提升管24的入口，再生提升管24的底部设有再生提升气入口25，再生提升管24的出口连接于反应器气固分离器4的入口。
在本申请的上述实施方案中，流化床再生器14可以是湍动流化床再生器；反应器气固分离器4、反应器气固分离器5和再生器气固分离器17可以是旋风分离器。
作为本申请的一个具体实施方案中，本申请所述的用于甲醇和/或二甲醚与苯生产对二甲苯联产低碳烯烃的方法包括：
a)将含有甲醇和/或二甲醚与苯的原料由快速流化床反应器1最下方的反应器进料分布器3-1送入快速流化床反应器1的反应区，将甲醇由快速流化床反应器1中的反应器进料分布器3-2至3-n送入快速流化床反应器1的反应区，与催化剂接触，生成含有对二甲苯和低碳烯烃产品的物流和含碳的待生催化剂；
b)将由快速流化床反应器1流出的含有对二甲苯和低碳烯烃产品的物流送入产品分离系统，经分离获得对二甲苯、乙烯、丙烯、丁烯、C5+链烃、芳烃副产物以及未转化的甲醇、二甲醚和苯，芳烃副产物包含甲苯、邻二甲苯、间二甲苯、乙苯和C9+芳烃，未转化的甲醇和二甲醚由反应器进料分布器3-2至3-n送入快速流化床反应器1的反应区，芳烃副产物和未转化的苯由反应器进料分布器3-1送入快速流化床反应器1的反应区，与催化剂接触，转化为产物；
c)待生催化剂经过反应器汽提器8、待生斜管10、待生滑阀11和待生提升管12进入流化床再生器14的稀相区；
d)将再生介质从再生器进料分布器16通入流化床再生器14的再生区，再生介质和待生催化剂发生烧炭反应，生成含有CO、CO2的烟气和再生催化剂，烟气经过再生器气固分离器17除尘后排放；
e)再生催化剂经过再生器汽提器20、再生斜管22、再生滑阀23和再生提升管24进入反应器气固分离器4入口，气固分离后，再生催化剂进入快速流化床反应器1中反应区的底部；
f)反应器汽提气由反应器汽提气入口9进入反应器汽提器8和待生催化剂逆流接触，然后进入快速流化床反应器1；待生提升气由待生提升气入口13进入待生提升管12和待生催化剂顺流接触，然后进入流化床再生器14的稀相区；
g)再生器汽提气由再生器汽提气入口21进入再生器汽提器20和再生催化剂逆流接触，然后进入流化床再生器14；再生提升气由再生提升气入口25进入再生提升管24和再生催化剂顺流接触，然后进入反应器气固分离器4的入口。
为更好地说明本申请，便于理解本申请的技术方案，本申请的典型但非限制性的实施例如下。
实施例1
采用图1所示的装置，但快速流化床反应器1中不包含反应器气固分离器4，再生提升管24直接连接于快速流化床反应器1的稀相区。快速流化床反应器1含有1个反应器进料分布器3-1。
快速流化床反应器1反应区反应条件为：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0.15MPa，床层密度约为350kg/m3。
流化床再生器14再生区反应条件为：气体表观线速度约为1.0m/s，再生温度约为650℃，再生压力约为0.15MPa，床层密度约为350kg/m3。
催化剂含有HZSM-5分子筛，再生催化剂碳含量约为0.15wt.％。
再生介质为空气；反应器汽提气、再生器汽提气、待生提升气和再生提升气为水蒸气。
由快速流化床反应器最下方的反应器进料分布器3-1进入的混合物中，芳烃和甲醇的物质量比为0.5。
实验结果为：苯转化率为17％，甲醇转化率为97％，对二甲苯选择性为99％，基于芳烃的对二甲苯质量单程收率为13％，低碳烯烃(乙烯+丙烯+丁烯)在链烃中的选择性为65％。
实施例2
采用图1所示的装置，快速流化床反应器1含有3个反应器进料分布器3-1至3-3，反应器气固分离器4置于反应器壳体2外部。
快速流化床反应器1反应区反应条件为：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0.15MPa，床层密度约为350kg/m3。
流化床再生器14再生区反应条件为：气体表观线速度约为1.0m/s，再生温度约为650℃，再生压力约为0.15MPa，床层密度约为350kg/m3。
催化剂含有HZSM-5分子筛，再生催化剂碳含量约为0.15wt.％。
再生介质为空气；反应器汽提气、再生器汽提气、待生提升气和再生提升气为水蒸气。
由快速流化床反应器最下方的反应器进料分布器3-1进入的混合物中，芳烃和甲醇的物质量比为2。
由反应器进料分布器3-2至3-3进入的含氧化合物和由反应器进料分布器3-1进入的甲醇的物质的量比为3。
实验结果为：苯转化率为40％，甲醇转化率为94％，对二甲苯选择性为97％，基于芳烃的对二甲苯质量单程收率为33％，低碳烯烃(乙烯+丙烯+丁烯)在链烃中的选择性为72％。
本案例和实施例1的区别在于：
①再生催化剂进入快速流化床反应器的底部，而实施例1中再生催化剂进入快速流化床反应器的稀相区；
②甲醇由3个反应器进料分布器(3-1至3-3)分别进入，而实施例1中甲醇由1个反应器进料分布器3-1进入。
对比本实施例和实施例1可知，催化剂首先接触高浓度的芳烃原料，大幅度的提高了苯转化率、对二甲苯收率和低碳烯烃选择性。
实施例3
采用图1所示的装置，快速流化床反应器1含有6个反应器进料分布器3-1至3-6，反应器气固分离器4置于反应器壳体2内部。
快速流化床反应器1反应区反应条件为：气体表观线速度约为6.0m/s，反应温度约为570℃，反应压力约为0.7MPa，床层密度约为60kg/m3。
流化床再生器14再生区反应条件为：气体表观线速度约为1.7m/s，再生温度约为600℃，再生压力约为0.7MPa，床层密度约为220kg/m3。
催化剂含有HZSM-11分子筛，再生催化剂碳含量约为0.1wt.％。
再生介质为空气；反应器汽提气、再生器汽提气、待生提升气和再生提升气为水蒸气。
由快速流化床反应器最下方的反应器进料分布器3-1进入的混合物中，芳烃和甲醇的物质量比为4。
由反应器进料分布器3-2至3-6进入的含氧化合物和由反应器进料分布器3-1进入的甲醇的物质量比为20。
实验结果为：苯转化率为44％，甲醇转化率为76％，对二甲苯选择性为90％，基于芳烃的对二甲苯质量单程收率为42％，低碳烯烃(乙烯+丙烯+丁烯)在链烃中的选择性为73％。
实施例4
采用图1所示的装置，快速流化床反应器1含有4个反应器进料分布器3-1至3-4，反应器气固分离器4置于反应器壳体2外部。
快速流化床反应器1反应区反应条件为：气体表观线速度约为3.0m/s，反应温度约为420℃，反应压力约为0.3MPa，床层密度约为180kg/m3。
流化床再生器14再生区反应条件为：气体表观线速度约为1.2m/s，再生温度约为 700℃，再生压力约为0.3MPa，床层密度约为330kg/m3。
催化剂含有HZSM-5分子筛，再生催化剂碳含量约为0.1wt.％。
再生介质为水蒸气；反应器汽提气、再生器汽提气、待生提升气和再生提升气为氮气。
由快速流化床反应器最下方的反应器进料分布器3-1进入的混合物中，芳烃和甲醇的物质量比为3。
由反应器进料分布器3-1至3-4进入的含氧化合物和由反应器进料分布器3-1进入的甲醇的物质量比为10。
实验结果为：苯转化率为42％，甲醇转化率为85％，对二甲苯选择性为93％，基于芳烃的对二甲苯质量单程收率为39％，低碳烯烃(乙烯+丙烯+丁烯)在链烃中的选择性为72％。
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 (24)

  1. 一种用于甲醇和/或二甲醚与苯制对二甲苯联产低碳烯烃的快速流化床反应器，其特征在于，所述快速流化床反应器包含第一反应器进料分布器和多个第二反应器进料分布器，所述第一反应器进料分布器和多个第二反应器进料分布器沿所述快速流化床反应器中气体流向依次布置。
  2. 根据权利要求1所述的快速流化床反应器，其特征在于，所述第二反应器进料分布器为2至10个。
  3. 根据权利要求1所述的快速流化床反应器，其特征在于，所述快速流化床反应器包括第一反应器气固分离器和第二反应器气固分离器，所述第一反应器气固分离器置于稀相区或反应器壳体外部，所述第二反应器气固分离器置于稀相区或反应器壳体外部；
    所述第一反应器气固分离器设有再生催化剂入口，所述第一反应器气固分离器的催化剂出口置于反应区的底部，所述第一反应器气固分离器的气体出口置于所述稀相区；
    所述第二反应器气固分离器的入口置于所述稀相区，所述第二反应器气固分离器的催化剂出口置于所述反应区，所述第二反应器气固分离器的气体出口连接于快速流化床反应器的产品气出口；
    所述反应区位于所述快速流化床反应器的下部，所述稀相区位于所述快速流化床反应器的上部。
  4. 根据权利要求3所述的快速流化床反应器，其特征在于，所述第一反应器气固分离器和第二反应器气固分离器是旋风分离器。
  5. 根据权利要求1所述的快速流化床反应器，其特征在于，所述快速流化床反应器包括反应器取热器，反应器取热器置于所述快速流化床反应器壳体内部或外部。
  6. 根据权利要求5所述的快速流化床反应器，其特征在于，所述反应器取热器设置于所述多个反应器进料分布器之间。
  7. 根据权利要求1所述的快速流化床反应器，其特征在于，所述快速流化床反应器包括反应器汽提器，所述反应器汽提器在快速流化床反应器的底部由外向内穿过反应器壳体并且开口于快速流化床反应器的反应区内，所述反应器汽提器的底部设有反应器汽提气入口和待生催化剂出口。
  8. 根据权利要求7所述的快速流化床反应器，其特征在于，所述反应器汽提器在 反应器壳体内部的开口的水平高度高于第一反应器进料分布器。
  9. 一种用于甲醇和/或二甲醚与苯制对二甲苯联产低碳烯烃的装置，其特征在于，所述装置包含权利要求1至8任一项所述的快速流化床反应器中的至少一种和用于再生催化剂的流化床再生器。
  10. 根据权利要求9所述的装置，其特征在于，所述流化床再生器是湍动流化床再生器，所述流化床再生器包含再生器壳体、再生器气固分离器、再生器取热器和再生器汽提器；流化床再生器的下部是再生区，流化床再生器的上部是再生器稀相区，再生器进料分布器置于再生区的底部，再生器取热器置于再生区，再生器气固分离器置于稀相区或再生器壳体外部；
    所述再生器气固分离器的入口置于再生器稀相区，所述再生器气固分离器的催化剂出口置于再生区，再生器汽提器开口于再生器壳体的底部。
  11. 根据权利要求9所述的装置，其特征在于，所述流化床再生器包含再生器壳体、再生器进料分布器、再生器气固分离器、再生器取热器、烟气出口和再生器汽提器；
    所述流化床再生器的下部是再生区，流化床再生器的上部是稀相区；
    再生器进料分布器置于再生区的底部，再生器取热器置于再生区，再生器气固分离器置于稀相区或再生器壳体外部，再生器气固分离器的入口置于稀相区，再生器气固分离器的催化剂出口置于再生区，再生器气固分离器的气体出口连接于烟气出口，再生器汽提器开口于再生器壳体的底部；
    所述反应器汽提器的待生催化剂出口连接于待生斜管的入口，待生斜管中设有待生滑阀，待生斜管的出口连接于待生提升管的入口，待生提升管的底部设有待生提升气入口，待生提升管的出口连接于流化床再生器的稀相区；
    所述再生器汽提器的底部设有再生器汽提气入口，再生器汽提器的底部连接于再生斜管的入口，再生斜管中设有再生滑阀，再生斜管的出口连接于再生提升管的入口，再生提升管的底部设有再生提升气入口，再生提升管的出口连接于第一反应器气固分离器的再生催化剂入口，所述第一反应器气固分离器置于流化床反应器的稀相区或反应器壳体外部。
  12. 一种甲醇和/或二甲醚与苯制对二甲苯联产低碳烯烃的方法，其特征在于，采用权利要求1至8任一项所述快速流化床反应器中的至少一种。
  13. 根据权利要求12所述的方法，其特征在于，将含有甲醇和/或二甲醚与甲苯的原料A由第一反应器进料分布器送入快速流化床反应器的反应区，将含有甲醇和/或二甲醚的原料B分别由多个第二反应器进料分布器送入快速流化床反应器的反应区 与催化剂接触，生成含有对二甲苯和低碳烯烃产品的物流C和待生催化剂。
  14. 根据权利要求13所述的方法，其特征在于，将所述物流C分离获得对二甲苯、低碳烯烃、C5+链烃、芳烃副产物以及未转化的甲醇、二甲醚和苯；
    其中，未转化的甲醇和二甲醚由多个第二反应器进料分布器送入快速流化床反应器的反应区，芳烃副产物和未转化的苯由第一反应器进料分布器送入快速流化床反应器的反应区与催化剂接触。
  15. 根据权利要求13所述的方法，其特征在于，所述待生催化剂经流化床再生器再生后，进入快速流化床反应器中反应区的底部。
  16. 根据权利要求12所述的方法，其特征在于，包括如下步骤：
    (1)将含有甲醇和/或二甲醚与苯的物流A由快速流化床反应器下方的第一反应器进料分布器送入快速流化床反应器的反应区与催化剂接触；
    (2)将含有甲醇和/或二甲醚的物流B分别由2至10个第二反应器进料分布器送入快速流化床反应器的反应区与催化剂接触，生成含有对二甲苯和低碳烯烃产品的物流C和待生催化剂；所述2至10个第二反应器进料分布器依次布置在第一反应器进料分布器的上方；
    (3)将步骤(2)得到物流C经分离得到未转化的甲醇和二甲醚物流C-1、未转化的苯和芳烃副产物物流C-2；将物流C-1分别由2至10个第二反应器进料分布器送入快速流化床反应器的反应区与催化剂接触；将物流C-2由第一反应器进料分布器送入快速流化床反应器的反应区与催化剂接触；
    所述芳烃副产物包含甲苯、邻二甲苯、间二甲苯、乙苯和C9+芳烃；
    (4)将步骤(2)得到待生催化剂经流化床再生器再生，再生催化剂经第一反应器气固分离器，气固分离后，进入快速流化床反应器中反应区的底部。
  17. 根据权利要求13或16所述的方法，其特征在于，由第一反应器进料分布器送入快速流化床反应器的混合物中，芳烃的分子摩尔数与甲醇和/或二甲醚的碳摩尔数之比大于0.5。
  18. 根据权利要求13或16所述的方法，其特征在于，由多个第二反应器进料分布器进入快速流化床反应器的混合物中的全部含氧化合物和由第一反应器进料分布器进入的甲醇的摩尔比大于1。
  19. 根据权利要求16所述的方法，其特征在于，催化剂再生采用权利要求9任一项所述的装置中的至少一种。
  20. 根据权利要求16所述的方法，其特征在于，待生催化剂经过反应器汽提器、 待生斜管、待生滑阀和待生提升管进入流化床再生器的稀相区；
    再生介质通入流化床再生器的再生区，和待生催化剂发生烧炭反应，生成含有CO、CO2的烟气和再生催化剂，烟气经过再生器气固分离器除尘后排放；
    再生催化剂经过再生器汽提器、再生斜管、再生滑阀和再生提升管进入第一反应器气固分离器入口，气固分离后，再生催化剂进入快速流化床反应器中反应区的底部；
    反应器汽提气由反应器汽提气入口进入反应器汽提器和待生催化剂逆流接触，然后进入快速流化床反应器；待生提升气由待生提升气入口进入待生提升管和待生催化剂顺流接触，然后进入流化床再生器的稀相区；
    再生器汽提气由再生器汽提气入口进入再生器汽提器和再生催化剂逆流接触，然后进入流化床再生器；再生提升气由再生提升气入口进入再生提升管和再生催化剂顺流接触，然后进入第一反应器气固分离器的入口，所述第一反应器气固分离器置于流化床反应器的稀相区或反应器壳体外部。
  21. 根据权利要求20所述的方法，其特征在于，所述再生催化剂碳含量≤0.5wt％。
  22. 根据权利要求20所述的方法，其特征在于，所述再生介质为空气、贫氧空气或水蒸气中的至少一种；和/或，所述反应器汽提气、再生器汽提气、待生提升气和再生提升气为水蒸气和/或氮气。
  23. 根据权利要求20所述的方法，其特征在于，所述快速流化床反应器反应区的反应条件为：气体表观线速度为1.0m/s～8.0m/s，反应温度为350℃～600℃，反应压力为0.1Mpa～1.0MPa，床层密度为50kg/m3～500kg/m3。
  24. 根据权利要求20所述的方法，其特征在于，所述流化床再生器再生区反应条件为：气体表观线速度为0.1m/s～2m/s，再生温度为500℃～750℃，再生压力为0.1Mpa～1.0MPa，床层密度为200kg/m3～1200kg/m3。
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