WO2021180150A1 - 含氧化合物生产低碳烯烃的方法 - Google Patents

含氧化合物生产低碳烯烃的方法 Download PDF

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WO2021180150A1
WO2021180150A1 PCT/CN2021/080114 CN2021080114W WO2021180150A1 WO 2021180150 A1 WO2021180150 A1 WO 2021180150A1 CN 2021080114 W CN2021080114 W CN 2021080114W WO 2021180150 A1 WO2021180150 A1 WO 2021180150A1
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catalyst
weight
reaction zone
fluidized bed
carbon
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PCT/CN2021/080114
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English (en)
French (fr)
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齐国祯
曹静
李晓红
王洪涛
金永明
高攀
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中国石油化工股份有限公司
中国石油化工股份有限公司上海石油化工研究院
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Priority claimed from CN202010173926.6A external-priority patent/CN113387763B/zh
Priority claimed from CN202010173939.3A external-priority patent/CN113387765B/zh
Application filed by 中国石油化工股份有限公司, 中国石油化工股份有限公司上海石油化工研究院 filed Critical 中国石油化工股份有限公司
Priority to CN202180020813.5A priority Critical patent/CN115605449A/zh
Priority to AU2021233959A priority patent/AU2021233959A1/en
Priority to US17/906,230 priority patent/US20230118436A1/en
Priority to BR112022018248A priority patent/BR112022018248A2/pt
Publication of WO2021180150A1 publication Critical patent/WO2021180150A1/zh
Priority to ZA2022/11242A priority patent/ZA202211242B/en

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    • 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
    • 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
    • C07C1/22—Preparation of hydrocarbons from one or more compounds, none of them being a hydrocarbon starting from organic compounds containing only oxygen atoms as heteroatoms by reduction
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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]
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    • B01J8/0055—Separating solid material from the gas/liquid stream using cyclones
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    • B01J8/08—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with moving particles
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    • B01J8/125—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with moving particles moved by gravity in a downward flow with multiple sections one above the other separated by distribution aids, e.g. reaction and regeneration sections
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    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
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    • B01J8/18—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles
    • B01J8/1818—Feeding of the fluidising gas
    • B01J8/1827—Feeding of the fluidising gas the fluidising gas being a reactant
    • B—PERFORMING OPERATIONS; TRANSPORTING
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    • B01J8/1845—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles with particles moving upwards while fluidised
    • B01J8/1863—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles with particles moving upwards while fluidised followed by a downward movement outside the reactor and subsequently re-entering it
    • 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
    • B01J8/00—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes
    • B01J8/18—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles
    • 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
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
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    • B01J8/00—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes
    • B01J8/18—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles
    • 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
    • B01J8/30—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 the edge of a lower bed projecting beyond the edge of the superjacent bed
    • 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
    • B01J2208/00—Processes carried out in the presence of solid particles; Reactors therefor
    • B01J2208/00796—Details of the reactor or of the particulate material
    • B01J2208/00991—Disengagement zone in fluidised-bed reactors
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C2529/00—Catalysts comprising molecular sieves
    • C07C2529/82—Phosphates
    • C07C2529/84—Aluminophosphates containing other elements, e.g. metals, boron
    • C07C2529/85—Silicoaluminophosphates (SAPO compounds)
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C2529/00—Catalysts comprising molecular sieves
    • C07C2529/89—Silicates, aluminosilicates or borosilicates of titanium, zirconium or hafnium
    • Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
    • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
    • Y02P20/00—Technologies relating to chemical industry
    • Y02P20/50—Improvements relating to the production of bulk chemicals
    • Y02P20/52—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
    • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
    • Y02P20/00—Technologies relating to chemical industry
    • Y02P20/50—Improvements relating to the production of bulk chemicals
    • Y02P20/584—Recycling of catalysts
    • 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
    • Y02P30/00—Technologies relating to oil refining and petrochemical industry
    • Y02P30/20—Technologies relating to oil refining and petrochemical industry using bio-feedstock
    • 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
    • Y02P30/00—Technologies relating to oil refining and petrochemical industry
    • Y02P30/40—Ethylene production

Definitions

  • the invention relates to a method for producing low-carbon olefins from oxygen-containing compounds.
  • Low-carbon olefins mainly ethylene and propylene
  • ethylene and propylene are two important basic chemical raw materials, and their demand is constantly increasing.
  • ethylene and propylene are produced through petroleum routes.
  • the cost of producing ethylene and propylene from petroleum resources is increasing.
  • people have begun to vigorously develop alternative raw materials conversion technology to produce ethylene and propylene.
  • an important type of alternative raw materials for the production of low-carbon olefins are oxygen-containing compounds, such as alcohols (methanol, ethanol), ethers (dimethyl ether, methyl ethyl ether), esters (dimethyl carbonate, methyl formate). Esters), etc.
  • oxygen-containing compounds can be converted from coal, natural gas, biomass and other energy sources. Certain oxygen-containing compounds can already be produced on a larger scale, such as methanol, which can be made from coal or natural gas. The process is very mature and can achieve a production scale of millions of tons. Due to the wide range of sources of oxygenates, coupled with the economics of the process of converting to low-carbon olefins, the process of converting oxygenates to olefins (OTO), especially the process of converting methanol to olefins (MTO), has been increasingly affected. More and more attention.
  • Document US6166282 discloses a technology and reactor for the conversion of methanol to low-carbon olefins.
  • a fast fluidized bed reactor is used. After the gas phase is reacted in a dense phase reaction zone with a lower gas velocity, it rises to a fast partition with a rapidly decreasing inner diameter. Afterwards, a special gas-solid separation device is used to initially separate most of the entrained catalyst. Since the product gas and the catalyst are quickly separated after the reaction, the occurrence of secondary reactions is effectively prevented.
  • the inner diameter of the fast fluidized bed reactor and the required storage capacity of the catalyst are greatly reduced. In this method, the yield of low-carbon olefin carbon base is generally around 77%.
  • Document CN101328103A discloses a method for converting methanol or dimethyl ether into low-carbon olefins.
  • the raw material including methanol or dimethyl ether enters the reaction zone of a fluidized bed reactor and contacts with a catalyst including silicoaluminophosphate molecular sieve; wherein , Under the conditions that the reaction pressure is 0.05-1MPa in gauge, the average temperature of the reaction zone is 450-550°C, and the average gas velocity of the empty tower in the reaction zone is 0.8-2.0 m/s, the average density of the reaction zone is 20-300 kg /M3, the average carbon deposit of the catalyst in the reaction zone is 1.5-4.5% by weight.
  • the carbon-based selectivity of low-carbon olefins is up to 81.51% by weight.
  • the inventor of the present invention found that in the process of converting oxygenates (especially methanol) into low-carbon olefins in the prior art, a certain amount of carbon deposits on the catalyst is necessary to ensure high selectivity of low-carbon olefins.
  • oxygenates especially methanol
  • a certain amount of carbon deposits on the catalyst is necessary to ensure high selectivity of low-carbon olefins.
  • the coke deposit of the catalyst in the reaction zone is an average concept. The mixing quality of the coke deposit catalyst is very important to improve the selectivity of low-carbon olefins.
  • the inventor of the present invention further found that the reaction performance of the raw materials on catalysts with different coke deposits varies greatly, especially when the coke deposits are less than 3wt% of the catalyst, the selectivity of low-carbon olefins is greatly reduced.
  • the inventor also unexpectedly discovered that it is essential to control the difference in the amount of coke deposited on the mixed catalyst in the reaction zone to achieve the desired technical effect.
  • the present invention has been completed based on these findings.
  • the present invention provides, for example, embodiments in the following aspects:
  • a method for producing low-carbon olefins from oxygenates which includes the steps of contacting a feedstock containing oxygenates with a molecular sieve catalyst in a fluidized bed reaction zone to produce products containing ethylene and/or propylene under effective conditions;
  • the effective conditions include the following: in the fluidized bed reaction zone, based on the mass of the molecular sieve in the catalyst, the ratio of the mass of the catalyst for controlling various coke deposits to the total catalyst mass in the fluidized bed reaction zone is as follows:
  • the ratio of the mass of the catalyst with a coke deposit less than 3% by weight to the total catalyst mass in the fluidized bed reaction zone is 1-20% by weight, preferably 1-15% by weight, 1.5-10% by weight, or 2-5% by weight %;
  • the catalyst with a coke deposit amount of 3 to less than 5% by weight accounts for 10 to 70% by weight, preferably 15 to 60% by weight, 20 to 50% by weight, or 30 to 45% by weight;
  • the catalyst with a coke deposit amount of 5-10% by weight accounts for 10 to 88% by weight, preferably 15 to 80% by weight, 20 to 70% by weight, or 30 to 60% by weight.
  • a method for producing low-carbon olefins from oxygenates which includes the steps of contacting oxygenate raw materials with molecular sieve catalysts in a fluidized bed reaction zone to produce ethylene and/or propylene products under effective conditions;
  • the raw material enters the fluidized bed reactor reaction zone through the distribution device, and the maximum difference between the distribution device and the carbon deposit on the catalyst particles within the range of 1/2 the bed height of the distribution device is less than 8% by weight , Preferably less than 5% by weight, or less than 3% by weight.
  • the ratio of the mass of the catalyst with a coke deposit less than 3% by weight to the total catalyst mass in the fluidized bed reaction zone is 1-20% by weight, preferably 1-15% by weight, 1.5-10% by weight, or 2-5% by weight %;
  • the catalyst with a coke deposit amount of 3 to less than 5% by weight accounts for 10 to 70% by weight, preferably 15 to 60% by weight, 20 to 50% by weight, or 30 to 45% by weight;
  • the catalyst with a coke deposit amount of 5-10% by weight accounts for 10 to 88% by weight, preferably 15 to 80% by weight, 20 to 70% by weight, or 30 to 60% by weight.
  • the molecular sieve is a silico-aluminum-phosphorus molecular sieve, preferably SAPO-18, SAPO-34, SAPO-5 or a combination thereof.
  • the fluidized bed reaction zone is a dense phase, turbulent or fast fluidized type, preferably a fast fluidized type.
  • the effective conditions further include: a reaction temperature of 400 to 550°C, and a reaction pressure of 0 to 1 MPaG.
  • the lower part of the reaction zone is provided with a regeneration pipeline outlet
  • the regeneration pipeline outlet is provided with a catalyst distributor
  • the catalyst distributor is substantially radially along the reaction zone. Horizontal layout.
  • a fluidized bed reactor for implementing the method for producing low-carbon olefins from oxygenates according to any one of the foregoing embodiments, comprising:
  • the reaction zone is used to receive the methanol feedstock and make it contact with the catalyst to generate olefin products, wherein the process makes the catalyst at least partially deactivated to obtain a spent catalyst;
  • Gas-solid rapid separation equipment used to separate the spent catalyst from the reaction zone
  • the cyclone separator is used to receive the gas-phase product separated by the gas-solid rapid separation device and part of the spent catalyst that has not been separated by the gas-solid rapid separation device for re-separation;
  • Stripping zone for receiving the spent catalyst from the legs of the cyclone separator
  • the catalyst outer circulation inclined pipe is used to return at least part of the stripped spent catalyst from the stripping zone to the bottom of the reaction zone.
  • Fig. 1 is a schematic diagram of the fluidized bed reactor in the method of the present invention.
  • 1 is the reactor raw material feed pipeline
  • both high raw material conversion rate and high low-carbon olefin selectivity can be ensured, and the low-carbon olefin selectivity can reach more than 84%.
  • reaction zone is used for fluidized bed reactors.
  • a fluidized bed reactor includes a reaction zone, an inlet zone, and a separation zone.
  • the "inlet zone” is the section where raw materials and catalysts are introduced into the reactor.
  • the “reaction zone” is the section in the reactor where the feed and the catalyst are in contact with each other under conditions effective to convert the oxygenated compounds of the feed into light olefin products.
  • the “separation zone” is the section in the reactor where the catalyst and any other solids in the reactor are separated from the product.
  • the reaction zone is located between the inlet zone and the separation zone.
  • the "gas phase” includes one or more of the heated and vaporized raw material methanol, diluent gas (such as water vapor), and gas phase reaction products (such as light olefins, C4 hydrocarbons, etc.) .
  • the remainder is the catalyst volume.
  • the ratio of the volume of the gas phase to the volume of the catalyst indicates the volume fraction of the catalyst solid particles in the gas-solid two-phase mixture in the reaction zone.
  • the method for calculating the amount of soot on the catalyst is the mass of soot on the catalyst divided by the mass of the catalyst.
  • the method for measuring the quality of coke deposits on the catalyst is as follows: weigh 0.1 to 1 g of the catalyst with carbon and burn it in a high-temperature carbon analyzer, and measure the mass of carbon dioxide generated by the combustion by infrared to obtain the quality of coke deposits on the catalyst. In order to determine the amount of carbon deposited on the catalyst in the reaction zone, a small amount of catalyst can be drawn out continuously or periodically or directly from various positions in the reaction zone.
  • the catalyst in the fluidized bed reaction zone is deactivated to form a spent catalyst (spent agent), and the spent catalyst enters the regenerator to be regenerated to form a regenerated catalyst (regenerant). Return to the fluidized bed reaction zone.
  • any two or more embodiments of the present invention can be combined arbitrarily, and the technical solutions formed thereby belong to a part of the original disclosure of this specification and also fall into the protection scope of the present invention.
  • the present invention relates to a method for producing low-carbon olefins from oxygen-containing compounds.
  • the method includes the steps of contacting oxygen-containing compound raw materials with a molecular sieve catalyst in a fluidized bed reaction zone to generate products containing ethylene and propylene under effective conditions.
  • the oxygen-containing compound raw material includes an aliphatic alcohol containing 1-20 carbon atoms, preferably 1-10 carbon atoms, more preferably 1-4 carbon atoms.
  • an aliphatic alcohol containing 1-20 carbon atoms, preferably 1-10 carbon atoms, more preferably 1-4 carbon atoms.
  • the effective conditions include: in the fluidized bed reaction zone, based on the mass of the molecular sieve in the catalyst, the mass of the catalyst with a coke deposit less than 3% by weight accounts for the total mass of the catalyst in the fluidized bed reaction zone.
  • the ratio of the catalyst mass is 1 to 20% by weight, preferably 1 to 15% by weight, preferably 1.5 to 10% by weight, more preferably 2 to 5% by weight.
  • the ratio of the catalyst with a carbon deposition of less than 3% by weight to the total catalyst in the fluidized bed reaction zone is controlled by adjusting the regenerator circulation volume and the carbon deposit of the regenerator (that is, the degree of coke burning of the regenerator).
  • the catalyst with a coke deposit less than 3% by weight is uniformly distributed in the fluidized bed reaction zone.
  • the effective conditions further include: the ratio of the volume of the gas phase to the volume of the total catalyst in the fluidized bed reaction zone is 1-15, preferably 5-12.
  • the effective conditions further include: a reaction temperature of 400 to 550° C., and a reaction pressure of 0 to 1 MPa.
  • the molecular sieve is a silico-aluminum-phosphorus molecular sieve, preferably SAPO-18, SAPO-34, SAPO-5 or a combination thereof, more preferably SAPO-34.
  • SAPO-18, SAPO-34, SAPO-5 or a combination thereof more preferably SAPO-34.
  • SAPO-34 The preparation method of SAPO molecular sieve or SAPO molecular sieve catalyst is well known in the art.
  • the fluidized bed reaction zone is a dense phase, turbulent or fast fluidized type, preferably a fast fluidized type.
  • the difference between the amount of coke deposited on the spent catalyst and the regenerated catalyst is not more than 7% by weight, preferably not more than 6% by weight, and more preferably not more than 5% by weight.
  • the coke deposit of the spent catalyst can be directly controlled by the reaction, and the coke deposit of the regenerated catalyst can be controlled by the degree of regeneration of the regenerator.
  • the gas phase and the catalyst in the fluidized bed reaction zone are quickly separated by a separation device after the reaction is completed or after leaving the fluidized bed reaction zone.
  • the separation device is preferably a cyclone separator.
  • FIG. 1 An exemplary embodiment 1 of the present invention is shown in FIG. 1.
  • the fluidized bed reactor used in the present invention is a fast fluidized bed, a dense phase fluidized bed, or a turbulent fluidized bed, and the stream including the methanol raw material enters the reactor through the feed line 1 and passes through the raw material distribution equipment 6 enters the reaction zone 2 and contacts with the molecular sieve catalyst to react to produce products containing low-carbon olefins, which make the catalyst at least partially deactivated to form a spent catalyst. Carry the spent catalyst through the gas-solid rapid separation zone 3 and enter the reactor separation zone 9.
  • the spent catalyst separated by the gas-solid rapid separation zone 3 and the cyclone separator 8 is combined, and after being stripped in the stripping zone 4, it is divided into two parts, and one part is returned to the bottom of the reaction zone 2 through the catalyst outer circulation inclined pipe 5; A part of it passes through the inclined tube 14 to be regenerated and enters the regenerator for charcoal burning regeneration, and the regenerated catalyst returns to the reaction zone 2 through the inclined tube 15 of regeneration.
  • the methanol feedstock (methanol purity 95wt%) enters the fast fluidized bed reaction zone, contacts with SAPO-34 molecular sieve catalyst, and produces products including ethylene and propylene under effective conditions.
  • the catalyst in the fluidized bed reaction zone is deactivated to form standby Catalyst, the spent catalyst enters the regenerator to be regenerated to form a regenerated catalyst, the regenerated catalyst is returned to the fluidized bed reaction zone, the difference between the coke amount of the spent catalyst and the regenerated catalyst is 5wt%, the fluidized bed reaction
  • the gas phase and catalyst in the zone are quickly separated by separation equipment after the reaction is completed or after leaving the fluidized bed reaction zone.
  • the effective conditions are: the reaction temperature is 450°C, the reaction gauge pressure is 0.15 MPa, and the mass of the catalyst in the reaction zone with a carbon deposit of less than 3 wt% is 14 wt%, and the carbon deposit is 3 to less than 5 wt. % Catalyst accounts for 68% by weight, catalysts with a carbon deposit of 5-10% by weight account for 16% by weight, and the ratio of the volume of the gas phase to the volume of the catalyst is 10. Sampling analysis results show that the methanol conversion rate is 99.96%, and the ethylene + propylene carbon-based selectivity is 84.52%.
  • the methanol feedstock (methanol purity 95%) enters the dense-phase fluidized bed reaction zone, contacts with SAPO-34 molecular sieve catalyst, and produces products including ethylene and propylene under effective conditions.
  • the catalyst in the fluidized bed reaction zone is deactivated to form a
  • the spent catalyst enters the regenerator to be regenerated to form a regenerated catalyst.
  • the regenerated catalyst is returned to the fluidized bed reaction zone.
  • the difference between the coke deposits of the spent catalyst and the regenerated catalyst is 6wt%.
  • the gas phase and catalyst in the reaction zone are quickly separated by separation equipment after the reaction is completed or after leaving the fluidized bed reaction zone.
  • the effective conditions are: the reaction temperature is 550°C, the reaction gauge pressure is 1 MPa, and based on the mass of the molecular sieve on the catalyst, the mass of the catalyst with a carbon deposit of less than 3 wt% in the reaction zone is 18 wt%, and the carbon deposit is 3 to less than 5 wt%.
  • the catalyst accounted for 63% by weight, the catalyst with a carbon deposit of 5-10% by weight accounted for 17% by weight, and the ratio of the gas phase volume to the catalyst volume was 1.
  • Sampling analysis results show that the methanol conversion rate is 99.32%, and the ethylene + propylene carbon-based selectivity is 82.14%.
  • the methanol feedstock (methanol purity 95%) enters the turbulent fluidized bed reaction zone, contacts with SAPO-34 molecular sieve catalyst, and produces products including ethylene and propylene under effective conditions.
  • the catalyst in the fluidized bed reaction zone is deactivated to form a
  • the spent catalyst enters the regenerator to be regenerated to form a regenerated catalyst.
  • the regenerated catalyst is returned to the fluidized bed reaction zone.
  • the difference in coke deposits between the spent catalyst and the regenerated catalyst is 3wt%.
  • the gas phase and catalyst in the reaction zone are quickly separated by separation equipment after the reaction is completed or after leaving the fluidized bed reaction zone.
  • the effective conditions are: the reaction temperature is 400° C., the reaction gauge pressure is 0.05 MPa, based on the mass of the molecular sieve on the catalyst, the mass of the catalyst with a coke deposit less than 3wt% in the reaction zone is 15wt%, and the coke deposit is 3 to less than 5 wt% % Catalyst accounts for 66% by weight, catalysts with a carbon deposit of 5-10% by weight account for 16% by weight, and the ratio of gas phase volume to catalyst volume is 3. Sampling analysis results show that the methanol conversion rate is 99.09%, and the ethylene + propylene carbon-based selectivity is 83.99%.
  • the methanol feedstock (methanol purity 99%) enters the fast fluidized bed reaction zone, contacts with SAPO-34 molecular sieve catalyst, and produces products including ethylene and propylene under effective conditions.
  • the catalyst in the fluidized bed reaction zone is deactivated to form a standby Catalyst, the spent catalyst enters the regenerator to be regenerated to form a regenerated catalyst, the regenerated catalyst is returned to the fluidized bed reaction zone, the difference between the coke amount of the spent catalyst and the regenerated catalyst is 5wt%, the fluidized bed reaction
  • the gas phase and catalyst in the zone are quickly separated by separation equipment after the reaction is completed or after leaving the fluidized bed reaction zone.
  • the effective conditions are: the reaction temperature is 480°C, the reaction gauge pressure is 0.2MPa, and the mass of the catalyst in the reaction zone where the carbon deposit is less than 3wt% is 5wt%, and the carbon deposit is 3 to less than 5 wt% based on the mass of the molecular sieve on the catalyst.
  • % Catalyst accounts for 50% by weight, catalysts with a carbon deposit of 5-10% by weight account for 42% by weight, and the ratio of gas phase volume to catalyst volume is 6. Sampling analysis results show that the methanol conversion rate is 99.90%, and the ethylene + propylene carbon-based selectivity is 84.22%.
  • the methanol feedstock enters the fast fluidized bed reaction zone, contacts with SAPO-34 molecular sieve catalyst, and produces products including ethylene and propylene under effective conditions.
  • the catalyst in the fluidized bed reaction zone is deactivated to form a standby Catalyst, the spent catalyst enters the regenerator to be regenerated to form a regenerated catalyst, the regenerated catalyst is returned to the fluidized bed reaction zone, the difference between the coke amount of the spent catalyst and the regenerated catalyst is 6wt%, the fluidized bed reaction
  • the gas phase and catalyst in the zone are quickly separated by separation equipment after the reaction is completed or after leaving the fluidized bed reaction zone.
  • the effective conditions are: a reaction temperature of 480°C, a reaction gauge pressure of 0.15 MPa, based on the mass of the molecular sieve on the catalyst, the mass of the catalyst with a coke deposit less than 3wt% in the reaction zone is 10wt%, and the coke deposit is 3 to less than 5 wt% % Catalyst accounts for 45% by weight, catalysts with a carbon deposit of 5-10% by weight account for 40% by weight, and the ratio of gas phase volume to catalyst volume is 12. Sampling analysis results show that the methanol conversion rate is 99.96%, and the ethylene + propylene carbon-based selectivity is 84.78%.
  • the mass of the catalyst with the carbon deposit less than 3wt% in the reaction zone is 30wt%, and the ratio of the gas phase volume to the catalyst volume is 0.5.
  • Sampling analysis results show that the methanol conversion rate is 99.99%, and the ethylene + propylene carbon-based selectivity is 80.32%.
  • the mass of the catalyst with the carbon deposit less than 3 wt% in the reaction zone is 30 wt%, and the ratio of the gas phase volume to the catalyst volume is 20.
  • Sampling analysis results show that the methanol conversion rate is 99.67%, and the ethylene + propylene carbon-based selectivity is 79.61%.
  • the mass of the catalyst with a carbon deposit less than 3 wt% in the reaction zone is 10 wt%, and the ratio of the gas phase volume to the catalyst volume is 20.
  • Sampling analysis results show that the methanol conversion rate is 99.07%, and the ethylene + propylene carbon-based selectivity is 83.98%.
  • the mass of the catalyst with the carbon deposit less than 3 wt% in the reaction zone is 0.5 wt%, and the ratio of the gas phase volume to the catalyst volume is 12.
  • Sampling analysis results show that the methanol conversion rate is 99.01%, and the ethylene + propylene carbon-based selectivity is 83.76%.
  • the method of the present invention can achieve the purpose of improving the yield of low-carbon olefins, and can be used in the industrial production of low-carbon olefins.
  • the methanol feedstock (methanol purity 95wt%) enters the fast fluidized bed reaction zone, contacts with SAPO-34 molecular sieve catalyst, and produces products including ethylene and propylene under effective conditions.
  • the catalyst in the fluidized bed reaction zone is deactivated to form standby The catalyst, the spent catalyst enters the regenerator to be regenerated to form a regenerated catalyst, the regenerated catalyst is returned to the fluidized bed reaction zone, the outlet of the regeneration pipeline is provided with a catalyst distributor, and the distributor is along the radial direction of the fluidized bed reactor Horizontally, the regenerated catalyst is evenly distributed on the radial plane of the reaction zone of the fluidized bed reactor.
  • the catalyst distributor is equipped with a conveying medium, which is steam; the fluidized bed layer is divided into the reaction zone of the fluidized bed reactor.
  • a conveying medium which is steam; the fluidized bed layer is divided into the reaction zone of the fluidized bed reactor.
  • the distribution equipment is located in the dense phase section of the fluidized bed at a distance of 1/2 the bed height of the distribution equipment (for the fast fluidized bed type, the height of the dense phase is the reaction zone Height); the maximum difference of carbon deposit on the catalyst particles within the range of 1/2 the bed height from the distribution device to the distribution device is 3.3%;
  • the effective reaction conditions are: the reaction temperature is 480°C, and the reaction gauge pressure is 0.15MPa;
  • the ratio of the regenerated catalyst to the spent catalyst in the reaction zone is controlled to be 0.1, the carbon deposit of the regenerated catalyst is 1.0% by weight, and the methanol conversion rate from the distribution equipment to the 1/2 bed height of the distribution equipment Is 85%.
  • Sampling analysis results show that the methanol conversion rate at the
  • the methanol feedstock (methanol purity 95wt%) enters the fast fluidized bed reaction zone, contacts with the SAPO-34 molecular sieve catalyst, and produces products including ethylene and propylene under effective conditions
  • a spent catalyst is formed.
  • the spent catalyst enters the regenerator to be regenerated to form a regenerated catalyst.
  • the regenerated catalyst is returned to the fluidized bed reaction zone.
  • a catalyst is provided at the outlet of the regeneration pipeline.
  • the distributor is arranged horizontally along the radial direction of the fluidized bed reactor to uniformly distribute the regenerated catalyst on the radial plane of the reaction zone of the fluidized bed reactor.
  • the catalyst distributor is provided with a conveying medium, which is steam
  • a conveying medium which is steam
  • the fluidized bed in the reaction zone of the fluidized bed reactor is divided into two sections, dense phase and dilute phase, and the distribution equipment is located in the dense phase section of the fluidized bed at a distance of 1/2 the bed height from the distribution equipment;
  • the maximum difference between the distribution equipment and the carbon deposits on the catalyst particles within the range of 1/2 the bed height from the distribution equipment is 2.5%;
  • the effective reaction conditions are: the reaction temperature is 480°C, and the reaction gauge pressure is 0.01 MPa;
  • the ratio of the controlled regenerated catalyst to the spent catalyst in the reaction zone is 0.3; the carbon deposit of the regenerated catalyst is 2% by weight; the methanol conversion rate from the distribution equipment to the 1/2 bed height of the distribution equipment is 82% .
  • Sampling analysis results show that the methanol conversion rate at the outlet of the reactor is 99.61%, and the ethylene + propylene carbon-based selectivity is 86.55%.
  • the methanol feedstock (methanol purity 95wt%) enters the fast fluidized bed reaction zone, contacts with SAPO-34 molecular sieve catalyst, and produces products including ethylene and propylene under effective conditions
  • a spent catalyst is formed.
  • the spent catalyst enters the regenerator to be regenerated to form a regenerated catalyst.
  • the regenerated catalyst is returned to the fluidized bed reaction zone.
  • a catalyst is provided at the outlet of the regeneration pipeline.
  • Distributor, the distributor is arranged horizontally along the radial direction of the fluidized bed reactor to uniformly distribute the regenerated catalyst on the radial plane of the reaction zone of the fluidized bed reactor.
  • the catalyst distributor is provided with a conveying medium, which is steam
  • a conveying medium which is steam
  • the fluidized bed in the reaction zone of the fluidized bed reactor is divided into two sections, dense phase and dilute phase, and the distribution equipment is located in the dense phase section of the fluidized bed at a distance of 1/2 the bed height from the distribution equipment;
  • the maximum difference between the distribution equipment and the carbon deposit on the catalyst particles within the range of 1/2 the bed height from the distribution equipment is 7%;
  • the effective reaction conditions are: the reaction temperature is 550° C., and the reaction gauge pressure is 1.0 MPa;
  • the ratio of the regenerated catalyst to the spent catalyst in the reaction zone is controlled to be 0.05; the amount of coke deposited on the regenerated catalyst is 0.01% by weight; the methanol conversion rate from the distribution equipment to the 1/2 bed height of the distribution equipment is 91% .
  • Sampling analysis results show that the methanol conversion rate at the outlet of the reactor is 99.99%, and the ethylene + propylene carbon-based selectivity
  • the methanol feedstock (methanol purity 95wt%) enters the fast fluidized bed reaction zone, contacts with the SAPO-34 molecular sieve catalyst, and produces products including ethylene and propylene under effective conditions
  • a spent catalyst is formed.
  • the spent catalyst enters the regenerator to be regenerated to form a regenerated catalyst.
  • the regenerated catalyst is returned to the fluidized bed reaction zone.
  • the outlet of the regeneration pipeline is provided with a catalyst.
  • the distributor is arranged horizontally along the radial direction of the fluidized bed reactor to uniformly distribute the regenerated catalyst on the radial plane of the reaction zone of the fluidized bed reactor.
  • the catalyst distributor is provided with a conveying medium, which is steam
  • a conveying medium which is steam
  • the fluidized bed in the reaction zone of the fluidized bed reactor is divided into two sections, dense phase and dilute phase, and the distribution equipment is located in the dense phase section of the fluidized bed at a distance of 1/2 the bed height from the distribution equipment;
  • the maximum difference between the distributing equipment and the carbon deposits on the catalyst particles within the range of 1/2 the bed height from the distributing equipment is 4.5%;
  • the effective reaction conditions are: the reaction temperature is 490°C, and the reaction gauge pressure is 0.17 MPa;
  • the ratio of the controlled regenerated catalyst to the spent catalyst in the reaction zone is 0.08; the carbon deposit of the regenerated catalyst is 0.5% by weight;
  • the methanol conversion rate from the distribution equipment to the 1/2 bed height of the distribution equipment is 88% .
  • Sampling analysis results show that the methanol conversion rate at the outlet of the reactor is 99.98%, and the ethylene + propylene carbon-based selectivity is 85.1
  • the methanol feedstock (methanol purity 95wt%) enters the fast fluidized bed reaction zone, contacts with the SAPO-34 molecular sieve catalyst, and produces products including ethylene and propylene under effective conditions
  • a spent catalyst is formed.
  • the spent catalyst enters the regenerator to be regenerated to form a regenerated catalyst.
  • the regenerated catalyst is returned to the fluidized bed reaction zone.
  • a catalyst is provided at the outlet of the regeneration pipeline.
  • the distributor is arranged horizontally along the radial direction of the fluidized bed reactor to uniformly distribute the regenerated catalyst on the radial plane of the reaction zone of the fluidized bed reactor.
  • the catalyst distributor is provided with a conveying medium, which is steam
  • a conveying medium which is steam
  • the fluidized bed in the reaction zone of the fluidized bed reactor is divided into two sections, dense phase and dilute phase, and the distribution equipment is located in the dense phase section of the fluidized bed at a distance of 1/2 the bed height from the distribution equipment;
  • the maximum difference between the distribution equipment and the carbon deposit on the catalyst particles within the range of 1/2 the bed height from the distribution equipment is 1.5%;
  • the effective reaction conditions are: the reaction temperature is 490°C, and the reaction gauge pressure is 0.15 MPa;
  • the ratio of the controlled regenerated catalyst to the spent catalyst in the reaction zone is 0.5; the carbon deposit of the regenerated catalyst is 5% by weight; the methanol conversion rate from the distribution equipment to the 1/2 bed height of the distribution equipment is 86%.
  • Sampling analysis results show that the methanol conversion rate at the outlet of the reactor is 99.92%, and the ethylene + propylene carbon-based selectivity is 85.99%.
  • the methanol feedstock (methanol purity 95wt%) enters the fast fluidized bed reaction zone, contacts with the SAPO-34 molecular sieve catalyst, and produces products including ethylene and propylene under effective conditions
  • a spent catalyst is formed.
  • the spent catalyst enters the regenerator to be regenerated to form a regenerated catalyst.
  • the regenerated catalyst is returned to the fluidized bed reaction zone.
  • a catalyst is provided at the outlet of the regeneration pipeline.
  • the distributor is arranged horizontally along the radial direction of the fluidized bed reactor to uniformly distribute the regenerated catalyst on the radial plane of the reaction zone of the fluidized bed reactor.
  • the catalyst distributor is provided with a conveying medium, which is steam
  • a conveying medium which is steam
  • the fluidized bed in the reaction zone of the fluidized bed reactor is divided into two sections, dense phase and dilute phase, and the distribution equipment is located in the dense phase section of the fluidized bed at a distance of 1/2 the bed height from the distribution equipment;
  • the maximum difference between the distribution equipment and the carbon deposits on the catalyst particles within the range of 1/2 the bed height from the distribution equipment is 3%;
  • the effective reaction conditions are: the reaction temperature is 400° C., and the reaction gauge pressure is 0.01 MPa;
  • the ratio of the regenerated catalyst to the spent catalyst in the reaction zone is controlled to be 0.05; the carbon deposit amount of the regenerated catalyst is 0.5% by weight; the methanol conversion rate from the distribution equipment to the 1/2 bed height of the distribution equipment is 81% .
  • Sampling analysis results show that the methanol conversion rate at the outlet of the reactor is 99.51%, and the ethylene + propylene carbon-based selectivity is 84.80%
  • the methanol feedstock (methanol purity 95wt%) enters the fast fluidized bed reaction zone, contacts with the SAPO-34 molecular sieve catalyst, and produces products including ethylene and propylene under effective conditions
  • a spent catalyst is formed.
  • the spent catalyst enters the regenerator to be regenerated to form a regenerated catalyst.
  • the regenerated catalyst is returned to the fluidized bed reaction zone.
  • a catalyst is provided at the outlet of the regeneration pipeline.
  • the distributor is arranged horizontally along the radial direction of the fluidized bed reactor to uniformly distribute the regenerated catalyst on the radial plane of the reaction zone of the fluidized bed reactor.
  • the catalyst distributor is provided with a conveying medium, which is steam
  • a conveying medium which is steam
  • the fluidized bed in the reaction zone of the fluidized bed reactor is divided into two sections, dense phase and dilute phase, and the distribution equipment is located in the dense phase section of the fluidized bed at a distance of 1/2 the bed height from the distribution equipment;
  • the maximum difference between the distribution equipment and the carbon deposits on the catalyst particles within the range of 1/2 the bed height from the distribution equipment is 8%;
  • the effective reaction conditions are: the reaction temperature is 550°C, and the reaction gauge pressure is 0.75 MPa;
  • the ratio of the regenerated catalyst to the spent catalyst in the reaction zone is controlled to be 0.15; the amount of coke deposited on the regenerated catalyst is 0.1% by weight; the methanol conversion rate from the distribution equipment to the 1/2 bed height of the distribution equipment is 93% .
  • Sampling analysis results show that the methanol conversion rate at the outlet of the reactor is 99.99%, and the ethylene + propylene carbon-based selectivity is 8
  • the maximum difference between the distribution equipment and the carbon deposits on the catalyst particles within 1/2 of the bed height range of the distribution equipment is 10%;
  • the methanol conversion rate at 1/2 the bed height of the distribution equipment is 71%.
  • Sampling analysis results show that the methanol conversion rate at the outlet of the reactor is 99.01%, and the ethylene + propylene carbon-based selectivity is 81.78%.
  • the method of the present invention can achieve the purpose of improving the yield of low-carbon olefins, and can be used in the industrial production of low-carbon olefins.

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Abstract

由含氧化合物生产低碳烯烃的方法，包括使包含含氧化合物的原料在流化床反应区与分子筛催化剂接触，在有效条件下生成含乙烯和/或丙烯产品的步骤；所述有效条件包括所述流化床反应区内，以催化剂中分子筛质量计，控制各种积炭量的催化剂质量占所述流化床反应区内全部催化剂质量的比例。

Description

含氧化合物生产低碳烯烃的方法 技术领域
本发明涉及一种含氧化合物生产低碳烯烃的方法。
背景技术
低碳烯烃，主要是乙烯和丙烯，是两种重要的基础化工原料，其需求量在不断增加。一般地，乙烯、丙烯是通过石油路线来生产，但由于石油资源有限的供应量及较高的价格，由石油资源生产乙烯、丙烯的成本不断增加。近年来，人们开始大力发展替代原料转化制乙烯、丙烯的技术。其中，一类重要的用于低碳烯烃生产的替代原料是含氧化合物，例如醇类(甲醇、乙醇)、醚类(二甲醚、甲乙醚)、酯类(碳酸二甲酯、甲酸甲酯)等，这些含氧化合物可以通过煤、天然气、生物质等能源转化而来。某些含氧化合物已经可以达到较大规模的生产，如甲醇，可以由煤或天然气制得，工艺十分成熟，可以实现上百万吨级的生产规模。由于含氧化合物来源的广泛性，再加上转化生成低碳烯烃工艺的经济性，所以由含氧化合物转化制烯烃(OTO)的工艺，特别是由甲醇转化制烯烃(MTO)的工艺受到越来越多的重视。
文献US4499327对磷酸硅铝分子筛催化剂应用于甲醇转化制烯烃工艺进行了详细研究，认为SAPO-34是MTO工艺的首选催化剂。SAPO-34催化剂具有很高的低碳烯烃选择性，而且活性也较高，可使甲醇转化为低碳烯烃的反应时间达到小于10秒的程度，更甚至达到提升管的反应时间范围内。
文献US6166282公布了一种甲醇转化为低碳烯烃的技术和反应器，采用快速流化床反应器，气相在气速较低的密相反应区反应完成后，上升到内径急速变小的快分区后，采用特殊的气固分离设备初步分离出大部分的夹带催化剂。由于反应后产物气与催化剂快速分离，有效的防止了二次反应的发生。经模拟计算，与传统的鼓泡流化床反应器相比，该快速流化床反应器内径及催化剂所需藏量均大大减少。该方法中低碳烯烃碳基收率一般均在77％左右。
文献CN101328103A公开了一种甲醇或二甲醚转化为低碳烯烃的方法，包括甲醇或二甲醚的原料进入流化床反应器的反应区中，与包 括硅铝磷酸盐分子筛的催化剂接触；其中，反应压力以表压计为0.05～1MPa、反应区平均温度为450～550℃、反应区平均空塔气速为0.8～2.0米/秒的条件下，反应区的平均密度为20～300千克/立方米，所述反应区内的催化剂平均积炭量为1.5～4.5％重量。低碳烯烃碳基选择性最高达到81.51％重量。
然而，随着市场上对乙烯、丙烯需求量的不断增加，对低碳烯烃生产技术提出了更高的要求。
发明内容
本发明的发明人发现，现有技术的含氧化合物(尤其是甲醇)转化为低碳烯烃过程中，催化剂上含有一定量的积碳对于保证低碳烯烃的高选择性是必需的，在流化床反应器内，存在反应-再生之间的催化剂循环，必然反应区内存在多股催化剂的混合问题，而反应区内催化剂的积炭量是一个平均的概念，低积炭量催化剂与高积炭量催化剂的混合质量对于提高低碳烯烃选择性至关重要。
在此基础上，本发明的发明人进一步发现，原料在不同积炭量的催化剂上的反应性能差异巨大，尤其是在积炭量小于3wt％催化剂上反应时，低碳烯烃选择性大幅降低，但是要保证高的原料转化率，积炭量小于3wt％催化剂又是必不可少的。控制这部分较低积炭量催化剂的含量和分布，同时匹配控制反应区内催化剂浓度，是解决这一矛盾的有效方法，既能保证高的甲醇转化率，又能保证高的低碳烯烃选择性。此外，发明人还出人意料地发现，将反应区内混合催化剂的积炭量差值控制好对于实现所需的技术效果是至关重要的。
本发明基于这些发现而完成。
具体而言，本发明例如提供了以下方面的实施方式：
1、一种由含氧化合物生产低碳烯烃的方法，包括使包含含氧化合物的原料在流化床反应区与分子筛催化剂接触，在有效条件下生成含乙烯和/或丙烯产品的步骤；
所述有效条件包括所述流化床反应区内，以催化剂中分子筛质量计，控制各种积炭量的催化剂质量占所述流化床反应区内全部催化剂质量的比例如下：
积炭量小于3重量％的催化剂质量占所述流化床反应区内全部催 化剂质量的比例为1～20重量％，优选为1～15重量％，1.5～10重量％，或2～5重量％；
积炭量为3至小于5重量％的催化剂占10～70重量％，优选为15～60重量％，20～50重量％，或30～45重量％；
积炭量为5-10重量％的催化剂占10～88重量％，优选为15～80重量％，20～70重量％，或30～60重量％。
2、根据实施方式1所述的方法，其中所述流化床反应区内，气相体积与所述流化床反应区内全部催化剂体积之比为1～15，优选5～12。
3、根据实施方式1或2所述的方法，其中所述原料经分布设备进入流化床反应器反应区，其中所述分布设备至距离所述分布设备1/2床层高度范围内的催化剂颗粒上的积碳量最大差值小于8重量％，优选小于5重量％、或小于3重量％。
4、根据实施方式3所述的方法，其中所述分布设备至距离所述分布设备1/2床层高度范围内的催化剂颗粒上的积碳量最大差值大于0.1重量％。
5、一种由含氧化合物生产低碳烯烃的方法，包括使含氧化合物原料在流化床反应区与分子筛催化剂接触，在有效条件下生成含乙烯和/或丙烯产品的步骤；
其中所述原料经分布设备进入流化床反应器反应区，其中所述分布设备至距离所述分布设备1/2床层高度范围内的催化剂颗粒上的积碳量最大差值小于8重量％，优选小于5重量％、或小于3重量％。
6、根据实施方式5所述的方法，其中所述分布设备至距离所述分布设备1/2床层高度范围内的催化剂颗粒上的积碳量最大差值大于0.1重量％。
7、根据实施方式5或6所述的方法，其中所述有效条件包括所述流化床反应区内，以催化剂中分子筛质量计，控制各种积炭量的催化剂质量占所述流化床反应区内全部催化剂质量的比例如下：
积炭量小于3重量％的催化剂质量占所述流化床反应区内全部催化剂质量的比例为1～20重量％，优选为1～15重量％，1.5～10重量％，或2～5重量％；
积炭量为3至小于5重量％的催化剂占10～70重量％，优选为15～60重量％，20～50重量％，或30～45重量％；
积炭量为5-10重量％的催化剂占10～88重量％，优选为15～80重量％，20～70重量％，或30～60重量％。
8、根据实施方式7所述的方法，其中所述流化床反应区内，气相体积与所述流化床反应区内全部催化剂体积之比为1～15，优选5～12。
9、根据前述实施方式任一所述的方法，其中所述含氧化合物原料包括甲醇。
10、根据前述实施方式任一所述的方法，其中所述分子筛为硅铝磷分子筛，优选SAPO-18、SAPO-34、SAPO-5或其组合。
11、根据前述实施方式任一所述的方法，其中所述流化床反应区为密相、湍动或快速流态化型式，优选快速流态化型式。
12、根据前述实施方式任一所述的方法，其中所述有效条件还包括：反应温度400～550℃，反应压力0～1MPaG。
13、根据前述实施方式任一所述的方法，其中待生催化剂与再生催化剂的积炭量之差不大于7重量％，优选不大于6重量％，更优选不大于5重量％。
14、根据前述实施方式任一所述的方法，其中所述积炭量小于3重量％的催化剂均匀分布于所述流化床反应区内。
15、根据前述实施方式任一所述的方法，其中所述流化床反应区内的气相和催化剂在反应完成后或离开所述流化床反应区后通过分离设备快速分离。
优选地，根据本发明的一个例示实施方式，其中所述反应区下部设置有再生管路出口，所述再生管路出口设置有催化剂分布器，所述催化剂分布器沿所述反应区径向基本水平布置。
16、根据实施方式3-15任一所述的方法，其中所述反应区内的催化剂失活后形成待生催化剂，所述待生催化剂通过待生管路进入再生器再生，形成再生催化剂，所述再生催化剂通过再生管路返回到所述流化床反应区；其中，所述反应区内控制再生催化剂与待生催化剂的比例为0.01～1，优选0.05～0.5，更优选0.07～0.3；再生催化剂的积炭量为0～5重量％，优选0.05～3重量％，更优选0.5～2重量％。
17、用于实施前述实施方式中任一项所述由含氧化合物生产低碳烯烃的方法的流化床反应器，包括：
反应区，用于接收甲醇原料，使其与催化剂接触，生成烯烃产品， 其中该过程使得催化剂至少部分失活，得到待生催化剂；
气固快速分离设备，用于使来自反应区的待生催化剂分离；
旋风分离器，用于接收气固快速分离设备分离出来的气相产品以及部分未被气固快速分离设备分离的待生催化剂，以进行再次分离；
汽提区，用于接收来自旋风分离器的料腿的待生催化剂；和
催化剂外循环斜管，用于使来自汽提区的至少部分经汽提的待生催化剂返回到反应区的底部。
附图说明
图1为本发明方法中所述流化床反应器的示意图。
图1中，
1为反应器原料进料管线；
2为反应器反应区；
3为气固快速分离区；
4为汽提区；
5为反应器外循环斜管；
6为原料分布设备；
8为反应器气固旋风分离器；
9为反应器分离区；
11为产品气出口管线；
14为待生斜管；
15为再生斜管；
技术效果
根据本发明方法，既能保证高的原料转化率，又能保证高的低碳烯烃选择性，低碳烯烃选择性可达到84％以上。
具体实施方式
下面对本发明的具体实施方式进行详细说明，但是需要指出的是，本发明的保护范围并不受这些具体实施方式的限制，而是由附录的权利要求书来确定。
本说明书提到的所有出版物、专利申请、专利和其它参考文献全都引于此供参考。除非另有定义，本说明书所用的所有技术和科学术 语都具有本领域技术人员常规理解的含义。在有冲突的情况下，以本说明书的定义为准。
当本说明书以词头“本领域技术人员公知”、“现有技术”或其类似用语来导出材料、物质、方法、步骤、装置或部件等时，该词头导出的对象涵盖本申请提出时本领域常规使用的那些，但也包括目前还不常用，却将变成本领域公认为适用于类似目的的那些。
在本说明书的上下文中，所谓“基本”、“大约”或类似表述指的是允许存在对于本领域技术人员而言可以接受或认为合理的偏差，比如偏差在±10％以内、±5％以内、±1％以内、±0.5％以内或者±0.1％以内。
在本说明书的上下文中，所使用的术语“反应区”，是针对流化床反应器而言的。理想地，流化床反应器包括反应区、进口区和分离区。“进口区”是在反应器中引入原料和催化剂的区段。“反应区”是在反应器中进料与催化剂在有效将进料的含氧化合物转化为轻烯烃产物的条件下接触的区段。“分离区”是在反应器中催化剂和反应器内的任何其它固体与产物分离的区段。典型地，反应区位于进口区和分离区之间。
在本说明书的上下文中，所述“气相”，包括加热汽化后的原料甲醇、稀释气(比如水蒸气)、以及气相的反应产物(比如轻烯烃、C4烃等)中的一种或几种。
在本说明书的上下文中，需要说明的是，反应区内除了气相所占的体积，剩余的就是催化剂体积。气相体积与催化剂体积之比表明了反应区内气固两相混合物中催化剂固体颗粒所占的体积分率。
在本说明书的上下文中，催化剂积炭量(或者平均积碳量)的计算方法为催化剂上的积炭质量除以所述催化剂质量。催化剂上的积炭质量测定方法如下：称量0.1～1克的带碳催化剂置于高温碳分析仪中燃烧，通过红外测定燃烧生成的二氧化碳质量，从而得到催化剂上的积炭质量。为了测定反应区内的催化剂积碳量，可以从反应区的各个位置，连续或周期性引出或者直接取出等量的小份催化剂。
在本说明书的上下文中，流化床反应区内的催化剂失活后形成待生催化剂(待生剂)，所述待生催化剂进入再生器再生，形成再生催化剂(再生剂)，所述再生催化剂返回到流化床反应区。
在没有明确指明的情况下，本说明书内所提到的所有百分数、份数、比率等都是以重量为基准的，而且压力是表压。
在本说明书的上下文中，本发明的任何两个或多个实施方式都可以任意组合，由此而形成的技术方案属于本说明书原始公开内容的一部分，同时也落入本发明的保护范围。
根据本发明的一个实施方式，涉及一种含氧化合物生产低碳烯烃的方法。所述方法包括含氧化合物原料在流化床反应区与分子筛催化剂接触，在有效条件下生成含乙烯、丙烯产品的步骤。
根据本发明的一个实施方式，所述含氧化合物原料包括含1-20个碳原子，优选1-10个碳原子，更优选1-4个碳原子的脂族醇。例如，甲醇、乙醇、正丙醇、异丙醇、甲基乙基醚、二甲醚、二乙醚、二异丙基醚、甲醛、碳酸二甲酯、二甲基酮、乙酸，和它们的混合物；优选甲醇、乙醇、二甲醚、二乙醚，和它们的混合物；更优选甲醇和二甲醚；最优选甲醇。
根据本发明的一个实施方式，所述有效条件包括：所述流化床反应区内，以催化剂中分子筛质量计，积炭量小于3重量％的催化剂质量占所述流化床反应区内全部催化剂质量的比例为1～20重量％，优选为1～15重量％，优选为1.5～10重量％，更优选为2～5重量％。通过调节再生剂循环量和再生剂的积炭量(即再生器烧炭程度)控制积炭量小于3重量％的催化剂占所述流化床反应区内全部催化剂的比例。
根据本发明的一个实施方式，所述积炭量小于3重量％的催化剂均匀分布于所述流化床反应区内。
根据本发明的一个实施方式，所述有效条件还包括：所述流化床反应区内，气相体积与全部催化剂体积之比为1～15，优选5～12。
根据本发明的一个实施方式，所述有效条件还包括：反应温度400～550℃，反应压力0～1MPa。
根据本发明的一个实施方式，所述分子筛为硅铝磷分子筛，优选SAPO-18、SAPO-34、SAPO-5或其组合，更优选为SAPO-34。SAPO分子筛或者SAPO分子筛催化剂的制备方法是为本领域所熟知的。
根据本发明的一个实施方式，所述流化床反应区为密相、湍动或快速流态化型式，优选快速流态化型式。
根据本发明的一个实施方式，待生催化剂与再生催化剂的积炭量之差不大于7重量％，优选不大于6重量％，更优选不大于5重量％。 其中，待生催化剂的积炭量可以通过反应直接控制，再生催化剂的积炭量可以通过再生器再生程度控制。
根据本发明的一个实施方式，所述流化床反应区内的气相和催化剂在反应完成后或离开所述流化床反应区后通过分离设备快速分离。所述分离设备优选旋风分离器。
本发明的一个例示性的实施方式1如图1所示。参见图1，本发明使用的流化床反应器为快速流化床、密相流化床或湍动流化床，包括甲醇原料的物流经进料管线1进入反应器，并经过原料分布设备6进入反应区2中，与分子筛催化剂接触，反应生成含有低碳烯烃的产品，其使得催化剂至少部分失活，形成待生催化剂。携带待生催化剂经过气固快速分离区3进入反应器分离区9，其中，气固快速分离设备3分离出来的大部分催化剂进入汽提区4，而气相产品以及部分未被气固快速分离设备分离的待生催化剂经入旋风分离器8分离进行再次分离。旋风分离器8分离出来的待生催化剂经过旋风分离器8的料腿返回到汽提区4，而分离出来的气相产品经出口管线11进入后续的分离工段。被气固快速分离区3和旋风分离器8分离出的待生催化剂合并，经过汽提区4汽提后分为两部分，一部分通过催化剂外循环斜管5返回到反应区2的底部；另外一部分经过待生斜管14进入再生器中烧炭再生，再生完成的催化剂通过再生斜管15返回反应区2。
实施例
以下将通过实施例和比较例对本发明进行进一步的详细描述，但本发明不限于以下实施例。
【实施例I-1】
使用快速流化床。甲醇原料(甲醇纯度95wt％)进入快速流化床反应区，与SAPO-34分子筛催化剂接触，在有效条件下生成包括乙烯、丙烯的产品，流化床反应区内的催化剂失活后形成待生催化剂，所述待生催化剂进入再生器再生，形成再生催化剂，所述再生催化剂返回到流化床反应区，所述待生催化剂与再生催化剂的积炭量之差为5wt％，流化床反应区内的气相、催化剂在反应完成后或离开所述流化床反应区后通过分离设备快速分离。所述有效条件为：反应温度450℃，反应表压为0.15MPa，以催化剂上分子筛质量计，反应区内积炭量 小于3wt％的催化剂质量为14wt％，积炭量为3至小于5重量％的催化剂占68重量％，积炭量为5-10重量％的催化剂占16重量％，气相体积与催化剂体积之比为10。取样分析结果表明，甲醇转化率为99.96％，乙烯+丙烯碳基选择性为84.52％。
【实施例I-2】
使用密相流化床。甲醇原料(甲醇纯度95％)进入密相流化床反应区，与SAPO-34分子筛催化剂接触，在有效条件下生成包括乙烯、丙烯的产品，流化床反应区内的催化剂失活后形成待生催化剂，所述待生催化剂进入再生器再生，形成再生催化剂，所述再生催化剂返回到流化床反应区，所述待生催化剂与再生催化剂的积炭量之差为6wt％，流化床反应区内的气相、催化剂在反应完成后或离开所述流化床反应区后通过分离设备快速分离。所述有效条件为：反应温度550℃，反应表压为1MPa，以催化剂上分子筛质量计，反应区内积炭量小于3wt％的催化剂质量为18wt％，积炭量为3至小于5重量％的催化剂占63重量％，积炭量为5-10重量％的催化剂占17重量％，气相体积与催化剂体积之比为1。取样分析结果表明，甲醇转化率为99.32％，乙烯+丙烯碳基选择性为82.14％。
【实施例I-3】
使用湍动流化床。甲醇原料(甲醇纯度95％)进入湍动流化床反应区，与SAPO-34分子筛催化剂接触，在有效条件下生成包括乙烯、丙烯的产品，流化床反应区内的催化剂失活后形成待生催化剂，所述待生催化剂进入再生器再生，形成再生催化剂，所述再生催化剂返回到流化床反应区，所述待生催化剂与再生催化剂的积炭量之差为3wt％，流化床反应区内的气相、催化剂在反应完成后或离开所述流化床反应区后通过分离设备快速分离。所述有效条件为：反应温度400℃，反应表压为0.05MPa，以催化剂上分子筛质量计，反应区内积炭量小于3wt％的催化剂质量为15wt％，积炭量为3至小于5重量％的催化剂占66重量％，积炭量为5-10重量％的催化剂占16重量％，气相体积与催化剂体积之比为3。取样分析结果表明，甲醇转化率为99.09％，乙烯+丙烯碳基选择性为83.99％。
【实施例I-4】
使用快速流化床。甲醇原料(甲醇纯度99％)进入快速流化床反应区，与SAPO-34分子筛催化剂接触，在有效条件下生成包括乙烯、丙烯的产品，流化床反应区内的催化剂失活后形成待生催化剂，所述待生催化剂进入再生器再生，形成再生催化剂，所述再生催化剂返回到流化床反应区，所述待生催化剂与再生催化剂的积炭量之差为5wt％，流化床反应区内的气相、催化剂在反应完成后或离开所述流化床反应区后通过分离设备快速分离。所述有效条件为：反应温度480℃，反应表压为0.2MPa，以催化剂上分子筛质量计，反应区内积炭量小于3wt％的催化剂质量为5wt％，积炭量为3至小于5重量％的催化剂占50重量％，积炭量为5-10重量％的催化剂占42重量％，气相体积与催化剂体积之比为6。取样分析结果表明，甲醇转化率为99.90％，乙烯+丙烯碳基选择性为84.22％。
【实施例I-5】
甲醇原料(甲醇纯度99％)进入快速流化床反应区，与SAPO-34分子筛催化剂接触，在有效条件下生成包括乙烯、丙烯的产品，流化床反应区内的催化剂失活后形成待生催化剂，所述待生催化剂进入再生器再生，形成再生催化剂，所述再生催化剂返回到流化床反应区，所述待生催化剂与再生催化剂的积炭量之差为6wt％，流化床反应区内的气相、催化剂在反应完成后或离开所述流化床反应区后通过分离设备快速分离。所述有效条件为：反应温度480℃，反应表压为0.15MPa，以催化剂上分子筛质量计，反应区内积炭量小于3wt％的催化剂质量为10wt％，积炭量为3至小于5重量％的催化剂占45重量％，积炭量为5-10重量％的催化剂占40重量％，气相体积与催化剂体积之比为12。取样分析结果表明，甲醇转化率为99.96％，乙烯+丙烯碳基选择性为84.78％。
【比较例I-1】
按照【实施例I-5】所述的条件和步骤，只是反应区内积炭量小于3wt％的催化剂质量为30wt％，气相体积与催化剂体积之比为0.5。取样分析结果表明，甲醇转化率为99.99％，乙烯+丙烯碳基选择性为80.32％。
【比较例I-2】
按照【实施例I-5】所述的条件和步骤，只是反应区内积炭量小于3wt％的催化剂质量为30wt％，气相体积与催化剂体积之比为20。取样分析结果表明，甲醇转化率为99.67％，乙烯+丙烯碳基选择性为79.61％。
【比较例I-3】
按照【实施例I-5】所述的条件和步骤，只是反应区内积炭量小于3wt％的催化剂质量为10wt％，气相体积与催化剂体积之比为20。取样分析结果表明，甲醇转化率为99.07％，乙烯+丙烯碳基选择性为83.98％。
【比较例I-4】
按照【实施例I-5】所述的条件和步骤，只是反应区内积炭量小于3wt％的催化剂质量为0.5wt％，气相体积与催化剂体积之比为12。取样分析结果表明，甲醇转化率为99.01％，乙烯+丙烯碳基选择性为83.76％。
【比较例I-5】
按照【实施例I-5】所述的条件和步骤，只是待生催化剂与再生催化剂的积炭量之差为9wt％。取样分析结果表明，甲醇转化率为98.97％，乙烯+丙烯碳基选择性为83.55％。
显然，采用本发明的方法，可以达到提高低碳烯烃收率的目的，可用于低碳烯烃的工业生产中。
【实施例II-1】
使用快速流化床。甲醇原料(甲醇纯度95wt％)进入快速流化床反应区，与SAPO-34分子筛催化剂接触，在有效条件下生成包括乙烯、丙烯的产品，流化床反应区内的催化剂失活后形成待生催化剂，所述待生催化剂进入再生器再生，形成再生催化剂，所述再生催化剂返回到流化床反应区，再生管路出口设有催化剂分布器，所述分布器沿流化床反应器径向水平布置，将再生催化剂均匀分布于流化床反应器反应区的径向平面上，催化剂分布器设有输送介质，输送介质为水蒸气；流化床反应器反应区内流态化床层分为密相和稀相两段，分布设备至距离所述分布设备1/2床层高度区域位于流态化床层的密相段(对于快速流化床型式，密相段高度即是反应区高度)；分布设备至距离所述分布设备1/2床层高度范围内的催化剂颗粒上的积碳量最大 差值为3.3％；所述有效反应条件为：反应温度480℃，反应表压为0.15MPa；所述反应区内控制再生催化剂与待生催化剂的比例为0.1，再生催化剂的积炭量为1.0重量％，分布设备至距离所述分布设备1/2床层高度处的甲醇转化率为85％。取样分析结果表明，反应器出口甲醇转化率为99.95％，乙烯+丙烯碳基选择性为84.36％。
【实施例II-2】
按照【实施例II-1】所述的条件和步骤，甲醇原料(甲醇纯度95wt％)进入快速流化床反应区，与SAPO-34分子筛催化剂接触，在有效条件下生成包括乙烯、丙烯的产品，流化床反应区内的催化剂失活后形成待生催化剂，所述待生催化剂进入再生器再生，形成再生催化剂，所述再生催化剂返回到流化床反应区，再生管路出口设有催化剂分布器，所述分布器沿流化床反应器径向水平布置，将再生催化剂均匀分布于流化床反应器反应区的径向平面上，催化剂分布器设有输送介质，输送介质为水蒸气；流化床反应器反应区内流态化床层分为密相和稀相两段，分布设备至距离所述分布设备1/2床层高度区域位于流态化床层的密相段；分布设备至距离所述分布设备1/2床层高度范围内的催化剂颗粒上的积碳量最大差值为2.5％；所述有效反应条件为：反应温度480℃，反应表压为0.01MPa；所述反应区内控制再生催化剂与待生催化剂的比例为0.3；再生催化剂的积炭量为2重量％；分布设备至距离所述分布设备1/2床层高度处的甲醇转化率为82％。取样分析结果表明，反应器出口甲醇转化率为99.61％，乙烯+丙烯碳基选择性为86.55％。
【实施例II-3】
按照【实施例II-1】所述的条件和步骤，甲醇原料(甲醇纯度95wt％)进入快速流化床反应区，与SAPO-34分子筛催化剂接触，在有效条件下生成包括乙烯、丙烯的产品，流化床反应区内的催化剂失活后形成待生催化剂，所述待生催化剂进入再生器再生，形成再生催化剂，所述再生催化剂返回到流化床反应区，再生管路出口设有催化剂分布器，所述分布器沿流化床反应器径向水平布置，将再生催化剂均匀分布于流化床反应器反应区的径向平面上，催化剂分布器设有输送介质，输送介质为水蒸气；流化床反应器反应区内流态化床层分为密相和稀相两段，分布设备至距离所述分布设备1/2床层高度区域位于流态化床层的密相段；分布设备至距离所述分布设备1/2床层高度范围内的催化剂 颗粒上的积碳量最大差值为7％；所述有效反应条件为：反应温度550℃，反应表压为1.0MPa；所述反应区内控制再生催化剂与待生催化剂的比例为0.05；再生催化剂的积炭量为0.01重量％；分布设备至距离所述分布设备1/2床层高度处的甲醇转化率为91％。取样分析结果表明，反应器出口甲醇转化率为99.99％，乙烯+丙烯碳基选择性为83.59％。
【实施例II-4】
按照【实施例II-1】所述的条件和步骤，甲醇原料(甲醇纯度95wt％)进入快速流化床反应区，与SAPO-34分子筛催化剂接触，在有效条件下生成包括乙烯、丙烯的产品，流化床反应区内的催化剂失活后形成待生催化剂，所述待生催化剂进入再生器再生，形成再生催化剂，所述再生催化剂返回到流化床反应区，再生管路出口设有催化剂分布器，所述分布器沿流化床反应器径向水平布置，将再生催化剂均匀分布于流化床反应器反应区的径向平面上，催化剂分布器设有输送介质，输送介质为水蒸气；流化床反应器反应区内流态化床层分为密相和稀相两段，分布设备至距离所述分布设备1/2床层高度区域位于流态化床层的密相段；分布设备至距离所述分布设备1/2床层高度范围内的催化剂颗粒上的积碳量最大差值为4.5％；所述有效反应条件为：反应温度490℃，反应表压为0.17MPa；所述反应区内控制再生催化剂与待生催化剂的比例为0.08；再生催化剂的积炭量为0.5重量％；分布设备至距离所述分布设备1/2床层高度处的甲醇转化率为88％。取样分析结果表明，反应器出口甲醇转化率为99.98％，乙烯+丙烯碳基选择性为85.19％。
【实施例II-5】
按照【实施例II-1】所述的条件和步骤，甲醇原料(甲醇纯度95wt％)进入快速流化床反应区，与SAPO-34分子筛催化剂接触，在有效条件下生成包括乙烯、丙烯的产品，流化床反应区内的催化剂失活后形成待生催化剂，所述待生催化剂进入再生器再生，形成再生催化剂，所述再生催化剂返回到流化床反应区，再生管路出口设有催化剂分布器，所述分布器沿流化床反应器径向水平布置，将再生催化剂均匀分布于流化床反应器反应区的径向平面上，催化剂分布器设有输送介质，输送介质为水蒸气；流化床反应器反应区内流态化床层分为 密相和稀相两段，分布设备至距离所述分布设备1/2床层高度区域位于流态化床层的密相段；分布设备至距离所述分布设备1/2床层高度范围内的催化剂颗粒上的积碳量最大差值1.5％；所述有效反应条件为：反应温度490℃，反应表压为0.15MPa；所述反应区内控制再生催化剂与待生催化剂的比例为0.5；再生催化剂的积炭量为5重量％；分布设备至距离所述分布设备1/2床层高度处的甲醇转化率为86％。取样分析结果表明，反应器出口甲醇转化率为99.92％，乙烯+丙烯碳基选择性为85.99％。
【实施例II-6】
按照【实施例II-1】所述的条件和步骤，甲醇原料(甲醇纯度95wt％)进入快速流化床反应区，与SAPO-34分子筛催化剂接触，在有效条件下生成包括乙烯、丙烯的产品，流化床反应区内的催化剂失活后形成待生催化剂，所述待生催化剂进入再生器再生，形成再生催化剂，所述再生催化剂返回到流化床反应区，再生管路出口设有催化剂分布器，所述分布器沿流化床反应器径向水平布置，将再生催化剂均匀分布于流化床反应器反应区的径向平面上，催化剂分布器设有输送介质，输送介质为水蒸气；流化床反应器反应区内流态化床层分为密相和稀相两段，分布设备至距离所述分布设备1/2床层高度区域位于流态化床层的密相段；分布设备至距离所述分布设备1/2床层高度范围内的催化剂颗粒上的积碳量最大差值为3％；所述有效反应条件为：反应温度400℃，反应表压为0.01MPa；所述反应区内控制再生催化剂与待生催化剂的比例为0.05；再生催化剂的积炭量为0.5重量％；分布设备至距离所述分布设备1/2床层高度处的甲醇转化率为81％。取样分析结果表明，反应器出口甲醇转化率为99.51％，乙烯+丙烯碳基选择性为84.80％。
【实施例II-7】
按照【实施例II-1】所述的条件和步骤，甲醇原料(甲醇纯度95wt％)进入快速流化床反应区，与SAPO-34分子筛催化剂接触，在有效条件下生成包括乙烯、丙烯的产品，流化床反应区内的催化剂失活后形成待生催化剂，所述待生催化剂进入再生器再生，形成再生催化剂，所述再生催化剂返回到流化床反应区，再生管路出口设有催化剂分布器，所述分布器沿流化床反应器径向水平布置，将再生催化剂均匀分布于流化床反应器反应区的径向平面上，催化剂分布器设有输送介质，输 送介质为水蒸气；流化床反应器反应区内流态化床层分为密相和稀相两段，分布设备至距离所述分布设备1/2床层高度区域位于流态化床层的密相段；分布设备至距离所述分布设备1/2床层高度范围内的催化剂颗粒上的积碳量最大差值为8％；所述有效反应条件为：反应温度550℃，反应表压为0.75MPa；所述反应区内控制再生催化剂与待生催化剂的比例为0.15；再生催化剂的积炭量为0.1重量％；分布设备至距离所述分布设备1/2床层高度处的甲醇转化率为93％。取样分析结果表明，反应器出口甲醇转化率为99.99％，乙烯+丙烯碳基选择性为84.33％。
【比较例II-1】
按照【实施例II-1】所述的条件和步骤，分布设备至距离所述分布设备1/2床层高度范围内的催化剂颗粒上的积碳量最大差值10％；分布设备至距离所述分布设备1/2床层高度处的甲醇转化率为71％。取样分析结果表明，反应器出口甲醇转化率为99.01％，乙烯+丙烯碳基选择性为81.78％。
显然，采用本发明的方法，可以达到提高低碳烯烃收率的目的，可用于低碳烯烃的工业生产中。

Claims (16)

  1. 一种由含氧化合物生产低碳烯烃的方法，包括使包含含氧化合物的原料在流化床反应区与分子筛催化剂接触，在有效条件下生成含乙烯和/或丙烯产品的步骤；
    所述有效条件包括所述流化床反应区内，以催化剂中分子筛质量计，控制各种积炭量的催化剂质量占所述流化床反应区内全部催化剂质量的比例如下：
    积炭量小于3重量％的催化剂质量占所述流化床反应区内全部催化剂质量的比例为1～20重量％，优选为1～15重量％，1.5～10重量％，或2～5重量％；
    积炭量为3至小于5重量％的催化剂占10～70重量％，优选为15～60重量％，20～50重量％，或30～45重量％；
    积炭量为5-10重量％的催化剂占10～88重量％，优选为15～80重量％，20～70重量％，或30～60重量％。
  2. 根据权利要求1所述的方法，其特征在于，所述流化床反应区内，气相体积与所述流化床反应区内全部催化剂体积之比为1～15，优选5～12。
  3. 根据权利要求1或2所述的方法，其中所述原料经分布设备进入流化床反应器反应区，其特征在于，所述分布设备至距离所述分布设备1/2床层高度范围内的催化剂颗粒上的积碳量最大差值小于8重量％，优选小于5重量％、或小于3重量％。
  4. 根据权利要求3所述的方法，其特征在于，所述分布设备至距离所述分布设备1/2床层高度范围内的催化剂颗粒上的积碳量最大差值大于0.1重量％。
  5. 一种由含氧化合物生产低碳烯烃的方法，包括使含氧化合物原料在流化床反应区与分子筛催化剂接触，在有效条件下生成含乙烯和/或丙烯产品的步骤；
    其中所述原料经分布设备进入流化床反应器反应区，其特征在于，所述分布设备至距离所述分布设备1/2床层高度范围内的催化剂颗粒上的积碳量最大差值小于8重量％，优选小于5重量％、或小于3重量％。
  6. 根据权利要求5所述的方法，其特征在于，所述分布设备至距离所述分布设备1/2床层高度范围内的催化剂颗粒上的积碳量最大差值大于0.1重量％。
  7. 根据前述权利要求任一所述的方法，其特征在于，所述含氧化合物原料包括甲醇。
  8. 根据前述权利要求任一所述的方法，其特征在于，所述分子筛为硅铝磷分子筛，优选SAPO-18、SAPO-34、SAPO-5或其组合。
  9. 根据前述权利要求任一所述的方法，其特征在于，所述流化床反应区为密相、湍动或快速流态化型式，优选快速流态化型式。
  10. 根据前述权利要求任一所述的方法，其特征在于，所述有效条件还包括：反应温度400～550℃，反应压力0～1MPaG。
  11. 根据前述权利要求任一所述的方法，其特征在于，待生催化剂与再生催化剂的积炭量之差不大于7重量％，优选不大于6重量％，更优选不大于5重量％。
  12. 根据前述权利要求任一所述的方法，其特征在于，所述积炭量小于3重量％的催化剂均匀分布于所述流化床反应区内。
  13. 根据前述权利要求任一所述的方法，其特征在于，所述流化床反应区内的气相和催化剂在反应完成后或离开所述流化床反应区后通过分离设备快速分离。
  14. 根据权利要求3-13任一所述的方法，其特征在于，所述反应区内的催化剂失活后形成待生催化剂，所述待生催化剂通过待生管路进入再生器再生，形成再生催化剂，所述再生催化剂通过再生管路返回到所述流化床反应区；其中，所述反应区内控制再生催化剂与待生催化剂的比例为0.01～1，优选0.05～0.5，更优选0.07～0.3；再生催化剂的积炭量为0～5重量％，优选0.05～3重量％，更优选0.5～2重量％。
  15. 用于实施前述权利要求中任一项所述由含氧化合物生产低碳烯烃的方法的流化床反应器，包括：
    反应区，用于接收甲醇原料，使其与催化剂接触，生成烯烃产品，其中该过程使得催化剂至少部分失活，得到待生催化剂；
    气固快速分离设备，用于使来自反应区的待生催化剂分离；
    旋风分离器，用于接收气固快速分离设备分离出来的气相产品以及部分未被气固快速分离设备分离的待生催化剂，以进行再次分离；
    汽提区，用于接收来自旋风分离器的料腿的待生催化剂；和
    催化剂外循环斜管，用于使来自汽提区的至少部分经汽提的待生催化剂返回到反应区的底部。
  16. 根据权利要求15所述的流化床反应器，其还包括分离区，位于所述反应区和所述气固快速分离设备之间，用于使所述待生催化剂中的至少部分发生沉降，以与气相初步分离。
PCT/CN2021/080114 2020-03-13 2021-03-11 含氧化合物生产低碳烯烃的方法 WO2021180150A1 (zh)

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BR112022018248A BR112022018248A2 (pt) 2020-03-13 2021-03-11 Método para produção de olefina leve a partir de composto contendo oxigênio
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CN101293803A (zh) * 2008-04-11 2008-10-29 中国石油化工股份有限公司 含氧化合物转化为低碳烯烃的方法
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