WO2015081489A1 - 一种含氧化合物制低碳烯烃的方法 - Google Patents

一种含氧化合物制低碳烯烃的方法 Download PDF

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Publication number
WO2015081489A1
WO2015081489A1 PCT/CN2013/088398 CN2013088398W WO2015081489A1 WO 2015081489 A1 WO2015081489 A1 WO 2015081489A1 CN 2013088398 W CN2013088398 W CN 2013088398W WO 2015081489 A1 WO2015081489 A1 WO 2015081489A1
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Prior art keywords
zone
catalyst
carbon
fluidized bed
phase fluidized
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PCT/CN2013/088398
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English (en)
French (fr)
Inventor
刘中民
叶茂
张涛
何长青
王贤高
赵银峰
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中国科学院大连化学物理研究所
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Priority to EP13898489.3A priority Critical patent/EP3078651B1/en
Priority to AU2013407180A priority patent/AU2013407180B2/en
Priority to PCT/CN2013/088398 priority patent/WO2015081489A1/zh
Priority to BR112016012633-5A priority patent/BR112016012633B1/pt
Priority to MYPI2016702007A priority patent/MY171803A/en
Priority to SG11201604429VA priority patent/SG11201604429VA/en
Priority to US15/101,297 priority patent/US9725375B2/en
Priority to KR1020167017634A priority patent/KR101847474B1/ko
Priority to RU2016126180A priority patent/RU2632905C1/ru
Priority to DK13898489.3T priority patent/DK3078651T3/en
Priority to JP2016535725A priority patent/JP6189544B2/ja
Publication of WO2015081489A1 publication Critical patent/WO2015081489A1/zh

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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
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J29/00—Catalysts comprising molecular sieves
    • B01J29/90—Regeneration or reactivation
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J38/00—Regeneration or reactivation of catalysts, in general
    • B01J38/04—Gas or vapour treating; Treating by using liquids vaporisable upon contacting spent catalyst
    • B01J38/06—Gas or vapour treating; Treating by using liquids vaporisable upon contacting spent catalyst using steam
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J38/00—Regeneration or reactivation of catalysts, in general
    • B01J38/04—Gas or vapour treating; Treating by using liquids vaporisable upon contacting spent catalyst
    • B01J38/12—Treating with free oxygen-containing gas
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C11/00—Aliphatic unsaturated hydrocarbons
    • C07C11/02—Alkenes
    • C07C11/04—Ethylene
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C11/00—Aliphatic unsaturated hydrocarbons
    • C07C11/02—Alkenes
    • C07C11/06—Propene
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C11/00—Aliphatic unsaturated hydrocarbons
    • C07C11/02—Alkenes
    • C07C11/08—Alkenes with four carbon atoms
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C5/00—Preparation of hydrocarbons from hydrocarbons containing the same number of carbon atoms
    • C07C5/32—Preparation of hydrocarbons from hydrocarbons containing the same number of carbon atoms by dehydrogenation with formation of free hydrogen
    • C07C5/327—Formation of non-aromatic carbon-to-carbon double bonds only
    • C07C5/333—Catalytic processes
    • 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
    • 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 present invention relates to a process for making low carbon olefins having improved low carbon olefin yield. Background technique
  • Low-carbon olefins namely ethylene and propylene
  • ethylene and propylene are two important basic chemical materials, and their demand is increasing.
  • ethylene and propylene are produced through petroleum routes, but the cost of producing ethylene and propylene from petroleum resources is increasing due to the limited supply of petroleum resources and higher prices.
  • people have begun to vigorously develop technologies for converting raw materials into ethylene and propylene.
  • MTO methanol conversion to olefins
  • the process of methanol conversion to olefins (MTO) has received increasing attention and has achieved a production scale of millions of tons.
  • MTO methanol conversion to olefins
  • the CMAI analysis said that by 2016, ethylene demand will grow at an average annual rate of 4.3%, propylene. Demand will grow at an average annual rate of 4.4%. Due to the rapid growth of China's economy, the annual growth rate of ethylene and propylene demand in China exceeds the world average.
  • SAPO-34 molecular sieve catalyst showed excellent catalytic performance when used in MTO reaction, with high low-carbon olefin selectivity and high activity. However, the catalyst loses its activity due to carbon deposition after a period of use.
  • the SAPO-34 molecular sieve catalyst has a significant induction period during use. During the induction period, the selectivity of olefins is lower, and the selectivity of hydrazines is higher. As the reaction time increases, the selectivity of low olefins gradually increases. After the induction period, the catalyst maintains high selectivity and high activity for a certain period of time, and the activity of the catalyst rapidly decreases as time continues to increase.
  • U.S. Patent 6,166,282 discloses a technique and a reactor for the conversion of methanol to lower olefins using a fast fluidized bed reactor. After the reaction of the gas phase in the dense phase reaction zone where the gas velocity is low, the gas phase rises to a fast zone where the inner diameter rapidly decreases. Afterwards, most of the entrained catalyst was separated by a special gas-solid separation device. Since the product gas after the reaction is rapidly separated from the catalyst, the occurrence of the secondary reaction is effectively prevented. According to the simulation calculation, the inner diameter of the fast fluidized bed reactor and the required reserves of the catalyst are greatly reduced compared with the conventional bubbling fluidized bed reactor.
  • CN101402538B discloses a method for increasing the yield of low carbon olefins by using a second reaction zone in the upper portion of the first reaction zone in which methanol is converted to a lower olefin, and the diameter of the second reaction zone is larger than the first reaction zone.
  • the unreacted methanol, the produced dimethyl ether and the hydrocarbons of more than four carbons continue to react, thereby achieving the purpose of increasing the yield of the low-carbon olefin.
  • the method can increase the yield of low-carbon olefins to a certain extent, since the catalyst from the first reaction zone already has more carbon deposits, and the hydrocarbon cracking of carbon four or more requires higher catalyst activity, In the method, the hydrocarbon conversion efficiency of carbon more than four in the second reaction zone is still low, resulting in a low yield of low carbon olefins.
  • CN102276406 A discloses a process for producing propylene.
  • the technique provides three reaction zones, a first fast bed reaction zone for methanol conversion to olefins, a riser reaction zone and a second fast bed reaction zone for series conversion of ethylene, carbon tetra or higher hydrocarbons and unreacted methanol or two Methyl ether.
  • hydrocarbons such as carbon four or more have a shorter residence time in the riser reaction zone and the second fast bed reaction zone, and the conversion efficiency is lower, resulting in a lower propylene yield.
  • a fluidized bed reactor for internally arranging a riser reactor for increasing the yield of light olefins is disclosed.
  • the first raw material enters the fluidized bed reaction zone, contacts with the catalyst to form a product including a low-carbon olefin, and simultaneously forms a catalyst to be produced; a part of the catalyst to be produced enters the regenerator to be regenerated, forms a regenerated catalyst, and a part enters the outlet end and is located inside the reaction zone.
  • the riser is in contact with the second raw material to raise the catalyst to be reacted into the reaction zone; the regenerated catalyst is returned to the reaction zone of the fluidized bed reactor.
  • the reaction device disclosed in this patent has no stripping part, and the raw catalyst will carry some product gas into the regenerator, burning with oxygen, and reducing low-carbon olefins.
  • the methanol to olefins technology disclosed in CN102875296A provides three reaction zones of a fast bed, a down bed and a riser.
  • the catalyst circulates between the regenerator, the fast bed, the riser and the descending bed.
  • the flow is very complicated, the flow distribution and control are very difficult, and the activity of the catalyst changes greatly.
  • the selectivity of the low olefins is closely related to the amount of carbon deposited on the catalyst.
  • a certain amount of carbon is required on the SAPO-34 catalyst.
  • the main reactor used in the MTO process is a fluidized bed, and the fluidized bed is close to the full mixed-flow reactor.
  • the distribution of catalyst coke is wide, which is not conducive to improving the selectivity of low-carbon olefins.
  • the MTO process has a small ratio of solvent to alcohol and a low coke yield.
  • it is necessary to control the carbon content and carbon content uniformity of the catalyst to a certain level in the regeneration zone. into The purpose of controlling the carbon deposition amount and carbon content uniformity on the catalyst in the reaction zone is achieved. Therefore, controlling the carbon deposition amount and carbon content uniformity in the reaction zone at a certain level is a key technology in the MTO process.
  • Low-carbon olefins are also very sensitive to process parameters such as reaction temperature.
  • the temperature of the regenerated catalyst is generally higher than 550 ° C, which is much higher than the temperature of the reaction zone. Local over-temperature at the inlet of the regenerated catalyst will reduce the selectivity of the olefins.
  • the technical problem to be solved by the present invention is the problem that the yield of low-carbon olefins in the prior art is not high, and the object is to provide a new method for increasing the yield of low-carbon olefins.
  • the method is used in the production of low-carbon olefins, and has the advantages of good catalyst carbon uniformity, high yield of low-carbon olefins, and good economics of low-carbon olefin production process.
  • the present invention provides a method for producing a low carbon olefin from an oxygen compound, comprising the following steps:
  • the spent catalyst flowing out from the nth secondary reaction zone is stripped and upgraded into a dense phase fluidized bed regenerator for regeneration;
  • the spent catalyst is serially passed through the first to mth secondary regeneration
  • the regeneration medium is fed into the first to mth secondary regeneration zones from the m feed zone of the regeneration zone in parallel, and the catalyst is contacted with the regeneration medium, and the carbon content is gradually decreased, and the catalyst after regeneration is completed.
  • the apparent apparent linear velocity of the gas in the material flow controller is less than or equal to the minimum fluidization velocity of the catalyst.
  • the apparent apparent linear velocity of the gas in the material flow controller is less than or equal to the minimum fluidization velocity of the catalyst.
  • the catalyst contains SAPO-34 molecular sieves.
  • the reaction conditions of the dense phase fluidized bed reactor are: the apparent carbon velocity of the pre-carbonized zone and the reaction zone is 0.1-1.5 m/s, and the pre-carbonized zone
  • the reaction temperature is 500-650 ° C
  • the reaction temperature of the reaction zone is 400-550 ° C
  • the bed reactor has a bed density of 200-1200 kg/m 3 .
  • the average carbon deposition amount of the catalyst in the first secondary pre-carbonized zone to the n-th secondary reaction zone in the dense phase fluidized bed reactor is sequentially increased, and the kth secondary pre-product
  • the average carbon deposition amount of the catalyst in the carbon zone is 0.5 to 3 wt%
  • the average carbon deposition amount of the catalyst in the n-th second-stage reaction zone is 7 to 10 wt%.
  • the reaction conditions of the dense phase fluidized bed regeneration zone are: an apparent gas velocity of 0.1-1.5 m/s, a reaction temperature of 500-700 ° C, and a bed density of 200-1200 kg. /m 3 o
  • the average carbon deposition amount of the catalyst in the first to mth secondary regeneration zone of the dense phase fluidized bed regeneration zone is successively decreased, and the average carbon deposition of the catalyst in the first secondary regeneration zone is averaged.
  • the amount of catalyst is 2-10wt%, and the average carbon deposition amount of the catalyst in the m-th secondary regeneration zone is 0-0.1wt% o
  • the oxygen-containing compound is methanol and/or dimethyl ether
  • the low-carbon olefin is any one of ethylene, propylene or butene or a mixture of any of the above
  • the hydrocarbon is any one or a mixture of any of naphtha, gasoline, condensate, light diesel oil, hydrogenated tail oil or kerosene
  • the regeneration medium is any of air, oxygen-depleted air or water vapor.
  • the catalyst after completion of the regeneration in the step is subsequently stripped and lifted back to the first secondary pre-carbonation zone of the dense phase fluidized bed, wherein the lift gas in the lifting process is water vapor, Any one or a mixture of any of a hydrocarbon of four or more carbon, naphtha, gasoline, condensate, light diesel oil, hydrogenated tail oil or kerosene.
  • FIG. 2 is a schematic structural view of a dense phase fluidized bed comprising two secondary pre-carbon zones and two second-stage reaction zones according to the present invention, wherein the arrows in the AA profile are secondary pre-carbon zones and two stages The flow direction of the catalyst in the reaction zone;
  • FIG. 3 is a schematic structural view of a dense phase fluidized bed including four secondary regeneration zones according to the present invention; Wherein the arrow in the BB profile is the flow direction of the catalyst in the secondary regeneration zone;
  • FIG. 4 is a schematic structural view of a stripper according to the present invention.
  • Figure 5 is a schematic view showing the structure of the material flow controller of the present invention.
  • the method provided by the present invention mainly comprises the following steps:
  • the spent catalyst flowing out from the nth secondary reaction zone is stripped and upgraded into a dense phase fluidized bed regenerator for regeneration;
  • the spent catalyst is serially passed through the first to mth secondary regeneration
  • the regeneration medium is fed into the first to mth secondary regeneration zones from the m feed zone of the regeneration zone in parallel, and the catalyst is contacted with the regeneration medium, and the carbon content is gradually decreased, and the catalyst after regeneration is completed.
  • the catalyst after completion of the regeneration in the step is subsequently stripped and lifted back to the first secondary pre-carbonation zone of the dense phase fluidized bed, and the lift gas in the stripping process may be water vapor or carbon four. Any one or a mixture of any of the above hydrocarbons, naphtha, gasoline, condensate, light diesel oil, hydrogenated tail oil or kerosene.
  • the apparent line velocity of the gas in the material flow controller is less than or equal to the minimum fluidization velocity of the catalyst.
  • the apparent line velocity of the gas in the material flow controller is less than or equal to the minimum fluidization velocity of the catalyst.
  • the catalyst contains SAPO-34 molecular sieves.
  • the reaction conditions of the dense phase fluidized bed reactor are: the apparent carbon velocity of the pre-carbon deposition zone and the reaction zone is 0.1-1.5 m/s, and the reaction temperature of the pre-carbonation zone is 500-650°. C, the reaction temperature of the reaction zone is 400-550 ° C, and the bed density is 200-1200 kg/m 3 .
  • the dense phase The average carbon deposition of the catalyst in the first secondary pre-carbon zone to the n-th second-stage reaction zone in the fluidized bed reactor increases in turn, and the average carbon deposition amount of the catalyst in the k-th secondary pre-carbon zone is 0.5-3 wt%, the average carbon deposition amount of the catalyst in the nth secondary reaction zone is 7-10 wt%.
  • reaction conditions of the dense phase fluidized bed regeneration zone are: gas apparent linear velocity of 0.1-1.5 m/s, reaction temperature of 500-700 ° C, and bed density of 200-1200 kg/m 3 .
  • the average carbon deposition amount of the catalyst in the first to mth secondary regeneration zone of the dense phase fluidized bed regeneration zone is successively decreased, and the average carbon deposition amount of the catalyst in the first secondary regeneration zone is 2- 10 wt%, the average carbon deposition amount of the catalyst in the m-th secondary regeneration zone is 0-0.1 wt%.
  • the oxygen-containing compound is methanol and/or dimethyl ether
  • the low-carbon olefin is any one of ethylene, propylene or butene or a mixture of any of the plurality
  • the carbon tetra hydrocarbon may also be derived from Any one or a mixture of any of naphtha, gasoline, condensate, light diesel oil, hydrogenated tail oil or kerosene.
  • a dense phase fluidized bed reactor comprising a pre-carbon deposition zone, a reaction zone, a gas-solid separation zone, a stripping zone, a pre-carbonation zone and a reaction zone separated by a material flow controller, pre-carbon deposition
  • the zone is separated by the material flow controller into k secondary pre-carbon zones, kl, and the reaction zone is divided into n secondary reaction zones by the material flow controller, ⁇ 1, each secondary pre-carbon zone and secondary reaction zone Can be fed independently;
  • a dense phase fluidized bed regenerator comprising a regeneration zone, a gas-solid separation zone, a stripping zone, and the regeneration zone is divided by the material flow controller into m secondary regeneration zones, m 2 , each secondary The regeneration zone can be fed independently.
  • the hydrocarbons of carbon four or more enter the k secondary secondary carbon deposition zones in the dense phase fluidized bed reactor in parallel, are contacted with the fully regenerated catalyst, and are converted into a stream including the low carbon olefin product, and at the same time, the catalyst
  • the oxygenate-containing feedstock enters the nth secondary reaction zone in the dense phase fluidized bed reactor in parallel, and is contacted with the pre-carbonated catalyst to form a stream comprising the low-carbon olefin product and the catalyst to be produced, and
  • the pre-carbon deposition catalyst sequentially passes through the first to the nth secondary reaction zones in series, and the carbon content is gradually increased;
  • the catalyst to be produced flowing out of the nth second-stage reaction zone is stripped and lifted into the catalyst.
  • the dense phase fluidized bed regenerator is regenerated, and the catalyst to be continuously passed through the first to mth secondary regeneration zones in series, in contact with the regeneration medium, the carbon content gradually decreases to near zero, and then is returned by stripping and lifting.
  • the low-carbon olefin product stream is separated from the entrained catalyst and enters a separation section, and the separated catalyst is introduced into the n-th second-stage reaction zone;
  • the carbon by-products of the carbon four or more obtained in the separation section are returned to the pre-carbonized zone in the dense phase fluidized bed reactor.
  • FIG. 1 a schematic diagram of the process for increasing the yield of low carbon olefins from the oxygenate to light olefins process of the present invention is shown in FIG.
  • the secondary reaction zone (2-3, 2-4) in the fluidized bed reactor (2) is contacted with a pre-carbon catalyst to form a gas phase product stream and a catalyst to be produced; a gas phase of the pre-carbon deposition zone and the reaction zone
  • the product stream and the entrained catalyst enter the cyclone separator (3), and the gas phase product stream enters the subsequent separation section through the outlet of the cyclone separator
  • the regenerated catalyst is serially passed through the first secondary pre-carbonized zone to the second secondary reaction zone (2-1, ..., 2-4) in the dense phase fluidized bed reactor (2).
  • the entrained regenerated catalyst enters the fourth secondary regeneration zone (10-4) through the feed leg of the cyclone separator; from the dense phase fluidized bed reactor (2)
  • the spent catalyst passes through the stripper (5) and the riser (7) into the dense phase fluidized bed regenerator (10), wherein the bottom of the stripper (5) is connected to the water vapor line (6), the riser (7) The bottom is connected to the lift gas line (8), and the catalyst to be produced is sequentially ordered in the dense phase fluidized bed regenerator (10).
  • the first to fourth secondary regeneration zones (10-1, ..., 10-4) are serially passed through to form a regenerated catalyst after charring.
  • the lifting gas in the riser (7) may be any one or a mixture of any one of water vapor, carbon four or more hydrocarbons, naphtha, gasoline, condensate, light diesel oil, hydrogenated tail oil or kerosene. .
  • FIG. 1 a schematic diagram of the structure of a dense phase fluidized bed reactor comprising two secondary pre-carbon regions and two secondary reaction zones in the reactor of the present invention is shown in FIG.
  • Three material flow controllers (17) and one baffle are vertically arranged to separate two secondary pre-carbon zones and two second-stage reaction zones, and the catalyst sequentially passes through the first secondary pre-carbonation zone.
  • the second secondary pre-carbon zone, the first secondary reaction zone, and the second secondary reaction zone are then passed to the stripper.
  • the structure of the dense phase fluidized bed regenerator of the reactor of the present invention comprising four secondary regeneration zones is shown in Figure 3.
  • Three material flow controllers (17) and one baffle are vertically arranged to divide the regeneration zone into four secondary regeneration zones.
  • the catalysts are serially passed through the first to fourth secondary regeneration zones, and then enter the steam. Lifter.
  • the structural schematic of the stripper of the present invention is shown in FIG.
  • the opening in the upper tube wall of the stripper serves as the material overflow port (18) between the nth secondary reaction zone (or the mth secondary regeneration zone:) and the stripper.
  • the material flow controller (17) consists of a partition (19), an orifice (20), a material downstream flow tube (21), a bottom baffle (22), and a heat take-up component (23).
  • the catalyst enters the material downstream flow tube from above the downstream flow tube, wherein the apparent line velocity of the gas is less than or equal to the minimum fluidization speed, and the catalyst in the downstream flow tube of the material is in a dense phase accumulation state, forming a material flow driving force, pushing the catalyst into the orifice through the orifice.
  • the heat taking part can be fixed on the partition by a coil structure.
  • the apparent carbon velocity in the pre-carbonized zone and the reaction zone in the dense phase fluidized bed reactor is 0.1-1.5 m/s; the apparent linear velocity of the gas in the dense phase fluidized bed regeneration zone is 0.1-1.5 m / s; the apparent flow velocity of the gas in the material flow controller is less than or equal to the minimum fluidization velocity of the catalyst;
  • the catalyst comprises SAPO-34 molecular sieve; k feeds are provided at the bottom of the pre-carbonation zone
  • the feed includes hydrocarbons of four or more carbons, naphtha, gasoline, etc.; the bottom of the reaction zone is provided with n feed ports, the feed includes methanol, dimethyl ether, etc.; the stripping medium of the stripping zone Containing water vapor; the bottom of the regeneration zone is provided with a regeneration medium inlet, and the regeneration medium includes air, oxygen-poor air, water vapor, and the like;
  • the reaction temperature of the pre-carbon deposition zone is 500-650 ° C, the reaction temperature of the reaction
  • the present invention may also use naphtha, gasoline, condensate, light diesel oil, hydrogenated tail oil or/and kerosene instead of carbon four or more as raw materials for the pre-carbonized zone in a dense phase fluidized bed reactor.
  • These hydrocarbons also have the effect of lowering the temperature of the regenerated catalyst and pre-carbonizing the regenerated catalyst.
  • the lift gas in the riser (15) may be steam, carbon or more hydrocarbons, naphtha, gasoline, condensate, light diesel oil, hydrogen tail oil or/and kerosene.
  • the purpose of controlling the carbon deposition amount of the catalyst, improving the uniformity of the carbon content and increasing the yield of the low carbon olefin can be achieved, and the technical advantage is large, and it can be used in the industrial production of the low carbon olefin.
  • the beneficial effects of the invention include: (1) the dense phase fluidized bed has a higher bed density, the catalyst speed is lower, and the wear is low; (2) the gas velocity in the material flow controller in the material flow controller Less than or equal to the minimum fluidization velocity of the catalyst, the catalyst is in a dense phase accumulation state, forming a unidirectional dense phase transport stream of the catalyst, avoiding catalyst backmixing between adjacent secondary reaction zones (or adjacent secondary regeneration zones) (3)
  • the heat taking part in the material flow controller has the function of controlling the temperature of the reaction zone; (4)
  • the material flow controller divides the dense phase fluidized bed reactor into the pre-carbonized zone and the reaction zone And dividing the pre-carbonation zone into k secondary pre-carbon zones, separating the reaction zone into n second-stage reaction zones, and the catalyst sequentially passes through the first secondary pre-carbon zone to the n-th stage In the reaction zone, the residence time distribution is narrow, and the uniformity of the carbon content of the pre-carbon catalyst and the catalyst to be produced is greatly improved; (5) The regenerated
  • k secondary pre-carbonation zones, n secondary reaction zones and m secondary regeneration zones can be independently fed, with strong operational flexibility; (9) Achieving more precise control of regenerated catalyst and waiting The carbon content of the catalyst, and the carbon content distribution is relatively uniform, the selectivity of the low-carbon olefin is improved, and the carbon content can be adjusted according to the demand to optimize the ratio of propylene/ethylene; (10) the carbon content distribution of the catalyst is relatively uniform, the reaction zone The required catalyst reserves are reduced; (11) The structure of a plurality of secondary pre-carbon deposition zones, reaction zones, and regeneration zones facilitates the enlargement of the reactor.
  • Example 1 In order to better explain the present invention, it is convenient to understand the technical solution of the present invention, and a typical but non-limiting embodiment of the present invention is as follows: Example 1
  • a second-stage pre-carbonation zone and three secondary reaction zones are arranged in the dense-phase fluidized bed reactor, and four secondary regeneration zones are arranged in the dense-phase fluidized bed regenerator.
  • the hydrocarbons of carbon 4 or higher enter the first secondary pre-carbonation zone in the dense-phase fluidized bed reactor, are contacted with the fully regenerated catalyst, and are converted into products including low-carbon olefins, and at the same time, the amount of carbon deposited on the catalyst reaches a certain value, forming a pre-carbon catalyst, the pre-carbon catalyst enters the reaction zone; the oxygen-containing material enters the first to third secondary reaction zones in parallel in the dense-phase fluidized bed reactor while pre-carbon deposition
  • the catalyst sequentially passes through the first to third secondary reaction zones, and the oxygenate-containing raw material is contacted with the pre-carbon catalyst to form a product including a low-carbon olefin and a deactivated spent catalyst; including a low-carbon olefin
  • the regenerated catalyst is formed after the reaction; the regenerated catalyst passes through the stripper and the riser and then enters the dense-phase fluidized bed reactor, and sequentially passes through the first secondary pre-carbonation zone and the first secondary reaction. From the zone to the fourth secondary reaction zone; the hydrocarbon by-product of the carbon four or more obtained in the separation section is returned to the first secondary pre-carbonation zone in the dense phase fluidized bed reactor; the lift gas in the riser 15 is carbon More than four hydrocarbons.
  • the reaction conditions of the dense phase fluidized bed reactor are as follows: the temperature of the first secondary pre-carbon deposition zone is 500 °C, the temperature of the first to third secondary reaction zone is 400 °C, and the gas phase linear velocity is 0.3 m/ s, the bed density is 1000k g / m 3 , the average carbon deposition of the first secondary pre-carbon zone is lwt%, the first The average carbon deposition in the secondary reaction zone is 5wt%, the average carbon deposition in the second secondary reaction zone is 8wt%, and the average carbon deposition in the third secondary reaction zone is 10wt%; dense phase fluidization
  • the bed regenerator reaction conditions are: reaction temperature is 550 °C, gas phase linear velocity is 0.3 m/s, bed density is 1000 kg/m 3 , and average carbon deposition in the first secondary regeneration zone is 5 wt%, second The average carbon deposition of the secondary regeneration zone is 2wt%, the average carbon deposition of the third secondary regeneration zone is 0.5wt%, and the average carbon
  • one secondary pre-carbonation zone and two secondary reaction zones are arranged, and two secondary regeneration zones are arranged in the dense phase fluidized bed regenerator.
  • the first secondary pre-carbonation zone entering the dense-phase fluidized bed reactor such as carbon four or more, is contacted with the fully regenerated catalyst, converted into a product including low-carbon olefins, and the carbon content on the catalyst is reached.
  • a pre-carbon catalyst the pre-carbon catalyst enters the reaction zone; the oxygen-containing material enters the first to the second secondary reaction zone in parallel in the dense-phase fluidized bed reactor while pre-carbon deposition
  • the catalyst sequentially passes through the first to the second secondary reaction zone, and the oxygenate-containing raw material is contacted with the pre-carbon catalyst to form a product including a low-carbon olefin and a deactivated catalyst to be produced; including a low-carbon olefin
  • the gas phase product stream and the entrained catalyst are introduced into the cyclone separator, and the gas phase product stream enters the subsequent separation section through the outlet of the cyclone separator, and the entrained catalyst is introduced into the second secondary reaction zone through the material leg of the cyclone separator;
  • the catalyst to be produced is passed from the second secondary reaction zone through the stripper and the riser to the dense phase fluidized bed regenerator, and sequentially passes through the first to the second secondary regeneration zone, and is regenerated.
  • the regenerated catalyst is formed after the reaction; the regenerated catalyst passes through the stripper and the riser and then enters the dense-phase fluidized bed reactor, and sequentially passes through the first secondary pre-carbonation zone and the first secondary reaction. Zone and the second secondary reaction zone; the hydrocarbon by-product of the carbon four or more obtained in the separation section is returned to the first secondary pre-carbonation zone in the dense phase fluidized bed reactor; the lift gas in the riser 15 is gasoline .
  • the reaction conditions of the dense phase fluidized bed reactor are as follows: the temperature of the first secondary pre-carbon deposition zone is 550 ° C, the temperature of the first to the second secondary reaction zone is 450 ° C, and the gas phase linear velocity is 0.5 m / s, the bed density is 900kg/m 3 , the average carbon deposition in the first secondary pre-carbonation zone is 2wt%, and the average carbon deposition in the first secondary reaction zone is 6wt%, the second secondary The average carbon deposition amount in the reaction zone is 8 wt%; the reaction condition of the dense phase fluidized bed regenerator is: the reaction temperature is 600 ° C, and the gas phase linear velocity is 0.7 m/s.
  • the bed density was 700 kg/m 3
  • the average carbon deposition in the first secondary regeneration zone was 3 wt%
  • the average carbon deposition in the second secondary regeneration zone was 0.1 wt%.
  • the reaction product was analyzed by on-line gas chromatography, and the yield of the low carbon olefin carbon group was 91.2% by weight.
  • one secondary pre-carbonation zone and five secondary reaction zones are arranged, and five secondary regeneration zones are arranged in the dense phase fluidized bed regenerator.
  • the naphtha and the carbon four or more hydrocarbons are mixed and then enter the first secondary pre-carbonation zone in the dense-phase fluidized bed reactor, contact with the fully regenerated catalyst, and converted into a product including a low-carbon olefin, and at the same time, a catalyst
  • the amount of carbon deposited reaches a certain value to form a pre-carbon catalyst, and the pre-carbon catalyst enters the reaction zone; the raw material containing the oxygen compound enters the first to fifth secondary reaction zones in parallel in the dense-phase fluidized bed reactor.
  • the pre-carbon catalyst is sequentially passed through the first to fifth secondary reaction zones, and the oxygenate-containing raw material is contacted with the pre-carbon catalyst to form a product including a low-carbon olefin and a deactivated catalyst.
  • the gas phase product stream including the low carbon olefin and the entrained catalyst are introduced into the cyclone separator, and the gas phase product stream enters the subsequent separation section through the outlet of the cyclone separator, and the entrained catalyst enters the fifth stage through the material leg of the cyclone separator
  • the second reaction zone; the catalyst to be produced is passed from the fifth secondary reaction zone through the stripper and the riser into the dense phase fluidized bed regenerator, and sequentially passes through the first to fifth secondary regenerations.
  • the regenerated catalyst in contact with the regeneration medium, forming a regenerated catalyst after the reaction; the regenerated catalyst passes through the stripper and the riser and then enters the dense phase fluidized bed reactor, and sequentially passes through the first secondary pre-carbonation zone, first From the secondary reaction zone to the fifth secondary reaction zone; the hydrocarbon by-product of the carbon four or more obtained in the separation section is returned to the first secondary pre-carbonation zone in the dense-phase fluidized bed reactor;
  • the lift gas uses hydrocarbons of carbon four or more.
  • the reaction conditions of the dense phase fluidized bed reactor are as follows: the temperature of the first secondary pre-carbon deposition zone is 650 ° C, the temperature of the first to fifth secondary reaction zone is 550 ° C, and the gas phase linear velocity is 0.7 m / s, the bed density is 700kg/m 3 , the average carbon deposition of the first secondary pre-carbon zone is 0.5wt%, and the average carbon deposition of the first secondary reaction zone is 2.5wt%, the second The average carbon deposition in the secondary reaction zone is 4 wt%, the average carbon deposition in the third secondary reaction zone is 5 wt%, and the average carbon deposition in the fourth secondary reaction zone is 6 wt%, the fifth secondary The average carbon deposition in the reaction zone is 7wt%; the reaction conditions of the dense phase fluidized bed regenerator are: reaction temperature is 700 ° C, gas phase linear velocity is 1.0 m / s, bed density is 500 kg / m 3 , the first The average carbon deposition in the secondary regeneration zone is 5wt
  • the amount is 1.5 wt%, the average carbon deposition amount of the fourth secondary regeneration zone is 0.05 wt%, and the average carbon deposition amount of the fifth secondary regeneration zone is 0.01 wt%.
  • the reaction product was analyzed by on-line gas chromatography, and the yield of the low carbon olefin carbon group was 92.5 wt%.
  • Two secondary pre-carbon zones and four second-stage reaction zones are arranged in the dense-phase fluidized bed reactor, and four secondary regeneration zones are arranged in the dense-phase fluidized bed regenerator.
  • Hydrocarbons of more than four carbons are fed in parallel to the first secondary pre-carbon zone and the second secondary pre-carbon zone in the dense-phase fluidized bed reactor, in contact with the fully regenerated catalyst, and converted to include low-carbon olefins.
  • the catalyst serially passes through the first secondary pre-carbonation zone and the second secondary pre-carbonation zone, and the carbon deposition amount reaches a certain value to form a pre-carbon deposition catalyst, and the pre-carbon deposition catalyst enters the reaction zone;
  • the oxygenate-containing feedstock enters the first to fourth secondary reaction zones in the dense phase fluidized bed reactor in parallel, and the pre-carbon deposition catalyst sequentially passes through the first to fourth secondary reaction zones in sequence.
  • the oxygenate-containing feedstock is contacted with a pre-carbonaceous catalyst to produce a product comprising a lower olefin and a deactivated spent catalyst; a gas phase product stream comprising a lower olefin and an entrained catalyst to enter the cyclone, a gas phase product
  • the flow enters the subsequent separation section through the outlet of the cyclone separator, and the entrained catalyst enters the fourth secondary reaction zone through the feed leg of the cyclone separator;
  • the catalyst to be produced is lifted by the fourth secondary reaction zone through the stripper Guan Jin a dense-phase fluidized bed regenerator, which passes through the first to fourth secondary regeneration zones in series, in contact with the regeneration medium, and forms a regenerated catalyst after the reaction;
  • the regenerated catalyst passes through the stripper, the riser and then enters the dense phase a fluidized bed reactor, and serially passes through the first secondary pre-carbonation zone, the second secondary pre-carbonation zone, the first secondary reaction zone to the fourth secondary reaction zone;
  • the reaction conditions of the dense phase fluidized bed reactor are: the temperature of the first secondary pre-carbonation zone and the second secondary pre-carbonation zone is 650 ° C, and the temperature of the first to fourth secondary reaction zones is 500. °C, the gas phase linear velocity is 1.0m/s, the bed density is 500kg/m 3 , and the average secondary carbon deposition in the first secondary pre-carbon deposition zone is 1.5wt%, the second secondary pre-carbonized zone
  • the average carbon deposition amount is 3.0 wt%, the average carbon deposition amount in the first secondary reaction zone is 4.5 wt%, and the average carbon deposition amount in the second secondary reaction zone is 6.0 wt%, and the third secondary reaction zone
  • the average carbon deposition amount is 7.0 wt%, and the average carbon deposition amount in the fourth secondary reaction zone is 8.0 wt%;
  • the dense phase fluidized bed regenerator reaction conditions are: reaction temperature is 700 ° C, gas phase The linear velocity is l.Om/s, the bed density is 500kg
  • the average carbon deposition amount of the third secondary regeneration zone is 1.2 wt%, and the average carbon deposition amount of the fourth secondary regeneration zone is 0.02 wt%.
  • the reaction product was analyzed by on-line gas chromatography, and the yield of the low carbon olefin carbon group was 93.2% by weight.
  • Two secondary pre-carbon zones and two second-stage reaction zones are arranged in the dense-phase fluidized bed reactor, and four secondary regeneration zones are arranged in the dense-phase fluidized bed regenerator. Hydrocarbons of more than four carbons are fed in parallel to the first secondary pre-carbon zone and the second secondary pre-carbon zone in the dense-phase fluidized bed reactor, in contact with the fully regenerated catalyst, and converted to include low-carbon olefins.
  • the catalyst serially passes through the first secondary pre-carbonation zone and the second secondary pre-carbonation zone, and the carbon deposition amount reaches a certain value to form a pre-carbon deposition catalyst, and the pre-carbon deposition catalyst enters the reaction zone;
  • the oxygenate-containing feedstock enters the first to the second secondary reaction zone in the dense phase fluidized bed reactor in parallel, and the pre-carbonated catalyst sequentially passes through the first to the second secondary reaction zone in sequence.
  • the oxygenate-containing feedstock is contacted with a pre-carbonaceous catalyst to produce a product comprising a lower olefin and a deactivated spent catalyst; a gas phase product stream comprising a lower olefin and an entrained catalyst to enter the cyclone, a gas phase product
  • the flow enters the subsequent separation section through the outlet of the cyclone separator, and the entrained catalyst enters the second secondary reaction zone through the material leg of the cyclone separator;
  • the catalyst to be produced is passed from the second secondary reaction zone through the stripper and is lifted Guan Jin a dense-phase fluidized bed regenerator, which passes through the first to fourth secondary regeneration zones in series, in contact with the regeneration medium, and forms a regenerated catalyst after the reaction;
  • the regenerated catalyst passes through the stripper and the riser and then enters the dense a fluidized bed reactor, and serially passes through a first secondary pre-carbonation zone, a second secondary pre-carbonation zone, a first secondary reaction zone, and
  • the reaction conditions of the dense phase fluidized bed reactor are: the temperature of the first secondary pre-carbonation zone and the second secondary pre-carbonation zone is 650 ° C, and the temperature of the first to the second secondary reaction zone is 500. °C, the gas phase linear velocity is 1.0m/s, the bed density is 500kg/m 3 , and the average secondary carbon deposition in the first secondary pre-carbon deposition zone is 1.5wt%, the second secondary pre-carbonized zone The average carbon deposition amount is 3.0 wt%, the average carbon deposition amount in the first secondary reaction zone is 6.0 wt%, and the average carbon deposition amount in the second secondary reaction zone is 8.5 wt%; dense phase fluidized bed regenerator
  • the reaction conditions are: the reaction temperature is 700 ° C, and the gas phase linear velocity is 1.0 m / s, The bed density is 500kg/m 3 , the average carbon deposition in the first secondary regeneration zone is 5.8wt%, and the average carbon deposition in the second secondary regeneration zone is 3wt%.
  • the third secondary regeneration zone The average carbon deposition amount was 1.1% by weight, and the average secondary carbon deposition amount of the fourth secondary regeneration zone was 0.02% by weight.
  • the reaction product was analyzed by on-line gas chromatography, and the yield of the low carbon olefin carbon group was 92.8 wt%.

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Abstract

本发明涉及一种提高含氧化合物制低碳烯烃工艺中提高低碳烯烃收率的方法。更具体地，本发明提供一种含氧化合物制低碳烯烃的方法，以包含k个二级预积碳区（K≥1）、n个二级反应区的多级密相流化床为反应器（n≥1）和包含m个二级再生区（m≥2）的多级密相流化床再生器为主体设备，通过将分离段所得碳四以上的烃回炼，或者在反应区添加石脑油、汽油、凝析油、轻柴油、加氢尾油或煤油的方法，主要解决现有技术中催化剂积碳量、碳含量均匀性难以控制，低碳烯烃收率较低的问题。

Description

一种含氧化合物制低碳烯烃的方法 技术领域
本发明涉及一种制低碳烯烃的工艺方法， 其具有提高的低碳烯烃收 率。 背景技术
低碳烯烃， 即乙烯和丙烯， 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。 甲醇转化 制烯烃 (MTO)的工艺受到越来越多的重视， 已实现百万吨级的生产规模。 随着世界经济的发展， 低碳烯烃， 特别是丙烯， 需求量与日倶增， 析迈公 司 (CMAI)分析称， 在 2016年前， 乙烯需求量将以年均 4.3%的速度增长， 丙烯需求量将以年均 4.4%的速度增长。 由于我国经济的高速增长， 我国 乙烯和丙烯的需求量的年增长率均超过世界平均水平。
20世纪 80年代初， UCC公司成功开发出了 SAPO系列分子筛，其中 SAPO-34分子筛催化剂在用于 MTO反应时表现出优异的催化性能， 具有 很高的低碳烯烃选择性， 而且活性很高， 但催化剂在使用一段时间后由于 积碳而失去活性。 SAPO-34 分子筛催化剂在使用过程中存在明显的诱导 期， 在诱导期内， 烯烃的选择性较低， 垸烃的选择性较高， 随着反应时间 的增加， 低碳烯烃选择性逐渐上升， 诱导期过后， 催化剂在一定时间内保 持高的选择性和高的活性，随着时间的继续延长，催化剂的活性迅速下降。
US6166282中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歩分离出大部分的夹 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。但该方法中低碳烯烃碳基收率一 般均在 77%左右， 存在低碳烯烃收率较低的问题。 CN101402538B公布了一种提高低碳烯烃收率的方法， 该方法采用在 甲醇转化为低碳烯烃的第一反应区上部设置一个第二反应区，且该第二反 应区直径大于第一反应区，以增加第一反应区出口的产品气体在第二反应 区内的停留时间， 使得未反应的甲醇、生成的二甲醚和碳四以上的烃继续 反应， 达到提高低碳烯烃收率的目的， 该方法虽然可以在一定程度上提高 低碳烯烃的收率， 但是由于第一反应区出来的催化剂已经带有较多的积 碳， 而碳四以上的烃裂解需要较高的催化剂活性， 因此该方法中第二反应 区内的碳四以上的烃转化效率仍然偏低， 从而导致低碳烯烃收率偏低。
CN102276406 A公布了一种增产丙烯的生产方法。该技术设置三个反 应区， 第一快速床反应区用于甲醇转化至烯烃， 提升管反应区和第二快速 床反应区串联用于转化乙烯、碳四以上的烃和未反应的甲醇或二甲醚。此 专利中碳四以上的烃等物质在提升管反应区和第二快速床反应区中的停 留时间较短， 转化效率偏低， 从而导致丙烯收率偏低。
CN102875289A 公布了一种内部布置提升管反应器的流化床反应装 置， 用于提高低碳烯烃的产率。第一原料进入流化床反应区， 与催化剂接 触， 生成包括低碳烯烃的产品， 同时形成待生催化剂； 待生催化剂一部分 进入再生器再生， 形成再生催化剂，一部分进入出口端位于反应区内部的 提升管， 与第二原料接触， 将待生催化剂提升至反应区内； 再生催化剂返 回流化床反应器反应区。此专利所披露的反应装置无汽提部分， 待生催化 剂将会携带部分产品气体进入再生器， 与氧气发生燃烧， 降低低碳烯烃的
CN102875296A公布的甲醇制烯烃技术设置了快速床、下行床和提升 管三个反应区。 催化剂在再生器、 快速床、 提升管和下行床之间循环， 流 向十分复杂、 流量分配和控制十分困难， 催化剂的活性变化较大。
本领域所公知的， 低碳烯烃的选择性和催化剂上的积碳量密切相关， 要保证高的低碳烯烃选择性， SAPO-34催化剂上需要一定数量的积碳。 目 前 MTO工艺所采用的主要反应器为流化床， 而流化床接近于全混流反应 器， 催化剂积炭分布很宽， 不利于提高低碳烯烃的选择性。 MTO工艺的 剂醇比很小， 生焦率较低， 要实现较大的、 容易控制的催化剂循环量， 就 需要在再生区中将催化剂上的积碳量、碳含量均匀性控制在一定水平， 进 而达到控制反应区内催化剂上的积碳量、碳含量均匀性的目的。 因此， 控 制反应区内的催化剂积碳量和碳含量均匀性于某一水平是 MTO工艺中的 关键技术。
另外， 在甲醇制备低碳烯烃的过程中， 不可避免的会产生碳四以上的 烃等副产物， 并且碳四以上的烃的碳基选择性一般在 10wt%左右。若将这 些副产物进一歩转化为目的产物， 将会提高低碳烯烃的收率和工艺经济 性。 我们的研究已表明不含碳的 SAPO-34分子筛催化剂用于碳四以上的 烃转化制低碳烯烃具有很好的效果， 并且可以实现在 SAPO-34分子筛催 化剂上预积碳， 提高其低碳烯烃选择性。低碳烯烃对反应温度等工艺参数 也十分敏感， 例如再生催化剂的温度一般高于 550°C， 远高于反应区的温 度， 再生催化剂入口处局部超温将降低析碳烯烃的选择性。
催化剂积炭分布不均、碳含量难以控制、反应温度波动大等均不利于 提高低碳烯烃的收率， 为解决上述问题， 一些研究者提出了在流化床内设 置上下两个反应区、 两个流化床串联、 流化床和提升管、 下行床串联等技 术， 取得了一定的有益效果， 但也同时增加 MTO工艺的复杂性、 控制难 度增加。本发明提出了通过在密相流化床中设置内构件形成多个二级反应 区 (再生区:)的方案来提高低碳烯烃的收率。 发明内容
本发明所要解决的技术问题是现有技术中存在的低碳烯烃收率不高 的问题， 目的在于提供一种新的提高低碳烯烃收率的方法。该方法用于低 碳烯烃的生产中， 具有催化剂积碳均匀性好、低碳烯烃收率较高、 低碳烯 烃生产工艺经济性较好的优点。
为实现上述目的， 本发明提供一种含氧化合物制低碳烯烃的方法， 包 括以下歩骤：
a) 将碳四以上的烃从 k个预积碳区进料支线并行通入密相流化床反应 器中的 k个二级预积碳区， 与再生完全和 /或新鲜的催化剂接触， 转化为 包含低碳烯烃产品的物流， 同时形成预积碳催化剂； 其中所述催化剂依序 串行通过第 1个至第 k个二级预积碳区， 碳含量逐渐增加； 其中所述密相 流化床反应器由物料流动控制器分隔为预积碳区和反应区；其中所述密相 流化床反应器的预积碳区由物料流动控制器分隔为 k个二级预积碳区，第
1个至第 k个二级预积碳区依序相联；
b) 将含有含氧化合物的原料从 n个反应区进料支线并行通入密相流化 床反应器中的 n个二级反应区， 与所述预积碳催化剂接触， 生成包含低碳 烯烃产品的物流和待生催化剂；其中所述从第 k个二级预积碳区流入的预 积碳催化剂依序串行通过第 1个至第 n个二级反应区， 碳含量逐渐增加； 其中所述密相流化床反应器的反应区由物料流动控制器分隔为 n个二级 反应区， 第 1个至第 n个二级反应区依序相联， 第 1个二级反应区连接于 第 k个二级预积碳区的下游；
c) 将由所述预积碳区和反应区流出的所述包含低碳烯烃产品的物流 与其携带的待生催化剂分离； 分离出的待生催化剂进入第 n个二级反应 区； 所述包含低碳烯烃产品的物流进入产品分离工段， 经分离、提纯得到 低碳烯烃产品，分离工段所得的碳四以上的烃副产品返回至所述密相流化 床反应器中的预积碳区；
d) 由第 n个二级反应区流出的待生催化剂经过汽提、提升进入密相流 化床再生器再生； 所述待生催化剂依序串行通过第 1个至第 m个二级再 生区； 再生介质从 m个再生区进料支线并行通入第 1个至第 m个二级再 生区， 所述待生催化剂与所述再生介质接触， 碳含量逐渐下降， 完成再生 后的催化剂随后经汽提、提升返回第 1个二级预积碳区； 其中所述密相流 化床再生器由物料流动控制器分隔为 m个二级再生区， 第 1个至第 n个 二级再生区依序相联；
其中， k≥l， n≥l， m≥2。 更优选地， 4 k 2， 8^n^3 , 8 m 3。 在一个优选实施方式中， 在所述密相流化床反应器中， 物料流动控制 器中气体表观线速度小于等于催化剂的最小流化速度。
在一个优选实施方式中， 在所述密相流化床再生器中， 物料流动控制 器中气体表观线速度小于等于催化剂的最小流化速度。
在一个优选实施方式中， 所述催化剂含有 SAPO-34分子筛。
在一个优选实施方式中， 所述密相流化床反应器的反应条件为： 所述 预积碳区和反应区的气体表观线速度为 0.1-1.5m/s， 所述预积碳区的反应 温度为 500-650°C， 所述反应区的反应温度为 400-550°C， 所述密相流化 床反应器的床层密度为 200-1200kg/m3。
在一个优选实施方式中， 所述密相流化床反应器中第 1个二级预积碳 区至第 n个二级反应区内催化剂平均积炭量依次递增，第 k个二级预积碳 区内的催化剂平均积炭量为 0.5-3wt%， 第 n个二级反应区内的催化剂平 均积炭量为 7-10wt%。
在一个优选实施方式中， 所述密相流化床再生区的反应条件为： 气体 表观线速度为 0.1-1.5m/s， 反应温度为 500-700 °C， 床层密度为 200-1200kg/m3 o
在一个优选实施方式中， 所述密相流化床再生区第 1个至第 m个二级 再生区内的催化剂平均积炭量依次递减，第 1个二级再生区内的催化剂平 均积炭量为 2-10wt%， 第 m 个二级再生区内的催化剂平均积炭量为 0-0.1wt% o
在一个优选实施方式中， 所述含氧化合物为甲醇和 /或二甲醚； 所述低 碳烯烃为乙烯、丙烯或丁烯中的任意一种或任意几种的混合物； 所述碳四 以上的烃为石脑油、 汽油、 凝析油、 轻柴油、 加氢尾油或煤油中的任意一 种或任意几种的混合物； 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。
在一个优选实施方式中， 歩骤 中完成再生后的催化剂随后经汽提、 提升返回密相流化床的第 1个二级预积碳区，所述提升过程中的提升气采 用水蒸气、 碳四以上的烃、 石脑油、 汽油、 凝析油、 轻柴油、 加氢尾油或 煤油中的任意一种或任意几种的混合物。
采用本发明的所述方法， 即有效提高了目的产物低碳烯烃的收率， 又 优化了能量分配与利用。 附图说明
图 1为本发明所述方法的流程示意图；
图 2为本发明所述包含 2个二级预积碳区和 2个二级反应区的密相流 化床的结构示意图， 其中 A-A剖面图中的箭头是二级预积碳区和二级反 应区间的催化剂流动方向；
图 3为本发明所述包含 4个二级再生区的密相流化床的结构示意图， 其中 B-B剖面图中的箭头是二级再生区间的催化剂流动方向；
图 4为本发明所述汽提器的结构示意图；
图 5为本发明所述物料流动控制器的结构示意图。
附图中的附图标记说明如下：
1-反应器进料管线； 1-1 第 1个二级预积碳区进料支线； 1-2 第 2个 二级预积碳区进料支线； 1-3 第 1个二级反应区进料支线； 1-4 第 2个二 级反应区进料支线； 2-密相流化床反应器； 2-1 第 1个二级预积碳区； 2-2 第 2个二级预积碳区； 2-3 第 1个二级反应区； 2-4 第 2个二级反应区； 3-旋风分离器； 4-产品物料管线； 5-汽提器； 6-水蒸气管线； 7-提升管； 8-提升气管线； 9-再生器进料管线； 9-1 第 1个二级再生区进料支线； 9-2 第 2个二级再生区进料支线； 9-3 第 3个二级再生区进料支线； 9-4 第 4 个二级再生区进料支线； 10-密相流化床再生器； 10-1 第 1个二级再生区； 10-2 第 2个二级再生区； 10-3 第 3个二级再生区； 10-4 第 4个二级再生 区； 11-旋风分离器； 12-废气管线； 13-汽提器； 14-水蒸气管线； 15-提升 管； 16-提升气管线； 17-物料流动控制器； 18-物料溢流口； 19-隔板； 20- 孔口； 21-物料下行流动管； 22-底部挡板； 23-取热部件。 具体实施方式
为了提高含氧化合物制低碳烯烃工艺中的低碳烯烃收率，本发明提供 的方法主要包括以下歩骤：
a) 将碳四以上的烃从 k个预积碳区进料支线并行通入密相流化床反 应器中的 k个二级预积碳区， 与再生完全和 /或新鲜的催化剂接触， 转化为包含低碳烯烃产品的物流， 同时， 形成预积碳催化剂； 其中 所述催化剂依序串行通过第 1个至第 k个二级预积碳区，碳含量逐 渐增加；其中所述密相流化床反应器由物料流动控制器分隔为预积 碳区和反应区；其中所述密相流化床反应器的预积碳区由物料流动 控制器分隔为 k个二级预积碳区，第 1个至第 k个二级预积碳区依 序相联；
b) 将含有含氧化合物的原料从 n个反应区进料支线并行通入密相流 化床反应器中的 n个二级反应区， 与所述预积碳催化剂接触， 生成 包含低碳烯烃产品的物流和待生催化剂；其中所述从第 k个二级预 积碳区流入的预积碳催化剂依序串行通过第 1个至第 n个二级反应 区， 碳含量逐渐增加； 其中所述密相流化床反应器的反应区由物料 流动控制器分隔为 n个二级反应区，第 1个至第 n个二级反应区依 序相联， 第 1个二级反应区连接于第 k个二级预积碳区的下游； c) 将由预积碳区和反应区流出的所述包含低碳烯烃产品的物流与其 携带的待生催化剂分离；分离出的待生催化剂进入第 n个二级反应 区； 所述包含低碳烯烃产品的物流进入产品分离工段， 经分离、提 纯得到低碳烯烃产品，分离工段所得的碳四以上的烃副产品返回至 所述密相流化床反应器中的预积碳区；
d) 由第 n个二级反应区流出的待生催化剂经过汽提、提升进入密相流 化床再生器再生； 所述待生催化剂依序串行通过第 1个至第 m个 二级再生区； 再生介质从 m个再生区进料支线并行通入第 1个至 第 m个二级再生区， 所述待生催化剂与所述再生介质接触， 碳含 量逐渐下降， 完成再生后的催化剂随后经汽提、提升返回第 1个二 级预积碳区；其中所述密相流化床再生器由物料流动控制器分隔为 m个二级再生区， 第 1个至第 n个二级再生区依序相联；
优选地，歩骤 中完成再生后的催化剂随后经汽提、提升返回密相流 化床的第 1个二级预积碳区， 所述汽提过程中的提升气可采用水蒸气、碳 四以上的烃、 石脑油、 汽油、 凝析油、 轻柴油、 加氢尾油或煤油中的任意 一种或任意几种的混合物。
优选地， k^ l , n^ l , m 2。优选地， 4 k 2， 8^n^3 , 8 m 3。 优选地， 密相流化床反应器中， 物料流动控制器中气体表观线速度小 于等于催化剂的最小流化速度。
优选地， 密相流化床再生器中， 物料流动控制器中气体表观线速度小 于等于催化剂的最小流化速度。
优选地， 所述催化剂含有 SAPO-34分子筛。
优选地， 所述密相流化床反应器的反应条件为： 预积碳区和反应区的 气体表观线速度为 0.1-1.5m/s，预积碳区的反应温度为 500-650°C，反应区 的反应温度为 400-550°C， 床层密度为 200-1200kg/m3。优选地， 所述密相 流化床反应器中第 1个二级预积碳区至第 n个二级反应区内的催化剂平均 积炭量依次递增， 第 k 个二级预积碳区内的催化剂平均积炭量为 0.5-3wt%, 第 n个二级反应区内的催化剂平均积炭量为 7-10wt%。
优选地， 所述密相流化床再生区的反应条件为： 气体表观线速度为 0.1-1.5m/s， 反应温度为 500-700 °C， 床层密度为 200-1200kg/m3。
优选地， 所述密相流化床再生区第 1个至第 m个二级再生区内的催 化剂平均积炭量依次递减， 第 1 个二级再生区内的催化剂平均积炭量为 2-10wt%, 第 m个二级再生区内的催化剂平均积炭量为 0-0.1wt%。
优选地， 所述含氧化合物为甲醇和 /或二甲醚； 所述低碳烯烃为乙烯、 丙烯或丁烯中的任意一种或任意几种的混合物；所述碳四的烃也可以来自 石脑油、 汽油、 凝析油、 轻柴油、 加氢尾油或煤油中的任意一种或任意几 种的混合物。
本发明提供的技术方案还涉及：
(l)提供一种密相流化床反应器， 包括预积碳区、 反应区、 气固分离 区、汽提区， 预积碳区和反应区由物料流动控制器隔开， 预积碳区由物料 流动控制器分隔为 k个二级预积碳区， k l， 反应区由物料流动控制器分 隔为 n个二级反应区， η 1， 各个二级预积碳区和二级反应区均可独立进 料；
(2；)提供一种密相流化床再生器， 包括再生区、 气固分离区、 汽提区， 再生区由物料流动控制器分隔为 m个二级再生区， m 2， 各个二级再生 区均可独立进料.
优选地，碳四以上的烃等并行进入密相流化床反应器中的 k个二级预 积碳区， 与完全再生的催化剂接触， 转化为包括低碳烯烃产品的物流， 同 时， 催化剂依序串行通过第 1个至第 k个二级预积碳区， 积碳量达到一定 值， 形成预积碳催化剂， 预积碳催化剂进入反应区；
优选地， 含有含氧化合物的原料并行进入密相流化床反应器中的第 n 个二级反应区， 与预积碳催化剂接触， 生成包括低碳烯烃产品的物流和待 生催化剂，同时，预积碳催化剂依序串行通过第 1个至第 n个二级反应区， 碳含量逐渐增加；
优选地， 由第 n个二级反应区流出的待生催化剂经过汽提、提升进入 密相流化床再生器再生， 待生催化剂依序串行通过第 1个至第 m个二级 再生区， 与再生介质接触， 碳含量逐渐下降至接近于零， 随后经汽提、 提 升返回第 1个二级预积碳区；
优选地，所述低碳烯烃产品物流与夹带的待生催化剂分离后进入分离 工段， 分离出的待生催化剂进入第 n个二级反应区；
优选地，分离工段所得的碳四以上的烃副产品返回至密相流化床反应 器中的预积碳区。
在一个优选的实施方案中，本发明所述的提高含氧化合物制低碳烯烃 技术的低碳烯烃收率的流程示意图如图 1所示。将碳四以上的烃从预积碳 区进料支线 （1-1， 1-2) 并行通入密相流化床反应器 （2) 中的二级预积 碳区 （2-1， 2-2)， 与包括 SAPO-34分子筛的催化剂接触， 生成气相产品 物流和预积碳催化剂； 将含有含氧化合物的原料从反应区进料支线（1-3， 1-4) 并行通入密相流化床反应器（2) 中的二级反应区 （2-3， 2-4)， 与预 积碳催化剂接触， 生成气相产品物流和待生催化剂； 预积碳区和反应区的 气相产品物流和夹带的催化剂进入旋风分离器（3 )， 气相产品物流经旋风 分离器的出口、 产品物料管线 （4) 进入后续分离工段， 夹带的催化剂经 旋风分离器的料腿进入第 2个二级反应区（2-4); 来自密相流化床再生器 ( 10) 的再生催化剂经过汽提器 （13 )、 提升管 （15) 进入密相流化床反 应器 （2)， 其中汽提器 （13 ) 底部连接水蒸气管线 （14)， 提升管 （15 ) 底部连接提升气管线（16)， 再生催化剂在密相流化床反应器（2) 中依序 串行经过第 1个二级预积碳区至第 2个二级反应区 （2-1， ……， 2-4)， 积碳后形成待生催化剂； 将再生介质从再生器进料管线 （9 ) 及其支线 ( 9-1， ……， 9-4 ) 并行通入密相流化床再生器 （10 ) 中的二级再生区 ( 10-1， ……， 10-4)， 与待生催化剂接触， 烧炭后生成尾气和再生催化 剂， 尾气和夹带的再生催化剂进入旋风分离器 （11 )， 尾气经旋风分离器 的出口、 废气管线（12)进入尾气处理工段， 处理后排放， 夹带的再生催 化剂经旋风分离器的料腿进入第 4个二级再生区 （10-4); 来自密相流化 床反应器（2) 的待生催化剂经过汽提器（5)、 提升管（7)进入密相流化 床再生器（10)， 其中汽提器（5)底部连接水蒸气管线 （6)， 提升管 （7) 底部连接提升气管线（8)， 待生催化剂在密相流化床再生器（10) 中依序 串行经过第 1个至第 4个二级再生区 （10-1， ……， 10-4)， 烧炭后形成 再生催化剂。 提升管 （7 ) 中的提升气可采用水蒸气、 碳四以上的烃、 石 脑油、 汽油、 凝析油、 轻柴油、 加氢尾油或煤油中的任意一种或任意几种 的混合物。
在一个具体的实施方案中，本发明的反应器包括 2个二级预积碳区和 2个二级反应区的密相流化床反应器的结构示意图如图 2所示。 3个物料 流动控制器（17 )和一个挡板竖直设置， 分隔出 2个二级预积碳区和 2个 二级反应区， 催化剂依序串行通过第 1个二级预积碳区、第 2个二级预积 碳区、 第 1个二级反应区、 第 2个二级反应区， 然后进入汽提器。
在一个具体的实施方案中，本发明的反应器包括 4个二级再生区的密 相流化床再生器的结构示意图如图 3所示。 3个物料流动控制器（17 ) 和 一个挡板竖直设置， 将再生区分隔为 4个二级再生区， 催化剂依序串行通 过第 1个至第 4个二级再生区， 然后进入汽提器。
在一个具体的实施方案中， 本发明所述的汽提器的结构示意图如图 4 所示。 汽提器上部管壁上开口作为第 n个二级反应区 (或第 m个二级再生 区:)与汽提器之间的物料溢流口 （18)。
在一个具体实施方案中，本发明的物料流动控制器的结构示意图如图
5所示。 物料流动控制器 （17 ) 由隔板 （19)、 孔口 （20)、 物料下行流动 管 （21 )、 底部挡板 （22) 和取热部件 （23 ) 组成。 催化剂由下行流动管 上方进入物料下行流动管， 其中气体表观线速度小于等于最小流化速度， 物料下行流动管内的催化剂处于密相堆积状态， 形成物料流动推动力， 推 动催化剂经过孔口流入其后的二级预积碳区 (或反应区、 或再生区：)。 取热 部件可采用盘管结构， 固定于隔板之上。
优选地，所述密相流化床反应器中预积碳区和反应区内气体表观线速 度为 0.1-1.5m/s; 所述密相流化床再生区内气体表观线速度为 0.1-1.5m/s; 所述物料流动控制器内气体表观线速度小于等于催化剂的最小流化速度； 所述催化剂包括 SAPO-34分子筛； 所述预积碳区底部设有 k个进料口， 进料包括碳四以上的烃、 石脑油、汽油等； 所述反应区底部设有 n个进料 口， 进料包括甲醇、 二甲醚等； 所述汽提区的汽提介质包含水蒸气； 所述 再生区底部设有再生介质入口，再生介质包括空气、贫氧空气、水蒸气等； 所述预积碳区的反应温度为 500-650°C， 反应区的反应温度为 400-550°C， 预积碳区和反应区床层密度为 200-1200kg/m3， 第 1个二级预积碳区至第 n个二级反应区内催化剂平均积炭量依次递增， 第 k个二级预积碳区的平 均积炭量为 0.5-3wt%，第 n二级反应区的平均积炭量为 7-10wt%；所述再 生区的反应温度为 500-700°C，床层密度为 200-1200kg/m3，第 1个至第 m 个二级再生区内催化剂平均积炭量依次递减，第 1个二级再生区的平均积 炭量为 2-10wt%， 第 m个二级再生区的平均积炭量为 0-0.1wt%。
优选地， 本发明也可采用石脑油、 汽油、 凝析油、 轻柴油、 加氢尾油 或 /和煤油替代碳四以上的烃作为密相流化床反应器中预积碳区的原料， 这些烃类同样具有降低再生催化剂温度、 在再生催化剂上预积碳的作用。
优选地， 提升管（15 ) 中的提升气可采用水蒸气、 碳四以上的烃、 石 脑油、 汽油、 凝析油、 轻柴油、 加氢尾油或 /和煤油。 采用本发明的方法， 可以达到控制催化剂积碳量、改善碳含量均匀性以及提高低碳烯烃收率的 目的， 具有较大的技术优势， 可用于低碳烯烃的工业生产中。
本发明能产生的有益效果包括： （1)密相流化床具有较高的床层密度， 催化剂速度较低、 磨损低； （2)物料流动控制器中的物料下行流动管中的 气速小于等于催化剂的最小流化速度， 催化剂处于密相堆积状态， 形成了 催化剂的单向密相输送流， 避免了相邻二级反应区 （或相邻二级再生区） 之间的催化剂返混， 停留时间分布窄； （3)物料流动控制器中的取热部件 具有控制反应区温度的作用； （4)物料流动控制器将密相流化床反应器分 割出预积碳区和反应区， 并将预积碳区分割为 k个二级预积碳区， 将反应 区分隔为 n个二级反应区，催化剂依次串行通过第 1个二级预积碳区至第 n个二级反应区， 停留时间分布窄， 预积碳催化剂和待生催化剂碳含量的 均匀性大幅度提高； （5)活性高、 温度高的再生催化剂返回到第 1 个二级 预积碳区， 有利于碳四以上的烃向低碳烯烃的转化， 反应后， 再生催化剂 上的积碳量达到一定值，其在 MTO反应中的低碳烯烃选择性也同时提高； (6)预积碳区内发生的碳四以上的烃转化为低碳烯烃的反应为吸热反应， 降低了再生催化剂的温度、 减轻了反应区的取热负荷、 有效利用了热量、 同时避免了高温催化剂和含氧化合物的接触； （7)物料流动控制器将再生 区分隔为 m个二级再生区， 待生催化剂依次通过第 1个至第 m个二级再 生区， 停留时间分布窄， 烧炭后所得的再生催化剂上的积碳量接近于零；
(8) k个二级预积碳区、 n个二级反应区和 m个二级再生区均可独立进料， 操作灵活性强； （9)实现了较为精确的控制再生催化剂和待生催化剂的碳 含量， 并且碳含量分布较为均匀， 提高了低碳烯烃的选择性， 并可根据需 求调控碳含量来优化丙烯 /乙烯的比率；（10)因催化剂的碳含量分布较为均 匀， 反应区所需的催化剂藏量降低； （11)多个二级预积碳区、 反应区、 再 生区的结构便于实现反应器的大型化。
为更好地说明本发明， 便于理解本发明的技术方案， 本发明的典型但 非限制性的实施例如下： 实施例 1
密相流化床反应器内设置 1个二级预积碳区和 3个二级反应区，密相 流化床再生器内设置 4个二级再生区。碳四以上的烃等进入密相流化床反 应器中的第 1个二级预积碳区， 与完全再生的催化剂接触， 转化为包括低 碳烯烃的产品， 同时， 催化剂上积碳量达到一定值， 形成预积碳催化剂， 预积碳催化剂进入反应区；含有含氧化合物的原料并行进入密相流化床反 应器中的第 1个至第 3个二级反应区，同时预积碳催化剂依序串行通过第 1个至第 3个二级反应区， 含有含氧化合物的原料与预积碳催化剂接触， 生成包括低碳烯烃的产品和失活的待生催化剂；包括低碳烯烃的气相产品 物流和夹带的待生催化剂进入旋风分离器，气相产品物流经旋风分离器的 出口进入后续分离工段， 夹带的待生催化剂经旋风分离器的料腿进入第 3 个二级反应区； 待生催化剂由第 3个二级反应区经过汽提器、提升管进入 密相流化床再生器， 并依序串行经过第 1个至第 4个二级再生区， 与再生 介质接触， 反应后形成再生催化剂； 再生催化剂经过汽提器、 提升管再进 入密相流化床反应器， 并依序串行经过第 1个二级预积碳区、第 1个二级 反应区至第 4个二级反应区；分离工段所得的碳四以上的烃副产品返回至 密相流化床反应器中的第 1个二级预积碳区； 提升管 15中的提升气采用 碳四以上的烃。密相流化床反应器反应条件为： 第 1个二级预积碳区温度 为 500 °C，第 1个至第 3个二级反应区温度为 400 °C，气相线速度为 0.3m/s， 床层密度为 1000kg/m3， 第 1个二级预积碳区的平均积炭量为 lwt%， 第 1 个二级反应区的平均积炭量为 5wt%， 第 2个二级反应区的平均积炭量为 8wt%, 第 3个二级反应区的平均积炭量为 10wt%； 密相流化床再生器反 应条件为： 反应温度为 550 °C， 气相线速度为 0.3m/s， 床层密度为 lOOOkg/m3 , 第 1个二级再生区的平均积炭量为 5wt%， 第 2个二级再生区 的平均积炭量为 2wt%， 第 3个二级再生区的平均积炭量为 0.5wt%， 第 4 个二级再生区的平均积炭量为 0.02wt%。 反应产品采用在线气相色谱分 析， 低碳烯烃碳基收率为 91.9wt%。 实施例 2
密相流化床反应器内设置 1个二级预积碳区和 2个二级反应区，密相 流化床再生器内设置 2个二级再生区。碳四以上飞烃等进入密相流化床反 应器中的第 1个二级预积碳区， 与完全再生的催化剂接触， 转化为包括低 碳烯烃的产品， 同时， 催化剂上积碳量达到一定值， 形成预积碳催化剂， 预积碳催化剂进入反应区；含有含氧化合物的原料并行进入密相流化床反 应器中的第 1个至第 2个二级反应区，同时预积碳催化剂依序串行通过第 1个至第 2个二级反应区， 含有含氧化合物的原料与预积碳催化剂接触， 生成包括低碳烯烃的产品和失活的待生催化剂；包括低碳烯烃的气相产品 物流和夹带的待生催化剂进入旋风分离器，气相产品物流经旋风分离器的 出口进入后续分离工段， 夹带的待生催化剂经旋风分离器的料腿进入第 2 个二级反应区； 待生催化剂由第 2个二级反应区经过汽提器、提升管进入 密相流化床再生器， 并依序串行经过第 1个至第 2个二级再生区， 与再生 介质接触， 反应后形成再生催化剂； 再生催化剂经过汽提器、 提升管再进 入密相流化床反应器， 并依序串行经过第 1个二级预积碳区、第 1个二级 反应区和第 2个二级反应区；分离工段所得的碳四以上的烃副产品返回至 密相流化床反应器中的第 1个二级预积碳区； 提升管 15中的提升气采用 汽油。密相流化床反应器反应条件为：第 1个二级预积碳区温度为 550°C， 第 1个至第 2个二级反应区温度为 450°C， 气相线速度为 0.5m/s， 床层密 度为 900kg/m3， 第 1个二级预积碳区的平均积炭量为 2wt%， 第 1个二级 反应区的平均积炭量为 6wt%， 第 2个二级反应区的平均积炭量为 8wt%； 密相流化床再生器反应条件为：反应温度为 600 °C，气相线速度为 0.7m/s， 床层密度为 700kg/m3， 第 1个二级再生区的平均积炭量为 3wt%， 第 2个 二级再生区的平均积炭量为 0.1wt%。 反应产品采用在线气相色谱分析， 低碳烯烃碳基收率为 91.2wt%。 实施例 3
密相流化床反应器内设置 1个二级预积碳区和 5个二级反应区，密相 流化床再生器内设置 5个二级再生区。石脑油和碳四以上的烃混合后进入 密相流化床反应器中的第 1个二级预积碳区， 与完全再生的催化剂接触， 转化为包括低碳烯烃的产品， 同时， 催化剂上积碳量达到一定值， 形成预 积碳催化剂， 预积碳催化剂进入反应区； 含有含氧化合物的原料并行进入 密相流化床反应器中的第 1个至第 5个二级反应区，同时预积碳催化剂依 序串行通过第 1个至第 5个二级反应区，含有含氧化合物的原料与预积碳 催化剂接触， 生成包括低碳烯烃的产品和失活的待生催化剂； 包括低碳烯 烃的气相产品物流和夹带的待生催化剂进入旋风分离器，气相产品物流经 旋风分离器的出口进入后续分离工段，夹带的待生催化剂经旋风分离器的 料腿进入第 5个二级反应区；待生催化剂由第 5个二级反应区经过汽提器、 提升管进入密相流化床再生器，并依序串行经过第 1个至第 5个二级再生 区， 与再生介质接触， 反应后形成再生催化剂； 再生催化剂经过汽提器、 提升管再进入密相流化床反应器， 并依序串行经过第 1个二级预积碳区、 第 1个二级反应区至第 5个二级反应区；分离工段所得的碳四以上的烃副 产品返回至密相流化床反应器中的第 1个二级预积碳区； 提升管 15中的 提升气采用碳四以上的烃。密相流化床反应器反应条件为： 第 1个二级预 积碳区温度为 650°C， 第 1个至第 5个二级反应区温度为 550°C， 气相线 速度为 0.7m/s，床层密度为 700kg/m3，第 1个二级预积碳区的平均积炭量 为 0.5wt%， 第 1个二级反应区的平均积炭量为 2.5wt%， 第 2个二级反应 区的平均积炭量为 4wt%， 第 3个二级反应区的平均积炭量为 5wt%， 第 4 个二级反应区的平均积炭量为 6wt%， 第 5个二级反应区的平均积炭量为 7wt%; 密相流化床再生器反应条件为： 反应温度为 700°C， 气相线速度为 1.0m/s, 床层密度为 500kg/m3， 第 1个二级再生区的平均积炭量为 5wt%， 第 2个二级再生区的平均积炭量为 3wt%， 第 3个二级再生区的平均积炭 量为 1.5wt%， 第 4个二级再生区的平均积炭量为 0.05wt%， 第 5个二级 再生区的平均积炭量为 0.01wt%。 反应产品采用在线气相色谱分析， 低碳 烯烃碳基收率为 92.5wt%。 实施例 4
密相流化床反应器内设置 2个二级预积碳区和 4个二级反应区，密相 流化床再生器内设置 4个二级再生区。碳四以上的烃并行通入密相流化床 反应器中的第 1个二级预积碳区和第 2个二级预积碳区，与完全再生的催 化剂接触， 转化为包括低碳烯烃的产品， 同时， 催化剂串行通过第 1个二 级预积碳区和第 2个二级预积碳区， 积碳量达到一定值， 形成预积碳催化 剂， 预积碳催化剂进入反应区； 含有含氧化合物的原料并行进入密相流化 床反应器中的第 1个至第 4个二级反应区，同时预积碳催化剂依序串行通 过第 1个至第 4个二级反应区，含有含氧化合物的原料与预积碳催化剂接 触， 生成包括低碳烯烃的产品和失活的待生催化剂； 包括低碳烯烃的气相 产品物流和夹带的待生催化剂进入旋风分离器，气相产品物流经旋风分离 器的出口进入后续分离工段，夹带的待生催化剂经旋风分离器的料腿进入 第 4个二级反应区； 待生催化剂由第 4个二级反应区经过汽提器、提升管 进入密相流化床再生器， 并依序串行经过第 1个至第 4个二级再生区， 与 再生介质接触， 反应后形成再生催化剂； 再生催化剂经过汽提器、 提升管 再进入密相流化床反应器， 并依序串行经过第 1个二级预积碳区、第 2个 二级预积碳区、第 1个二级反应区至第 4个二级反应区； 分离工段所得的 碳四以上的烃副产品返回至密相流化床反应器中的第 1 个二级预积碳区 和第 2个二级预积碳区； 提升管 15中的提升气采用碳四以上的烃。 密相 流化床反应器反应条件为：第 1个二级预积碳区和第 2个二级预积碳区温 度为 650°C， 第 1 个至第 4个二级反应区温度为 500°C， 气相线速度为 1.0m/s , 床层密度为 500kg/m3， 第 1 个二级预积碳区的平均积炭量为 1.5wt%, 第 2个二级预积碳区的平均积炭量为 3.0wt%， 第 1个二级反应 区的平均积炭量为 4.5wt%， 第 2个二级反应区的平均积炭量为 6.0wt%， 第 3个二级反应区的平均积炭量为 7.0wt%， 第 4个二级反应区的平均积 炭量为 8.0wt%； 密相流化床再生器反应条件为： 反应温度为 700°C， 气相 线速度为 l .Om/s,床层密度为 500kg/m3，第 1个二级再生区的平均积炭量 为 5.5wt%， 第 2个二级再生区的平均积炭量为 3wt%， 第 3个二级再生区 的平均积炭量为 1.2wt%， 第 4个二级再生区的平均积炭量为 0.02wt%。 反应产品采用在线气相色谱分析， 低碳烯烃碳基收率为 93.2wt%。 实施例 5
密相流化床反应器内设置 2个二级预积碳区和 2个二级反应区，密相 流化床再生器内设置 4个二级再生区。碳四以上的烃并行通入密相流化床 反应器中的第 1个二级预积碳区和第 2个二级预积碳区，与完全再生的催 化剂接触， 转化为包括低碳烯烃的产品， 同时， 催化剂串行通过第 1个二 级预积碳区和第 2个二级预积碳区， 积碳量达到一定值， 形成预积碳催化 剂， 预积碳催化剂进入反应区； 含有含氧化合物的原料并行进入密相流化 床反应器中的第 1个至第 2个二级反应区，同时预积碳催化剂依序串行通 过第 1个至第 2个二级反应区，含有含氧化合物的原料与预积碳催化剂接 触， 生成包括低碳烯烃的产品和失活的待生催化剂； 包括低碳烯烃的气相 产品物流和夹带的待生催化剂进入旋风分离器，气相产品物流经旋风分离 器的出口进入后续分离工段，夹带的待生催化剂经旋风分离器的料腿进入 第 2个二级反应区； 待生催化剂由第 2个二级反应区经过汽提器、提升管 进入密相流化床再生器， 并依序串行经过第 1个个至第 4个二级再生区， 与再生介质接触， 反应后形成再生催化剂； 再生催化剂经过汽提器、提升 管再进入密相流化床反应器， 并依序串行经过第 1个二级预积碳区、 第 2 个二级预积碳区、第 1个二级反应区、 第 2个二级反应区； 分离工段所得 的碳四以上的烃副产品返回至密相流化床反应器中的第 1 个二级预积碳 区和第 2个二级预积碳区； 提升管 15中的提升气采用碳四以上的烃。 密 相流化床反应器反应条件为：第 1个二级预积碳区和第 2个二级预积碳区 温度为 650°C， 第 1个至第 2个二级反应区温度为 500°C， 气相线速度为 1.0m/s , 床层密度为 500kg/m3， 第 1 个二级预积碳区的平均积炭量为 1.5wt%, 第 2个二级预积碳区的平均积炭量为 3.0wt%， 第 1个二级反应 区的平均积炭量为 6.0wt%， 第 2个二级反应区的平均积炭量为 8.5wt%； 密相流化床再生器反应条件为：反应温度为 700°C，气相线速度为 l .Om/s, 床层密度为 500kg/m3， 第 1个二级再生区的平均积炭量为 5.8wt%， 第 2 个二级再生区的平均积炭量为 3wt%， 第 3个二级再生区的平均积炭量为 l .lwt%, 第 4个二级再生区的平均积炭量为 0.02wt%。反应产品采用在线 气相色谱分析， 低碳烯烃碳基收率为 92.8wt%。
以上已对本发明进行了详细描述，但本发明并不局限于本文所描述具 体实施方式。本领域技术人员理解， 在不背离本发明范围的情况下， 可以 作出其他更改和变形。 本发明的范围由所附权利要求限定。

Claims

权 利 要 求
1. 一种含氧化合物制低碳烯烃的方法， 包括以下歩骤：
a) 将碳四以上的烃从 k个预积碳区进料支线并行通入密相流化床反应 器中的 k个二级预积碳区， 与再生完全和 /或新鲜的催化剂接触， 转化为 包含低碳烯烃产品的物流， 同时， 形成预积碳催化剂； 其中所述催化剂依 序串行通过第 1个至第 k个二级预积碳区， 碳含量逐渐增加； 其中所述密 相流化床反应器由物料流动控制器分隔为预积碳区和反应区；其中所述密 相流化床反应器的预积碳区由物料流动控制器分隔为 k个二级预积碳区， 第 1个至第 k个二级预积碳区依序相联；
b) 将含有含氧化合物的原料从 n个反应区进料支线并行通入密相流化 床反应器中的 n个二级反应区， 与所述预积碳催化剂接触， 生成包含低碳 烯烃产品的物流和待生催化剂；其中所述从第 k个二级预积碳区流入的预 积碳催化剂依序串行通过第 1个至第 n个二级反应区， 碳含量逐渐增加； 其中所述密相流化床反应器的反应区由物料流动控制器分隔为 n个二级 反应区， 第 1个至第 n个二级反应区依序相联， 第 1个二级反应区连接于 第 k个二级预积碳区的下游；
c) 将由所述预积碳区和反应区流出的所述包含低碳烯烃产品的物流 与其所携带的待生催化剂分离；分离出的待生催化剂进入第 n个二级反应 区； 所述包含低碳烯烃产品的物流进入产品分离工段， 经分离、提纯得到 低碳烯烃产品，分离工段所得的碳四以上的烃副产品返回至所述密相流化 床反应器中的预积碳区；
d) 由第 n个二级反应区流出的待生催化剂经过汽提、提升进入密相流 化床再生器再生； 所述待生催化剂依序串行通过第 1个至第 m个二级再 生区； 再生介质从 m个再生区进料支线并行通入第 1个至第 m个二级再 生区， 所述待生催化剂与所述再生介质接触， 碳含量逐渐下降， 完成再生 后的催化剂随后经汽提、提升返回第 1个二级预积碳区； 其中所述密相流 化床再生器由物料流动控制器分隔为 m个二级再生区， 第 1个至第 n个 二级再生区依序相联；
其中， k l， n^ l , m^2 o
2. 根据权利要求 1所述的方法， 其特征在于， 4 k≥2， 8^n^3 , 8 ^
3. 根据权利要求 1所述的方法， 其特征在于， 在所述密相流化床反应 器中， 物料流动控制器中气体表观线速度小于等于催化剂的最小流化速 度。
4. 根据权利要求 1所述的方法， 其特征在于， 在所述密相流化床再生 器中， 物料流动控制器中气体表观线速度小于等于催化剂的最小流化速 度。
5. 根据权利要求 1所述的方法，其特征在于，所述催化剂含有 SAPO-34 分子筛。
6. 根据权利要求 1所述的方法， 其特征在于， 所述密相流化床反应器 的反应条件为： 所述预积碳区和反应区的气体表观线速度为 0.1-1.5m/s， 所述预积碳区的反应温度为 500-650°C， 所述反应区的温度为 400-550°C， 所述密相流化床反应器的床层密度为 200-1200kg/m3。
7. 根据权利要求 1所述的方法， 其特征在于， 所述密相流化床反应器 中第 1个二级预积碳区至第 n个二级反应区内的催化剂平均积炭量依次递 增， 第 k个二级预积碳区内的催化剂平均积炭量为 0.5-3wt%， 第 n个二 级反应区内的催化剂平均积炭量为 7-10wt%。
8. 根据权利要求 1所述的方法， 其特征在于， 所述密相流化床再生区 的反应条件为： 气体表观线速度为 0.1-1.5m/s， 反应温度为 500-700 °C， 床 层密度为 200-1200kg/m3。
9. 根据权利要求 1所述的方法， 其特征在于， 所述密相流化床再生区 第 1个至第 m个二级再生区内催化剂平均积炭量依次递减， 第 1个二级 再生区内的催化剂平均积炭量为 2-10wt%， 第 m个二级再生区内的催化 剂平均积炭量为 0-0.1wt%。
10. 根据权利要求 1 所述的方法， 其特征在于， 所述含氧化合物为甲 醇和 /或二甲醚； 所述低碳烯烃为乙烯、 丙烯或丁烯中的任意一种或任意 几种的混合物； 所述碳四以上的烃为石脑油、 汽油、 凝析油、 轻柴油、 加 氢尾油或煤油中的任意一种或任意几种的混合物； 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。
11. 根据权利要求 1所述的方法， 其特征在于， 歩骤 d)中完成再生后 的催化剂随后经汽提、提升返回密相流化床的第 1个二级预积碳区， 所述 提升过程中的提升气采用水蒸气、碳四以上的烃、石脑油、汽油、凝析油、 轻柴油、 加氢尾油或煤油中的任意一种或任意几种的混合物。
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* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN109694298A (zh) * 2017-10-20 2019-04-30 中国石油化工股份有限公司 甲醇转化制烃的方法
JP2020513468A (ja) * 2016-12-14 2020-05-14 エクソンモービル リサーチ アンド エンジニアリング カンパニーＥｘｘｏｎ Ｒｅｓｅａｒｃｈ Ａｎｄ Ｅｎｇｉｎｅｅｒｉｎｇ Ｃｏｍｐａｎｙ 流動接触分解装置におけるオキシジェネート転化のための方法
WO2022077457A1 (zh) * 2020-10-16 2022-04-21 中国科学院大连化学物理研究所 焦调控反应器、装置及含氧化合物制备低碳烯烃的方法
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EP4082655A4 (en) * 2020-10-16 2023-04-26 Dalian Institute Of Chemical Physics, Chinese Academy Of Sciences FLUIDIZED BED REGENERATOR, APPARATUS FOR PREPARING LOW CARBON OLEFIN AND APPLICATION THEREOF
EP4082658A4 (en) * 2020-10-16 2023-04-26 Dalian Institute Of Chemical Physics, Chinese Academy Of Sciences COKE CONTROL REACTOR, DEVICE FOR THE PREPARATION OF LOW CARBON OLEFIN FROM AN OXYGEN-CONTAINING COMPOUND AND USE THEREOF
RU2798851C1 (ru) * 2020-10-16 2023-06-28 Далянь Инститьют Оф Кемикал Физикс, Чайниз Академи Оф Сайенсез Реактор для управления содержанием кокса, а также устройство и способ получения низкоуглеродистых олефинов из кислородсодержащего соединения

Families Citing this family (8)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US9827544B2 (en) * 2013-12-03 2017-11-28 Dalian Institute Of Chemical Physics, Chinese Academy Of Sciences Reaction device for preparing light olefins from methanol and/or dimethyl ether
AU2013407175B2 (en) * 2013-12-03 2017-06-15 Dalian Institute Of Chemical Physics, Chinese Academy Of Sciences Method for preparing a light olefin using an oxygen-containing compound, and device for use thereof
AU2016426746B2 (en) * 2016-10-19 2020-02-27 Dalian Institute Of Chemical Physics, Chinese Academy Of Sciences Turbulent fluidized-bed reactor, device, and method using oxygen-containing compound for manufacturing propene and C4 hydrocarbon
EP3530641B1 (en) * 2016-10-19 2024-01-24 Dalian Institute Of Chemical Physics, Chinese Academy of Sciences Method and device for manufacturing propene and c4 hydrocarbon
CN114364454B (zh) * 2019-07-31 2023-11-07 沙特基础工业全球技术公司 用于最大化btx产率的密相流化床反应器
CN113387762A (zh) * 2020-03-13 2021-09-14 中国石油化工股份有限公司 在含要求温度控制的流化床反应器内转化甲醇的方法
CN111875464B (zh) * 2020-07-10 2023-11-03 中石化洛阳工程有限公司 一种高效的含氧化合物生产低碳烯烃的方法
KR20230013253A (ko) * 2020-10-16 2023-01-26 달리안 인스티튜트 오브 케미컬 피직스, 차이니즈 아카데미 오브 사이언시즈 재생장치, 저탄소 올레핀을 제조하는 장치 및 그 응용

Citations (7)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US6166282A (en) 1999-08-20 2000-12-26 Uop Llc Fast-fluidized bed reactor for MTO process
CN102276406A (zh) 2010-06-11 2011-12-14 中国石油化工股份有限公司 增产丙烯生产方法
CN102463138A (zh) * 2010-11-17 2012-05-23 中国石油化工股份有限公司 Sapo-34催化剂的两段再生方法
CN102816591A (zh) * 2011-06-09 2012-12-12 中国石油化工股份有限公司 一种催化裂化方法
CN101402538B (zh) 2008-11-21 2013-01-09 中国石油化工股份有限公司 提高低碳烯烃收率的方法
CN102875289A (zh) 2011-07-12 2013-01-16 中国石油化工股份有限公司 制备低碳烯烃的方法
CN102875296A (zh) 2011-07-12 2013-01-16 中国石油化工股份有限公司 甲醇制低碳烯烃的反应装置

Family Cites Families (10)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US7057083B2 (en) 2003-11-12 2006-06-06 Exxonmobil Chemical Patents Inc. Catalyst pretreatment with C4-C7 olefins in an oxygenate to olefins reaction system
US7465845B2 (en) * 2004-12-22 2008-12-16 Exxonmobil Chemical Patents Inc. Increasing ethylene and/or propylene production in an oxygenate to olefins reaction systems
US7763765B2 (en) 2006-03-31 2010-07-27 Exxonmobil Chemical Patents Inc. Method of high pressure and high capacity oxygenate conversion with catalyst exposure cycle
CN101239868B (zh) * 2007-02-07 2011-05-18 中国石油化工股份有限公司 提高乙烯、丙烯收率的方法
CN101348404B (zh) * 2007-07-18 2011-11-30 中国石油化工股份有限公司 甲醇或二甲醚转化过程中提高乙烯、丙烯收率的方法
CN101362668B (zh) * 2007-11-15 2011-09-14 陕西煤化工技术工程中心有限公司 一种制取丙烯的方法
EP2300396B1 (en) 2008-06-25 2019-08-07 Total Research & Technology Feluy Process to make olefins from oxygenates
CN102463183A (zh) * 2010-11-19 2012-05-23 江苏中油天工机械有限公司 一种破碎机刀盘轴固定框
US20140310980A1 (en) 2011-05-12 2014-10-23 Glatt Ingenieurtechnik Gmbh Device for the continuous treatment of solids in a fluidized bed apparatus
AU2013407175B2 (en) * 2013-12-03 2017-06-15 Dalian Institute Of Chemical Physics, Chinese Academy Of Sciences Method for preparing a light olefin using an oxygen-containing compound, and device for use thereof

Patent Citations (7)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US6166282A (en) 1999-08-20 2000-12-26 Uop Llc Fast-fluidized bed reactor for MTO process
CN101402538B (zh) 2008-11-21 2013-01-09 中国石油化工股份有限公司 提高低碳烯烃收率的方法
CN102276406A (zh) 2010-06-11 2011-12-14 中国石油化工股份有限公司 增产丙烯生产方法
CN102463138A (zh) * 2010-11-17 2012-05-23 中国石油化工股份有限公司 Sapo-34催化剂的两段再生方法
CN102816591A (zh) * 2011-06-09 2012-12-12 中国石油化工股份有限公司 一种催化裂化方法
CN102875289A (zh) 2011-07-12 2013-01-16 中国石油化工股份有限公司 制备低碳烯烃的方法
CN102875296A (zh) 2011-07-12 2013-01-16 中国石油化工股份有限公司 甲醇制低碳烯烃的反应装置

Cited By (22)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
JP2020513468A (ja) * 2016-12-14 2020-05-14 エクソンモービル リサーチ アンド エンジニアリング カンパニーＥｘｘｏｎ Ｒｅｓｅａｒｃｈ Ａｎｄ Ｅｎｇｉｎｅｅｒｉｎｇ Ｃｏｍｐａｎｙ 流動接触分解装置におけるオキシジェネート転化のための方法
JP7039595B2 (ja) 2016-12-14 2022-03-22 エクソンモービル リサーチ アンド エンジニアリング カンパニー 流動接触分解装置におけるオキシジェネート転化のための方法
CN109694298A (zh) * 2017-10-20 2019-04-30 中国石油化工股份有限公司 甲醇转化制烃的方法
CN109694298B (zh) * 2017-10-20 2021-12-28 中国石油化工股份有限公司 甲醇转化制烃的方法
EP4082655A4 (en) * 2020-10-16 2023-04-26 Dalian Institute Of Chemical Physics, Chinese Academy Of Sciences FLUIDIZED BED REGENERATOR, APPARATUS FOR PREPARING LOW CARBON OLEFIN AND APPLICATION THEREOF
JP2023528819A (ja) * 2020-10-16 2023-07-06 中国科学院大▲連▼化学物理研究所 流動床反応器、装置及び応用
US20230001370A1 (en) * 2020-10-16 2023-01-05 Dalian Institute Of Chemical Physics, Chinese Academy Of Sciences Coke control reactor, and device and method for preparing low-carbon olefins from oxygen-containing compound
US20230001371A1 (en) * 2020-10-16 2023-01-05 Dalian Institute Of Chemical Physics, Chinese Academy Of Sciences Fluidized bed reactor, device, and use thereof
EP4082654A4 (en) * 2020-10-16 2023-01-25 Dalian Institute Of Chemical Physics, Chinese Academy Of Sciences FLUID BED REACTOR, APPARATUS AND PROCESS FOR PRODUCTION OF LOW CARBON OLEFINS FROM OXYGEN CONTAINING COMPOUNDS
EP4088811A4 (en) * 2020-10-16 2023-04-12 Dalian Institute Of Chemical Physics, Chinese Academy Of Sciences FLUID BED REACTOR, DEVICE AND APPLICATION
WO2022077457A1 (zh) * 2020-10-16 2022-04-21 中国科学院大连化学物理研究所 焦调控反应器、装置及含氧化合物制备低碳烯烃的方法
EP4082658A4 (en) * 2020-10-16 2023-04-26 Dalian Institute Of Chemical Physics, Chinese Academy Of Sciences COKE CONTROL REACTOR, DEVICE FOR THE PREPARATION OF LOW CARBON OLEFIN FROM AN OXYGEN-CONTAINING COMPOUND AND USE THEREOF
RU2798851C1 (ru) * 2020-10-16 2023-06-28 Далянь Инститьют Оф Кемикал Физикс, Чайниз Академи Оф Сайенсез Реактор для управления содержанием кокса, а также устройство и способ получения низкоуглеродистых олефинов из кислородсодержащего соединения
WO2022077452A1 (zh) * 2020-10-16 2022-04-21 中国科学院大连化学物理研究所 流化床反应器、装置以及含氧化合物制备低碳烯烃的方法
JP2023528818A (ja) * 2020-10-16 2023-07-06 中国科学院大▲連▼化学物理研究所 コークス制御反応器、装置及び含酸素化合物から軽オレフィンを製造する方法
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