WO2012167708A1 - 由含氧化合物制烯烃的流化床反应器和方法 - Google Patents

由含氧化合物制烯烃的流化床反应器和方法 Download PDF

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WO2012167708A1
WO2012167708A1 PCT/CN2012/076249 CN2012076249W WO2012167708A1 WO 2012167708 A1 WO2012167708 A1 WO 2012167708A1 CN 2012076249 W CN2012076249 W CN 2012076249W WO 2012167708 A1 WO2012167708 A1 WO 2012167708A1
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catalyst
mpa mpa
gas
fluidized bed
dense phase
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PCT/CN2012/076249
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English (en)
French (fr)
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魏飞
魏小波
王垚
朱畅
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富德（北京）能源化工有限公司
清华大学
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Priority to RU2013157347/05A priority Critical patent/RU2561985C2/ru
Priority to CA2859723A priority patent/CA2859723C/en
Priority to AU2012267002A priority patent/AU2012267002B2/en
Priority to US14/124,502 priority patent/US9403735B2/en
Priority to EP12797214.9A priority patent/EP2719451B8/en
Publication of WO2012167708A1 publication Critical patent/WO2012167708A1/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
    • 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
    • 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
    • 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
    • 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
    • B01J8/005—Separating solid material from the gas/liquid stream
    • B01J8/0055—Separating solid material from the gas/liquid stream using cyclones
    • 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
    • B01J8/1836—Heating and cooling the reactor
    • 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
    • B01J8/1845—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
    • 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
    • C07C2529/00—Catalysts comprising molecular sieves
    • C07C2529/82—Phosphates
    • C07C2529/84—Aluminophosphates containing other elements, e.g. metals, boron
    • C07C2529/85—Silicoaluminophosphates (SAPO compounds)
    • Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
    • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
    • 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 fluidized bed reactor for the production of olefins from oxygenates and a process for the production of olefins from oxygenates using the fluidized bed reactors.
  • a silicoaluminophosphate (SAP0) type molecular sieve catalyst can be used to convert a low carbon oxide such as methanol and / dimethyl ether or the like into a low carbon olefin such as ethylene, propylene, butylene or the like.
  • SAP0-type molecular sieve catalysts such as SAP0-5, SAP0-11, SAP0-17, SAP0-41, SAP0-34, and SAP0-41 have been developed for the conversion of oxygenates to olefins, and It is recognized that SAP0-34 has excellent catalytic performance when used for methanol and/or dimethyl ether to produce olefin (MT0) because the molecular sieve catalyst has a small pore size and good hydrothermal stability.
  • olefins from oxygenates is carried out primarily by catalytic cracking, which is generally an exothermic reaction.
  • the target products are ethylene and propylene, but a small amount of butene, pentene, hexene and corresponding alkanes are formed during the reaction.
  • various product olefins may undergo secondary reactions such as interconversion, for example, ethylene and/or propylene may be further oligomerized to C 4 or higher.
  • CN1166478A discloses a process for producing a lower olefin such as ethylene or propylene from methanol or dimethyl ether, wherein the reaction and regeneration are continuously carried out in a dense phase circulating fluidized bed reactor using a SAP0-34 molecular sieve catalyst
  • US 4547616 discloses A continuous process for the production of lower olefins from oxygenates using a turbulent fluidized bed wherein the turbulent fluidized bed is also a dense phase fluidized bed reactor
  • US6023005 discloses the use of molecular sieve catalysts for oxygenation.
  • a process for the conversion of a compound to an olefin wherein a riser reactor is used.
  • No. 6,166,282 discloses a fast fluidized bed reactor for an NOx process, the reactor comprising an upper separation zone and a lower reaction zone, wherein the reaction zone comprises a dense phase zone and a transition zone above the dense phase zone, wherein The reaction mixture after the reaction in the dense phase zone is completely converted in the transition zone.
  • the fast fluidized bed significantly reduces the size of the reactor and reduces the catalyst inventory, thereby saving costs, but there is still a problem of gas-solid back mixing, and the gas entering the settling chamber needs to be longer. Time enters the cyclone, so this part of the gas still undergoes a secondary reaction, which is not conducive to ethylene and The total selectivity of propylene.
  • the present invention provides a fluidized bed reactor for the production of olefins from oxygenates and a process for the production of olefins from oxygenates using the fluidized bed reactors.
  • the invention provides a fluidized bed reactor for the production of an olefin from an oxygenate, the fluidized bed reactor comprising:
  • reaction zone located at a lower portion of the fluidized bed reactor, comprising a lower dense phase zone and an upper riser, wherein the dense phase zone and the riser are transitionally connected;
  • a separation zone the separation zone being located at an upper portion of the fluidized bed reactor, comprising a settling chamber, a gas-solid quick-distribution device, a cyclone separator, and a plenum, the riser extending upward into the separation zone and
  • An outlet is in communication with an inlet of the gas-solid quick-distribution device, an outlet of the gas-solid quick-distribution device is in communication with an inlet of the cyclone through a rapid air duct, and an outlet of the cyclone is connected to a gas collection chamber
  • the gas collection chamber is disposed below and connected to the outlet of the reactor;
  • a catalyst recycle line from the settling chamber to the dense phase zone, and a deactivated catalyst derivation line from the settling chamber and/or the dense phase zone to the catalyst regeneration unit and from the catalyst regeneration unit to the dense phase
  • the regenerated catalyst of the zone is returned to the pipeline.
  • the invention also provides a process for the preparation of an olefin from an oxygenate, the process utilizing the fluidized bed reactor described above, and comprising the steps of:
  • reaction mixture containing the low carbon olefin, the unconverted feedstock and the catalyst is passed from the dense phase zone into the riser where the unconverted feedstock continues to react under the action of the catalyst and is substantially completely converted, resulting in an upgrade comprising primarily low carbon olefins and catalysts.
  • the riser effluent directly enters the separation zone located in the upper part of the fluidized bed reactor, first enters the gas-solid quick-distribution device, where most of the catalyst is separated, and then enters the cyclone through the fast air-conducting pipe, where the remaining is further separated After the catalyst enters the plenum and then exits through the outlet of the reactor, the separated catalyst falls into the settling chamber located at the lower portion of the separation zone;
  • a portion of the catalyst is withdrawn from the settling chamber via the catalyst recycle line and recycled back to the dense phase region, and a portion of the catalyst is withdrawn from the settling chamber and/or the dense phase region via the deactivated catalyst derivation line into the catalyst regeneration unit for regeneration.
  • a portion of the regenerated catalyst is then withdrawn from the catalyst regeneration unit via the regenerated catalyst return line and returned to the dense phase zone.
  • an inner member may be disposed in the dense phase region of the fluidized bed reactor, and the inner member may be used to remove heat of reaction and/or inhibit gas-solid back mixing, for example, the inner member may
  • the inner member may be used to remove heat of reaction and/or inhibit gas-solid back mixing
  • the inner member may
  • heat exchange members such as heat exchangers or heat exchange coils
  • back-mixing components such as baffles or baffles, and even in some cases, for the simplicity and efficiency of the design, it is possible to combine these components. Better control of reaction temperature and / or higher conversion of raw materials.
  • the singularity is 0. 6-5, more preferably 0. 6-5, more preferably 0. 8-5, the above-mentioned 5-8 ⁇
  • the height ratio of the diameter of the riser is 2-20, preferably 3-15, more preferably 4-10, and the diameter ratio of the dense phase zone to the riser is 2-10, preferably 2. 5-8 More preferably 3-6; 1 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 2 0. 2- 1. 5m / s, more preferably 0. 3- 1. 2m / s, and a residence time of 0.
  • the singular gas velocity in the riser is from 2 to 20 m/s, preferably from 4 to 18 m/s, more preferably from 5 to 15 m/s, and the residence time is from 0. 3-5 s, preferably from 0.4 to 4 s, more Preferably, it is 0.5-3s, so that the load of the reaction zone can be reasonably distributed to achieve the target conversion rate.
  • the settling chamber of the fluidized bed reactor is located at a lower portion of the separation zone for collecting the catalyst separated in the separation zone, and a vapor stripping gas distributor may be provided in the settling chamber for introducing the stripping gas.
  • the stripping gas may be nitrogen or steam, preferably steam, and the stripping operation may further separate the raw material gas and/or product gas entrained in the catalyst, and in the settling chamber
  • a stripping mixture withdrawal line may be provided above for extracting the stripped gas mixture via the cyclone.
  • the gas-solid quick-distribution device in the fluidized bed reactor may be any suitable gas-solid fast-distribution device commonly used in the art.
  • the gas-solid quick-distribution device may be selected from a coarse-rotating gas.
  • Solid-speed sub-assembly, ejection gas-solid quick-distribution device, inverted L-type gas-solid quick-distribution device, T-type gas-solid quick-distribution device, multi-leaf gas-solid quick-distribution device and swirling arm gas-solid quick-distribution device The gas-solid two phases in the riser effluent can be quickly separated.
  • the cyclone separator in the fluidized bed reactor may be any suitable cyclone separator commonly used in the art, and the cyclone separator may be one or more sets of cyclones, and each group
  • the cyclone separator can be formed by connecting one, two or three cyclones in series to further separate the solid catalyst entrained in the product stream.
  • the rapid air duct in the fluidized bed reactor makes the gas-solid
  • the quick-distribution device is in communication with the cyclone separator and is sized such that the residence time of the stream within it is as short as possible, for example, such that the residence time generally does not exceed 5 s, preferably does not exceed 4 s, more preferably does not exceed 3 s.
  • reaction mixture leaving the fluidized bed reactor into a product aftertreatment unit for separation and/or purification to obtain a final product olefin such as ethylene and/or propylene, said separation and/or purification operation.
  • a product aftertreatment unit for separation and/or purification to obtain a final product olefin such as ethylene and/or propylene, said separation and/or purification operation.
  • This can be carried out by any suitable means known in the art such as a rectification column and an absorption column.
  • the oxygenated compound to be converted may be a lower alcohol and/or ether material which is commonly used in the art.
  • the oxygenate may be selected from the group consisting of methanol, ethanol, propanol, dimethyl ether, diethyl ether, and The propyl ether and mixtures thereof, preferably methanol and/or dimethyl ether; and the lower olefin product may be selected from the group consisting of ethylene, propylene, butene and mixtures thereof, preferably ethylene and/or propylene.
  • a diluent is added to the raw material in order to reduce the partial pressure of the raw material and the corresponding reaction product
  • the diluent may be any suitable gas inert to the reaction of converting the oxygenate into an olefin, for example, It may be steam or nitrogen, preferably steam; here, the diluent may be added in an amount of 5 to 80 mol%, preferably 10 to 60 mol%, more preferably 15 to 50 mol%, of the raw material mixture.
  • the catalyst for converting an oxygenate to an olefin is generally
  • the silicoaluminophosphate-based molecular sieve catalyst for example, the catalyst may be selected from the group consisting of SAP0-5, SAP0-11, SAP0-17, SAP0-41, SAP0-34, and SAP0-41, preferably SAP0-34.
  • reaction temperature at which the oxygen-containing compound is reacted on the SAP0 type molecular sieve catalyst those skilled in the art can accordingly select according to the oxygen-containing compound to be used, and generally, the oxygen-containing compound is converted into
  • the reaction temperature of the olefin is from 300 to preferably from 400 to 550.
  • the reaction pressure for reacting the oxygenate on the SAP0 type molecular sieve catalyst can also be considered by those skilled in the art depending on the oxygenate used and the design and operation cost of the fluidized bed reactor. 5 ⁇
  • the pressure is 0. 1-0. 5MPa absolute pressure.
  • the pressure is 0. 1-0. 5MPa absolute pressure.
  • an oxygen-containing gas such as air, oxygen-enriched air or pure oxygen, preferably air, is used for deactivating the catalyst by reacting the oxygenate on a SAP0 type molecular sieve catalyst to cause carbon deposition.
  • a SAP0 type molecular sieve catalyst 0-1 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0
  • the fluidized bed reactor combines a dense phase zone and a riser in the reaction zone, wherein a majority of the feedstock stream is converted in the dense phase zone, and the remaining feedstock stream is further converted in the riser pipe, according to which Under the premise of determining the target conversion rate of the raw material stream, the introduction of the riser reduces the catalyst storage load in the dense phase region, and the catalyst storage and catalyst average can be further adjusted through the catalyst circulation and regeneration operation.
  • the activity in turn, further adjusts the depth of reaction of the feed stream and thereby optimizes the conversion of the feed stream.
  • the separation zone of the fluidized bed reactor combines a gas-solid quick-distribution device and a cyclone separator, wherein the gas-solid fast-distribution device rapidly separates most of the catalyst, The remaining catalyst is then further separated by a cyclone, wherein the gas-solid fast-distribution device and the cyclone are connected by means of a rapid gas conduit, thereby minimizing the gas-solid separation time and correspondingly inhibiting and/or reducing the product olefin.
  • the second reaction therefore, increases the selectivity of the target product.
  • the fluidized bed reactor of the present invention combines a dense phase zone and a riser to control the reaction depth and optimize the conversion rate, and further combines the gas-solid quick-distribution device and the cyclone separator to achieve rapid gas-solid separation and suppress secondary The reaction ensures the selectivity of the desired product olefin.
  • the fluidized bed reactor of the present invention can adjust and/or control the total conversion and total selectivity of the conversion of oxygenates to olefins, thereby achieving an optimum reaction effect, thereby achieving the object of the present invention.
  • FIG. 1 is a specific example of a fluidized bed reactor of the present invention, wherein a dense phase zone and a riser are connected by a conical pipe section, and the gas-solid quick-distribution device adopts a coarse-rotation gas-solid quick-distribution device, a cyclone separator A secondary cyclone separator in series with two cyclones, and wherein a portion of the deactivated catalyst is withdrawn from the settling chamber into the catalyst regeneration unit for regeneration.
  • Fig. 1 The specific meanings of the reference numerals in Fig. 1 are as follows: 1-raw material inlet distributor; 2 - dense phase zone; 3-internal component; 4-lift pipe; 5 rotary gas-solid quick-distribution device; 6 rotary gas-solid Fast-distributing device material legs; 7-fast air guiding tube; 8-cyclone separator; 9-cyclone separator material leg; 10-collection chamber; 11-reactor outlet; 12-product post-treatment device; 13-steam stripping gas Distributor; 14-settling chamber; 15- stripping mixture derivation line; 16-deactivated catalyst derivation line; 17-deactivated catalyst flow regulating valve; 18-catalyst regenerating unit; 19-regenerated catalyst return line; 20-regenerated catalyst Return flow regulating valve; 21 - catalyst circulation line; 22 ⁇ flL chemical circulating flow regulating valve. detailed description
  • a feedstock stream containing an oxygenate such as methanol and/or dimethyl ether and a diluent such as steam is passed through a feed inlet distributor 1 into a dense phase zone 2 in the lower portion of the fluidized bed reactor where it is held for a certain period of time.
  • 4- 10s is contacted with the catalyst and reacts to convert most of the raw materials into olefins, and at the same time, some of the catalysts are deactivated due to carbon deposition, and the heat released by the reaction is removed by the internal member 3 to keep the reaction temperature within a certain range, such as 400-550. , and the reaction pressure is maintained at 0. 1-0. 5MPa absolute pressure, the inner member 3 simultaneously suppresses gas-solid back mixing;
  • the reaction mixture containing the olefin, the unconverted raw material and the catalyst is advanced from the dense phase zone 2 into the riser 4, and is held therein for a certain period of time, such as 1 - 3 s, so that the unconverted raw material continues to react under the action of the catalyst and is substantially completely converted, resulting in a main a riser effluent containing a low carbon olefin and a catalyst;
  • the riser effluent first enters the coarse-rotating gas-solid quick-distribution device 5, after which most of the catalyst is separated, and then the short-term residence of the rapid air-guiding pipe 7 does not exceed 3 s, and then enters the cyclone separator 8, after further separating the remaining catalyst. Entering the plenum 10, then exiting through the reactor outlet 11, and then going to the product aftertreatment unit 12 for separation to obtain the final product olefin, and the separated catalysts are respectively passed through the legs 6 of the coarse-rotation gas-solid quick-distribution device 5. And the material leg 9 of the cyclone separator 8 falls into the settling chamber 14 at the lower part of the separation zone;
  • a portion of the catalyst is withdrawn from the settling chamber 14 and returned to the dense phase zone 2 via the catalyst recycle line 21, and a portion of the catalyst is withdrawn from the settling chamber 14 via the deactivated catalyst derivation line 16 into the catalyst regeneration unit 18 at a temperature of 500-750 and 0.1. -0. 5MPa under pressure to burn off carbon deposits for regeneration, and then regenerated catalyst
  • the return line 19 leads a part of the regenerated catalyst from the catalyst regeneration unit 18 to return to the dense phase zone 2, where the circulating flow rate of the catalyst, the deducted flow rate of the deactivated catalyst, and the return flow rate of the regenerated catalyst are adjusted by the valves on the respective lines, thereby making the dense phase
  • the catalyst activity in zone 2 is more suitable for the conversion reaction;
  • the fluidized bed reactor shown in Fig. 1 was used, and a mixture of methanol and steam was used as a raw material, steam as a diluent accounted for 30 mol% of the mixture, and the total flow rate of the raw materials was 67 kmol/h; the catalyst was SAP0-34, and the loading amount was 0. 5t; keep the reaction temperature at 485, the reaction pressure is 0. 2 MPa absolute pressure, the apparent gas velocity in the dense phase region is 0.
  • the conversion reaction of methanol was carried out in the same manner as in Example 1, except that conventional bubbling was carried out.
  • a fluidized bed reactor in which a dense phase zone is used as a reaction zone, and a separation zone is a three-stage cyclone separator in which three cyclones are connected in series, wherein the loading amount of the catalyst is
  • Example 1 shows that the fluidized bed reactor of the present invention can achieve higher conversion rates and higher at significantly reduced catalyst inventory compared to conventional bubbling bed reactors.
  • the total selectivity of ethylene + propylene indicates that significant improvements have been made in the fluidized bed reactor of the present invention.

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  • Chemical & Material Sciences (AREA)
  • Organic Chemistry (AREA)
  • Chemical Kinetics & Catalysis (AREA)
  • Engineering & Computer Science (AREA)
  • Combustion & Propulsion (AREA)
  • Devices And Processes Conducted In The Presence Of Fluids And Solid Particles (AREA)
  • Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)

Abstract

一种流化床反应器以及使用所述反应器由含氧化合物制备烯烃的方法，所述流化床反应器包括：反应区，所述反应区位于所述流化床反应器的下部，包括下部的密相区（2）和上部的提升管（4），其中所述密相区（2）和所述提升管（4）之间过渡连接；分离区，所述分离区位于所述流化床反应器的上部，包括沉降室（14）、气固快分装置（5）、旋风分离器（8）和集气室（10），所述提升管（4）向上延伸进入所述分离区和在其出口处与所述气固快分装置（5）的入口连通，所述气固快分装置（5）的出口经快速导气管（7）与所述旋风分离器（8）的入口连通，所述旋风分离器（8）的出口与集气室（10）连通，所述集气室（10）设在反应器出口下方并与之连通；和从所述沉降室（14）至所述密相区（2）的催化剂循环管线（21），以及从所述沉降室（14）和/或所述密相区（2）至催化剂再生装置（18）的失活催化剂导出管线（16）和从催化剂再生装置（18）至所述密相区（2）的再生催化剂返回管线（19）。

Description

由含氧化合物制烯烃的流化床反应器和方法 技术领域
本发明涉及由含氧化合物制烯烃的流化床反应器以及利用所 述流化床反应器由含氧化合物制烯烃的方法。 背景技术
已知可利用硅铝磷酸盐（SAP0)类分子筛催化剂将低碳含氧化 合物如甲醇和 /二甲醚等转化成低碳烯烃如乙烯、 丙烯和丁烯等。 迄今为止， 已经开发研制了一系列的 SAP0 类分子筛催化剂例如 SAP0- 5、 SAP0- 11、 SAP0- 17、 SAP0- 41、 SAP0- 34和 SAP0- 41等用 于将含氧化合物转化成烯烃， 并且其中公认的是 SAP0- 34在用于 甲醇和 /或二甲醚制烯烃（MT0)时具有优异的催化性能， 这是因为 该分子筛催化剂孔径较小且水热稳定性很好。
由含氧化合物制烯烃主要通过催化裂解进行， 该过程一般为 放热反应。 具体地， 对于由甲醇和 /或二甲醚制乙烯、 丙烯等低碳 烯烃来说， 目标产物是乙烯和丙烯， 但反应过程中还生成少量丁 烯、 戊烯、 己烯和相应的烷烃， 这是由于在反应过程中， 除了甲 醇和 /或二甲醚催化裂解成烯烃外，各种产物烯烃还可能发生相互 转化等二次反应， 例如乙烯和 /或丙烯可进一步低聚成 C4以上的 高碳烯烃。
因此， 为了提高过程的乙烯和丙烯产量， 不仅要提高过程的 总转化率使反应物尽可能完全转化掉， 还要提高过程对乙烯和 / 或丙烯的总选择性。据此，对于确定的 SAP0类分子筛催化剂来说， 既要保证反应物与催化剂的充分接触以使反应物充分转化， 还要 尽量避免或减少产物气体与催化剂的接触来相应避免或减少乙烯 和 /或丙烯低聚成高碳烯烃等二次反应。
对于上述由含氧化合物制烯烃如 MT0的反应过程， 现有技术 中开发应用了各种反应器， 其中主要包括密相流化床反应器和提 升管反应器。 例如， CN1166478A公开了一种由甲醇或二甲醚制乙 烯、 丙烯等低碳烯烃的方法， 其中使用 SAP0- 34分子筛催化剂在 密相循环流化床反应装置中连续进行反应和再生， US4547616 则 公开了一种利用湍动流化床进行含氧化合物制低碳烯烃的连续方 法，其中的湍动流化床也是一种密相流化床反应器；而 US6023005 则公开了使用分子筛催化剂将含氧化合物转化成烯烃的方法， 其 中使用的是提升管反应器。
对于密相流化床反应器来说， 可在床层中加入移热设备， 使 反应温度易于控制， 但由于密相区内气固返混比较严重， 因此需 要大的催化剂藏量来保证原料的转化率， 同时需要较大的沉降室 来分离催化剂和产物气体， 这就使二次反应几率大大增加， 不利 于过程的乙烯和丙烯总选择性。
对于提升管反应器来说， 由于气固并流上行， 返混较小， 因 此降低了催化剂藏量， 但提升管内的反应温度不易控制； 另外， 由于含氧化合物制烯烃的过程的反应速度并不快， 因此提升管反 应器很难实现原料的完全转化。
US6166282公开了一种用于 MT0过程的快速流化床反应器， 该反应器包括上部的分离区和下部的反应区， 其中反应区包括密 相区和在密相区上部的过渡区， 其中经密相区反应后的反应混合 物在过渡区完全转化。 与传统鼓泡床相比， 该快速流化床明显缩 小了反应器尺寸和降低了催化剂藏量， 从而节约了成本， 但仍存 在气固返混问题， 并且其中进入沉降室的气体需较长时间才进入 旋风分离器， 因此这部分气体仍会发生二次反应， 不利于乙烯和 丙烯的总选择性。
因此， 还需要进一步改进由含氧化合物制烯烃的反应器和方 法来提高反应物的转化率及产物的选择性。 发明内容
基于现有技术状况， 本发明提出了一种用于由含氧化合物制 烯烃的流化床反应器以及利用所述流化床反应器由含氧化合物制 烯烃的方法。
一方面， 本发明提供一种用于由含氧化合物制烯烃的流化床 反应器， 所述流化床反应器包括：
反应区， 所述反应区位于所述流化床反应器的下部， 包括下 部的密相区和上部的提升管， 其中所述密相区和所述提升管之间 过渡连接；
分离区， 所述分离区位于所述流化床反应器的上部， 包括沉 降室、 气固快分装置、 旋风分离器和集气室， 所述提升管向上延 伸进入所述分离区和在其出口处与所述气固快分装置的入口连 通， 所述气固快分装置的出口经快速导气管与所述旋风分离器的 入口连通， 和所述旋风分离器的出口与集气室连通， 所述集气室 设在反应器出口下方并与之连通； 和
从所述沉降室至所述密相区的催化剂循环管线， 以及从所述 沉降室和 /或所述密相区至催化剂再生装置的失活催化剂导出管 线和从催化剂再生装置至所述密相区的再生催化剂返回管线。
另一方面， 本发明还提供一种由含氧化合物制烯烃的方法， 所述方法利用上述流化床反应器， 和包括如下步骤：
使含有含氧化合物和稀释剂的原料物流经原料入口分布器进 入所述流化床反应器下部的密相区， 在其中与催化剂接触并发生 反应， 使部分原料转化成低碳烯烃， 同时使部分催化剂由于碳沉 积而失活；
使含有低碳烯烃、 未转化原料和催化剂的反应混合物由密相 区上行进入提升管， 在其中未转化原料在催化剂作用下继续反应 并基本上完全转化， 产生主要含有低碳烯烃和催化剂的提升管流 出物；
提升管流出物直接进入位于所述流化床反应器上部的分离 区， 首先进入气固快分装置， 在此分离出大部分催化剂后经快速 导气管进入旋风分离器， 在此进一步分离出剩余催化剂后进入集 气室， 然后经反应器出口离开， 所分离出来的催化剂均落入位于 所述分离区下部的沉降室； 和
经催化剂循环管线从所述沉降室导出一部分催化剂循环回到 所述密相区，以及经失活催化剂导出管线从所述沉降室和 /或所述 密相区导出一部分催化剂进入催化剂再生装置进行再生， 然后经 再生催化剂返回管线从催化剂再生装置导出一部分再生催化剂返 回所述密相区。
按照本发明， 所述流化床反应器的密相区内还可以设有内构 件， 所述内构件可以用于移除反应热和 /或抑制气固返混， 例如， 所述内构件可以为换热构件如换热器或换热盘管， 也可以为抑制 返混的构件如导流板或挡板， 甚至在一些情况下为了设计的简洁 和高效可以为这些构件的组合，从而可以更好地控制反应温度和 / 或达到更高的原料转化率。
按照本发明， 所述流化床反应器的尺寸设计使得所述密相区 的高径比为 0. 5-10， 优选为 0. 6-8， 更优选为 0. 8-5， 所述提升 管的高径比为 2-20， 优选为 3-15， 更优选为 4-10， 和所述密相 区与所述提升管的直径比为 2-10， 优选为 2. 5-8， 更优选为 3-6； 而在利用所述流化床反应器进行含氧化合物制烯烃的过程中， 操 作所述反应器使得物流在所述密相区内的表观气速为 0. 1- 2m/s， 优选为 0. 2- 1. 5m/s， 更优选为 0. 3- 1. 2m/s， 和停留时间为 0. 5-20s , 优选为 l-15s， 更优选为 2-10s， 和在所述提升管内的 表观气速为 2- 20m/s， 优选为 4- 18m/s， 更优选为 5- 15m/s， 和停 留时间为 0. 3-5s， 优选为 0. 4-4s， 更优选为 0. 5-3s， 从而可以 合理分配反应区的负荷以实现目标转化率。
按照本发明， 所述流化床反应器的沉降室位于分离区的下部 用于收集在分离区内分离出来的催化剂， 在所述沉降室中可以设 有汽提气分布器用于引入汽提气以汽提催化剂， 在此， 所述汽提 气可以为氮气或蒸汽， 优选为蒸汽， 所述汽提操作可以进一步分 离出催化剂中夹带的原料气体和 /或产物气体，和在所述沉降室上 方可以设有汽提混合物导出管线用于经旋风分离器导出汽提出来 的气体混合物。
按照本发明， 所述流化床反应器中的气固快分装置可以为本 领域中常用的任何合适的气固快分装置， 例如， 所述气固快分装 置可以选自粗旋式气固快分装置、 弹射式气固快分装置、 倒 L型 气固快分装置、 T 型气固快分装置、 多叶型气固快分装置和旋流 臂式气固快分装置， 从而可以使提升管流出物中的气固两相快速 分离。
按照本发明， 所述流化床反应器中的旋风分离器可以为本领 域中常用的任何合适的旋风分离器， 并且所述旋风分离器可以为 一组或多组旋风分离器， 和每组旋风分离器可以由一个、 二个或 三个旋风分离器串联而成， 从而可以进一步更为彻底地分离出产 物物流中夹带的固体催化剂。
按照本发明， 所述流化床反应器中的快速导气管使所述气固 快分装置与所述旋风分离器连通， 其尺寸设计应使物流在其内的 停留时间尽可能短， 例如， 使所述停留时间一般不超过 5s， 优选 不超过 4s， 更优选不超过 3s。
按照本发明， 对于所述流化床反应器的从所述分离区下部的 沉降室和 /或所述反应区的密相区至催化剂再生装置的失活催化 剂导出管线和从催化剂再生装置至所述密相区的再生催化剂返回 管线、 以及从所述分离区下部至所述密相区的催化剂循环管线， 这些管线均可带有相应的阀门来调节其中的催化剂流量， 通过调 节催化剂的循环流量和 /或再生催化剂的返回流量可以调节所述 密相区内的催化剂藏量和催化剂平均活性， 从而可以进一步调节 物料在密相区和提升管内所达到的转化率。
按照本发明， 还可以将离开所述流化床反应器的反应混合物 引入产物后处理装置进行分离和 /或纯化，以获得最终产物烯烃如 乙烯和 /或丙烯， 所述分离和 /或纯化操作可以通过本领域已知的 任何合适的装置如精馏塔和吸收塔等进行。
按照本发明， 进行转化的含氧化合物可以为本领域常用的低 碳醇和 /或醚类原料，例如，所述含氧化合物可以选自甲醇、乙醇、 丙醇、 二甲醚、 二乙醚、 二丙醚和它们的混合物， 优选为甲醇和 / 或二甲醚； 和所述低碳烯烃产物可以选自乙烯、 丙烯、 丁烯和它 们的混合物， 优选为乙烯和 /或丙烯。
按照本发明， 在所述原料中添加稀释剂是为了降低原料及相 应的反应产物的分压， 所述稀释剂可以为对所述含氧化合物转化 成烯烃的反应呈惰性的任何合适气体，例如，可以为蒸汽或氮气， 优选为蒸汽； 在此， 所述稀释剂的添加量可以为所述原料混合物 的 5- 80mol%， 优选为 10- 60mol%， 更优选为 15- 50mol%。
按照本发明， 用于将含氧化合物转化成烯烃的催化剂一般为 硅铝磷酸盐类分子筛催化剂，例如，所述催化剂可以选自 SAP0-5、 SAP0-11、 SAP0-17、 SAP0-41、 SAP0-34和 SAP0- 41，优选为 SAP0-34。
按照本发明，对于使所述含氧化合物在 SAP0类分子筛催化剂 上发生反应的反应温度， 本领域技术人员可以根据所使用的含氧 化合物相应地进行选择， 一般地， 所述含氧化合物转化成烯烃的 反应温度为 300- 优选为 400- 550 。
按照本发明，对于使所述含氧化合物在 SAP0类分子筛催化剂 上发生反应的反应压力， 本领域技术人员亦可以根据所使用的含 氧化合物并考虑所述流化床反应器的设计和操作成本相应地进行 选择， 一般地， 所述含氧化合物转化成烯烃的反应压力为 0. 05-lMPa绝压， 优选为 0. 1-0. 5MPa绝压。
按照本发明， 对于使所述含氧化合物在 SAP0类分子筛催化剂 上发生反应产生碳沉积造成所述催化剂失活来说， 使用含氧气体 如空气、 富氧空气或纯氧、 优选为空气烧除催化剂表面上的积碳 而使所述催化剂再生， 此时， 在催化剂再生装置内的温度一般为 500- 优选为 350- 和压力一般为 0. 05- IMPa绝压， 优选为 0. 15-0. 8MPa绝压。
按照本发明， 所述流化床反应器在反应区内组合了密相区和 提升管， 其中在密相区内转化大部分原料物流， 在提升管内进一 步转化剩余的原料物流， 据此， 在确定原料物流目标转化率的前 提下， 提升管的引入降低了密相区内的催化剂藏量负荷， 并且其 中可以通过催化剂的循环和再生操作来进一步调节密相区内的催 化剂藏量和催化剂平均活性， 从而可以进一步调节原料物流的反 应深度和进而优化原料物流的转化率。
按照本发明， 所述流化床反应器的分离区组合了气固快分装 置和旋风分离器， 其中气固快分装置快速分离出大部分催化剂， 然后利用旋风分离器进一步分离剩余催化剂， 其中利用快速导气 管连通所述气固快分装置和旋风分离器， 从而尽可能地缩短了气 固分离时间和相应地抑制和 /或减少了产物烯烃的二次反应， 因 此， 增加了目标产物的选择性。
因此， 本发明的流化床反应器组合了密相区和提升管来控制 反应深度并优化转化率， 并进一步组合了气固快分装置和旋风分 离器来实现气固快速分离而抑制二次反应， 从而保证了目标产物 烯烃的选择性。
因此，本发明的流化床反应器可以调节和 /或控制含氧化合物 转化成烯烃的总转化率和总选择性， 从而可以达到最佳的反应效 果， 进而实现本发明的目的。 附图说明
图 1是本发明的流化床反应器的一个具体实例， 其中密相区 和提升管之间通过锥形管段过渡连接， 气固快分装置采用粗旋式 气固快分装置， 旋风分离器为二个旋风分离器串联的二级旋风分 离器， 和其中从沉降室导出一部分失活催化剂进入催化剂再生装 置进行再生。
附图 1中各附图标记的具体含义如下： 1-原料入口分布器； 2 -密相区； 3-内构件； 4-提升管； 5 旋式气固快分装置； 6 旋式气固快分装置料腿； 7-快速导气管； 8-旋风分离器； 9-旋风 分离器料腿； 10-集气室； 11-反应器出口； 12-产物后处理装置； 13 -汽提气分布器； 14-沉降室； 15-汽提混合物导出管线； 16 -失 活催化剂导出管线； 17-失活催化剂流量调节阀； 18-催化剂再生 装置； 19-再生催化剂返回管线； 20-再生催化剂返回流量调节阀； 21 -催化剂循环管线； 22~flL化剂循环流量调节阀。 具体实施方式
下面结合附图进一步详细描述本发明的流化床反应器及相应 的由含氧化合物制烯烃的方法。
参考图 1，使含有含氧化合物如甲醇和 /或二甲醚及稀释剂如 蒸汽的原料物流经原料入口分布器 1进入流化床反应器下部的密 相区 2，在其中停留一定时间如 4- 10s与催化剂接触并发生反应， 使大部分原料转化成烯烃，同时使部分催化剂由于碳沉积而失活， 反应放出的热量经内构件 3 移出使反应温度保持在一定范围如 400-550 内， 和反应压力保持为 0. 1-0. 5MPa绝压， 内构件 3同 时抑制气固返混；
使含有烯烃、 未转化原料和催化剂的反应混合物由密相区 2 上行进入提升管 4，在其中停留一定时间如 1- 3s使未转化原料在 催化剂作用下继续反应并基本上完全转化， 产生主要含有低碳烯 烃和催化剂的提升管流出物；
提升管流出物首先进入粗旋式气固快分装置 5， 在此分离出 大部分催化剂后经快速导气管 7短暂停留如不超过 3s后进入旋风 分离器 8， 在此进一步分离出剩余催化剂后进入集气室 10， 然后 经反应器出口 11离开， 再然后去往产物后处理装置 12进行分离 得到最终产物烯烃， 而分离出的催化剂分别经粗旋式气固快分装 置 5的料腿 6及旋风度分离器 8的料腿 9落入分离区下部的沉降 室 14;
经催化剂循环管线 21从沉降室 14导出一部分催化剂循环回 到密相区 2， 和经失活催化剂导出管线 16从沉降室 14导出一部 分催化剂进入催化剂再生装置 18 在 500- 750 的温度及 0. 1-0. 5MPa绝压的压力下烧除积碳进行再生， 然后经再生催化剂 返回管线 19从催化剂再生装置 18导出一部分再生催化剂返回密 相区 2， 在此， 通过各管线上的阀门调节催化剂的循环流量、 失 活催化剂的导出流量以及再生催化剂的返回流量， 从而使密相区 2内的催化剂活性更适合所述转化反应； 和
经汽提气分布器 13向沉降室 14内引入蒸汽作为汽提气以汽 提催化剂，和汽提出来的气体混合物经沉降室 14上方的汽提混合 物导出管线 15导出后经旋风分离器 8离开。
下面通过具体实施例和对比例进一步描述本发明， 所述实施 例不用于限制本发明的范围。
实施例
实施例 1
采用图 1所示的流化床反应器， 使用甲醇和蒸汽的混合物作 为原料， 蒸汽作为稀释剂占混合物的 30mol%， 原料总流量为 67kmol/h; 催化剂采用 SAP0-34 , 装填量为 0. 5t; 保持反应温度 为 485 , 反应压力为 0. 2 MPa绝压， 密相区内的表观气速为 0. 5 m/s和停留时间为 5s; 提升管内的表观气速为 8 m/s和停留时间 为 1. 5 s， 快速导气管内的停留时间为 1 s， 汽提剂采用水蒸汽， 水蒸汽的流量是 50 kg/h; 从沉降室至密相区的催化剂循环流量 为 3t/h， 从沉降室至催化剂再生装置的失活催化剂流量是 0. 5 t/h， 催化剂再生装置内的再生温度为 650 和再生压力为 0. 22 MPa绝压， 和返回密相区的再生催化剂流量是 0. 5 t/h; 在反应器 出口处进行取样测量确定甲醇的转化率为 99. 95mol%， 乙烯 +丙烯 的总选择性为 82. 3mol%e 对比例 1
与实施例 1相同进行甲醇的转化反应， 只是采用传统的鼓泡 式流化床反应器， 其中采用密相区作为反应区， 和分离区采用三 个旋风分离器串联的三级旋风分离器， 其中催化剂的装填量为
2t， 密相区内的表观气速为 0. 3 m/s和停留时间为 8 s; 在反应 器出口处进行取样测量确定甲醇的转化率为 99. 65mol%， 乙烯 +丙 烯的总选择性为 78. 3mol%。 实施例 1和对比例 1的结果表明， 与传统的鼓泡床反应器相 比， 采用本发明的流化床反应器可在明显降低的催化剂藏量下获 得更高的转化率及更高的乙烯 +丙烯的总选择性，这表明本发明的 流化床反应器取得了明显的技术进步。

Claims

权利要求
1. 一种用于由含氧化合物制烯烃的流化床反应器， 所述流化 床反应器包括：
反应区， 所述反应区位于所述流化床反应器的下部， 包括下 部的密相区和上部的提升管， 其中所述密相区和所述提升管之间 过渡连接；
分离区， 所述分离区位于所述流化床反应器的上部， 包括沉 降室、 气固快分装置、 旋风分离器和集气室， 所述提升管向上延 伸进入所述分离区和在其出口处与所述气固快分装置的入口连 通， 所述气固快分装置的出口经快速导气管与所述旋风分离器的 入口连通， 和所述旋风分离器的出口与集气室连通， 所述集气室 设在反应器出口下方并与之连通； 和
从所述沉降室至所述密相区的催化剂循环管线， 以及从所述 沉降室和 /或所述密相区至催化剂再生装置的失活催化剂导出管 线和从催化剂再生装置至所述密相区的再生催化剂返回管线。
2. 权利要求 1的流化床反应器， 其中所述密相区内还设有内 构件， 所述内构件用于移除反应热和 /或抑制气固返混。
3. 权利要求 1或 2的流化床反应器， 其中所述流化床反应器 的尺寸设计使得所述密相区的高径比为 0. 5-10， 优选为 0. 6-8， 更优选为 0. 8-5， 所述提升管的高径比为 2-20， 优选为 3-15， 更 优选为 4-10，和所述密相区与所述提升管的直径比为 2-10，优选 为 2. 5-8， 更优选为 3-6。
4. 权利要求 1-3任一项的流化床反应器， 其中所述沉降室位 于所述分离区的下部用于收集在分离区内分离出来的催化剂， 在 所述沉降室中设有汽提气分布器用于引入汽提气以汽提催化剂， 和在所述沉降室上方设有汽提混合物导出管线用于经旋风分离器 导出汽提出来的气体混合物。
5. 权利要求 1-4任一项的流化床反应器， 其中所述气固快分 装置选自粗旋式气固快分装置、 弹射式气固快分装置、 倒 L型气 固快分装置、 T 型气固快分装置、 多叶型气固快分装置和旋流臂 式气固快分装置。
6. 权利要求 1-5任一项的流化床反应器， 其中所述旋风分离 器为一组或多组旋风分离器， 和每组旋风分离器由一个、 二个或 三个旋风分离器串联而成。
7. 权利要求 1-6任一项的流化床反应器， 其中所述快速导气 管的尺寸设计使得物流在其内的停留时间不超过 5s，优选不超过 4s , 更优选不超过 3s。
8. 一种由含氧化合物制烯烃的方法， 所述方法利用权利要求 1-7任一项的流化床反应器， 和包括如下步骤：
使含有含氧化合物和稀释剂的原料物流经原料入口分布器进 入所述流化床反应器下部的密相区， 在其中与催化剂接触并发生 反应， 使部分原料转化成低碳烯烃， 同时使部分催化剂由于碳沉 积而失活；
使含有低碳烯烃、 未转化原料和催化剂的反应混合物由密相 区上行进入提升管， 在其中未转化原料在催化剂作用下继续反应 并基本上完全转化， 产生主要含有低碳烯烃和催化剂的提升管流 出物；
提升管流出物直接进入位于所述流化床反应器上部的分离 区， 首先进入气固快分装置， 在此分离出大部分催化剂后经快速 导气管进入旋风分离器， 在此进一步分离出剩余催化剂后进入集 气室， 然后经反应器出口离开， 所分离出来的催化剂均落入位于 所述分离区下部的沉降室； 和
经催化剂循环管线从所述沉降室导出一部分催化剂循环回到 所述密相区，以及经失活催化剂导出管线从所述沉降室和 /或所述 密相区导出一部分催化剂进入催化剂再生装置进行再生， 然后经 再生催化剂返回管线从催化剂再生装置导出一部分再生催化剂返 回所述密相区。
9. 权利要求 8的方法， 还包括经所述密相区内的内构件移除 反应热和 /或抑制气固返混。
10. 权利要求 8或 9的方法， 其中物流在所述密相区内的表 观气速为 0. l-2m/s , 优选为 0. 2-1. 5m/s， 更优选为 0. 3-1. 2m/s， 和停留时间为 0. 5- 20s， 优选为 1- 15s， 更优选为 2- 10s， 和在所 述提升管内的表观气速为 2-20m/s， 优选为 4-18m/s， 更优选为 5-15m/s，和停留时间为 0. 3-5s，优选为 0. 4-4s，更优选为 0. 5-3s。
11. 权利要求 8- 10任一项的方法， 还包括经汽提气分布器向 所述沉降室内引入汽提气以汽提催化剂， 所述汽提气为氮气或蒸 汽， 优选为蒸汽， 和汽提出来的气体混合物经所述沉降室上方的 汽提混合物导出管线导出后经旋风分离器离开。
12. 权利要求 8-11 任一项的方法， 其中所述催化剂选自 SAP0- 5、 SAP0- 11、 SAP0- 17、 SAP0- 41、 SAP0- 34和 SAP0- 41， 优 选为 SAP0-34。
13. 权利要求 12的方法， 其中所述流化床反应器内的反应温 度为 300- 优选为 400- 550 , 和反应压力为 0. 05- IMPa 绝压， 优选为 0. 1-0. 5MPa绝压。
14. 权利要求 8- 13任一项的方法， 其中所述含氧化合物选自 甲醇、 乙醇、 丙醇、 二甲醚、 二乙醚、 二丙醚和它们的混合物， 优选为甲醇和 /或二甲醚，所述稀幹剂为氮气或蒸汽，优选为蒸汽， 和所述烯烃选自乙烯、 丙烯、 丁烯和它们的混合物， 优选为乙烯 和 /或丙烯。
15. 权利要求 8- 14任一项的方法， 还包括将离开所述流化床 反应器的反应混合物引入产物处理装置进行分离和 /或纯化，以获 得最终烯烃产物。
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