WO2016061906A1 - 一种低碳烯烃的制造方法 - Google Patents

一种低碳烯烃的制造方法 Download PDF

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WO2016061906A1
WO2016061906A1 PCT/CN2015/000705 CN2015000705W WO2016061906A1 WO 2016061906 A1 WO2016061906 A1 WO 2016061906A1 CN 2015000705 W CN2015000705 W CN 2015000705W WO 2016061906 A1 WO2016061906 A1 WO 2016061906A1
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catalyst
reaction
reactor
produced
mpa
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PCT/CN2015/000705
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English (en)
French (fr)
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崔守业
许友好
于敬川
李明罡
宗保宁
唐津莲
王新
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中国石油化工股份有限公司
中国石油化工股份有限公司石油化工科学研究院
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Priority to SG11201703273SA priority Critical patent/SG11201703273SA/en
Priority to CN201580011656.6A priority patent/CN106687428B/zh
Priority to EP15852702.8A priority patent/EP3210960B1/en
Priority to RU2017115709A priority patent/RU2698107C2/ru
Priority to AU2015336835A priority patent/AU2015336835B2/en
Priority to JP2017521188A priority patent/JP6782233B2/ja
Priority to US15/520,729 priority patent/US10737991B2/en
Publication of WO2016061906A1 publication Critical patent/WO2016061906A1/zh

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    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J29/00—Catalysts comprising molecular sieves
    • B01J29/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites
    • B01J29/06—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof
    • 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
    • C07C5/42—Preparation of hydrocarbons from hydrocarbons containing the same number of carbon atoms by dehydrogenation with a hydrogen acceptor
    • C07C5/48—Preparation of hydrocarbons from hydrocarbons containing the same number of carbon atoms by dehydrogenation with a hydrogen acceptor with oxygen as an acceptor
    • 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
    • B01J21/00—Catalysts comprising the elements, oxides, or hydroxides of magnesium, boron, aluminium, carbon, silicon, titanium, zirconium, or hafnium
    • B01J21/02—Boron or aluminium; Oxides or hydroxides thereof
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J29/00—Catalysts comprising molecular sieves
    • B01J29/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites
    • B01J29/06—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof
    • B01J29/40—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof of the pentasil type, e.g. types ZSM-5, ZSM-8 or ZSM-11, as exemplified by patent documents US3702886, GB1334243 and US3709979, respectively
    • 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
    • B01J29/82—Phosphates
    • B01J29/84—Aluminophosphates containing other elements, e.g. metals, boron
    • B01J29/85—Silicoaluminophosphates [SAPO compounds]
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • 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
    • B01J38/12—Treating with free oxygen-containing gas
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C1/00—Preparation of hydrocarbons from one or more compounds, none of them being a hydrocarbon
    • C07C1/20—Preparation of hydrocarbons from one or more compounds, none of them being a hydrocarbon starting from organic compounds containing only oxygen atoms as heteroatoms
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C1/00—Preparation of hydrocarbons from one or more compounds, none of them being a hydrocarbon
    • C07C1/20—Preparation of hydrocarbons from one or more compounds, none of them being a hydrocarbon starting from organic compounds containing only oxygen atoms as heteroatoms
    • C07C1/24—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 elimination of water
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C11/00—Aliphatic unsaturated hydrocarbons
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C4/00—Preparation of hydrocarbons from hydrocarbons containing a larger number of carbon atoms
    • C07C4/02—Preparation of hydrocarbons from hydrocarbons containing a larger number of carbon atoms by cracking a single hydrocarbon or a mixture of individually defined hydrocarbons or a normally gaseous hydrocarbon fraction
    • C07C4/06—Catalytic processes
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J29/00—Catalysts comprising molecular sieves
    • B01J29/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites
    • B01J29/06—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof
    • B01J29/061—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof containing metallic elements added to the zeolite
    • 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
    • B01J8/1872—Details of the fluidised bed reactor
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C2529/00—Catalysts comprising molecular sieves
    • C07C2529/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites, pillared clays
    • C07C2529/06—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof
    • C07C2529/40—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof of the pentasil type, e.g. types ZSM-5, ZSM-8 or ZSM-11
    • 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)
    • F—MECHANICAL ENGINEERING; LIGHTING; HEATING; WEAPONS; BLASTING
    • F28—HEAT EXCHANGE IN GENERAL
    • F28D—HEAT-EXCHANGE APPARATUS, NOT PROVIDED FOR IN ANOTHER SUBCLASS, IN WHICH THE HEAT-EXCHANGE MEDIA DO NOT COME INTO DIRECT CONTACT
    • F28D21/00—Heat-exchange apparatus not covered by any of the groups F28D1/00 - F28D20/00
    • F28D2021/0019—Other heat exchangers for particular applications; Heat exchange systems not otherwise provided for
    • F28D2021/0022—Other heat exchangers for particular applications; Heat exchange systems not otherwise provided for for chemical reactors
    • 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

  • This invention relates to a process for the manufacture of light olefins from oxygenate feedstocks. More specifically, the present invention relates to a process for increasing the production of light olefins in a process for producing a light olefin from an oxygenate feedstock.
  • low-carbon olefins As a basic organic chemical raw material, low-carbon olefins (C 2 -C 4 olefins) play an important role in the modern petroleum and chemical industries.
  • the methods for making low-carbon olefins can be broadly classified into two broad categories, namely, conventional oil routes and emerging non-oil routes. Since the 1910s, countries around the world have begun to develop non-oil resources (especially oxygenated feedstocks) to make low-carbon olefins, and some progress has been made.
  • the catalyst is circulated between the reactor and the regenerator.
  • the reactor and regenerator are typically operated at substantially the same pressure.
  • the reactor is a hydrocarbon atmosphere
  • the regenerator is an oxygen-containing atmosphere. If the two are not well separated, there is a great safety hazard.
  • the prior art low-carbon olefin production unit generally adopts a cyclone separator similar to the catalytic cracking unit, and the natural running loss of the catalyst during the production process is unavoidable, especially when the catalyst has a particle size of not more than 20 ⁇ m.
  • the fine powder is increased, this adversely affects the subsequent product separation and is also disadvantageous for repeated use of the catalyst.
  • An object of the present invention is to provide a process for producing a low-carbon olefin which overcomes the aforementioned disadvantages of the prior art and which can easily achieve the purpose of increasing the production of light olefins by directly utilizing the existing reactor.
  • the inventors of the present invention have surprisingly discovered that if the weight hourly space velocity of the oxygenate raw material is increased correspondingly while increasing the reaction pressure, the yield of the low carbon olefin can be maintained to be comparable or even more than the prior art. a high level, which does not decrease as previously expected in the prior art, as a result of which, for existing reactors, it is possible to increase the reaction pressure and the weight hourly space velocity of the reactor by the provisions of the present invention.
  • the amount of oxygenate feedstock treatment in the reactor is increased in magnitude to increase the yield of light olefins (increasing production of light olefins). This finding by the inventors has broken through the conventional knowledge of those skilled in the art and has completed the present invention based on this finding.
  • the present invention relates to the following aspects.
  • a method for producing a low-carbon olefin (or a method for increasing yield), characterized in that the dehydration reaction is carried out in a method of producing a light olefin by continuously causing a dehydration reaction by contacting an oxygen-containing raw material with a catalyst.
  • the reaction pressure P is 0.5-10 MPa, preferably 0.75-3.5 MPa, more preferably 0.8-3 MPa, most preferably 1-2 MPa
  • the weight hourly space velocity H of the dehydration reaction is 7-250 h -1 , preferably 8-150 h -1 . More preferably, it is 10-100 h -1 , more preferably 15-80 h -1 , and most preferably 15-50 h -1 .
  • H f(P)
  • P unit is MPa
  • H belongs to the interval [0.55, 10.0]
  • H (the unit is h -1 ) belongs to the interval [7, 250], preferably belongs to the interval [8, 150]
  • reaction pressure P of the dehydration reaction is at least 0.35 MPa higher than the regeneration pressure of the regeneration reaction, preferably at least 0.4 MPa higher, at least 0.5 MPa higher, at least 0.6 MPa higher, at least 0.7 MPa higher, at least 0.8 MPa higher, at least high.
  • 0.9MPa at least 1.0MPa, at least 1.1MPa, at least 1.2MPa, at least 1.3MPa, at least 1.4MPa, at least 1.5MPa, at least 1.6MPa, at least 1.7MPa, at least 1.8MPa, at least high 1.9 MPa or at least 2.0 MPa high.
  • At least a portion of the regenerated catalyst and/or at least a portion of the further regenerated catalyst are recycled to the dehydration reaction and/or the further reaction.
  • the number of reactors for performing the dehydration reaction and/or the number of reactors for performing the further reaction is one or more, and each is independently selected a fluidized bed reactor, a dense phase bed reactor, a riser reactor, an ebullated bed reactor, a slurry bed reactor, and a composite form of two or more of these reactors, preferably selected from the riser reaction More preferably, each of them is independently selected from the group consisting of a constant diameter riser reactor, a constant line riser reactor, a variable diameter riser reactor, and a riser composite dense phase bed reactor.
  • R1 and R2 are the same or different from each other, each independently selected from hydrogen and C1-6 branched or straight chain
  • the alkyl groups preferably each independently selected from the group consisting of hydrogen and a C 1-4 branched or linear alkyl group, provided that at most one of R 1 and R 2 is hydrogen, more preferably selected from the group consisting of methanol, ethanol, and At least one of methyl ether, diethyl ether, methyl ethyl ether, dimethyl carbonate, and methyl formate.
  • the catalyst and the further catalyst are the same or different from each other, each independently selected from at least one of molecular sieve catalysts, preferably each independently selected from the group consisting of silicoaluminophosphates. At least one of a salt molecular sieve catalyst and an aluminosilicate molecular sieve catalyst.
  • reaction conditions of the regeneration reaction comprise: a reaction temperature of 450 to 850 ° C, preferably 550 to 700 ° C; a reaction pressure of 0.1 to 0.5 MPa, preferably 0.15 to 0.3 MPa; An oxygen atmosphere, preferably an air atmosphere or an oxygen atmosphere.
  • the manufacturing method is capable of increasing the yield of low olefins by 50% while maintaining the size and number of reactors for carrying out the dehydration reaction. Preferably, it is increased by 100%, more preferably by 150%, 200%, 500% or 790%, and most preferably by 1000% or more.
  • the olefin-rich oil and gas and the catalyst to be produced are separated in the oil separation zone, and the separated olefin-rich oil and gas is sent to the product separation and recovery system, and the catalyst to be produced is passed through the stripping section steam in the riser type reactor. Extracted from the riser type reactor and sent to the catalyst receiver to be produced;
  • the catalyst to be produced in the catalyst receiver to be produced is directly sent to the regenerator through the catalyst hopper, or first to the catalyst feeder through the catalyst hopper and then to the regenerator, and in the regenerator under an oxygen-containing atmosphere. Charring regeneration to obtain a regenerated catalyst;
  • the method for producing a low carbon olefin of the present invention has the following advantages in comparison with the prior art.
  • the weight hourly space velocity of the oxygenate raw material is increased correspondingly while increasing the reaction pressure, and the lower olefin can be made without changing the size and the number of the existing reactor or the reaction device.
  • the yield is maintained at a level that is comparable to or even higher than the prior art, and ultimately a large (up to 790%) increase in the production of light olefins.
  • the method for producing a low-carbon olefin of the present invention belongs to a method for increasing the yield of a low-carbon olefin, and can be applied to the modification or capacity upgrading of an existing low-carbon olefin production unit.
  • the process for producing a low-carbon olefin according to the present invention can significantly reduce the size and amount of the reactor or the reaction device, thereby reducing the overall production of low-carbon olefins, while ensuring that a predetermined low-carbon olefin production is achieved, as compared with the prior art.
  • the low carbon olefin production method of the present invention is a new generation of high-capacity low-carbon olefin production method, which can be applied to the construction of a smaller scale and lower investment cost than the existing low-carbon olefin production apparatus.
  • the regenerator is maintained to operate at a lower pressure while operating the reactor at a higher pressure, thereby reducing the complexity of the entire low carbon olefin production method and manufacturing apparatus.
  • the reaction pressure of the reactor is significantly higher than the regeneration pressure of the regenerator, whereby the hydrocarbon atmosphere and regenerator of the reactor can be realized by using a pressure switching device such as a lock hopper or a catalyst hopper.
  • a pressure switching device such as a lock hopper or a catalyst hopper.
  • FIG. 1 is a schematic view showing the flow of a method for producing a low-carbon olefin from an oxygen-containing compound according to a first embodiment of the present invention.
  • Figure 2 is a schematic view showing the flow of a method for preparing a low-carbon olefin from an oxygen-containing compound according to a second embodiment of the present invention.
  • Figure 3 is a schematic view showing the flow of a method for producing a light olefin from an oxygen-containing compound according to a third embodiment of the present invention.
  • Figure 4 is a flow chart showing the process for preparing a low-carbon olefin from an oxygen-containing compound according to a fourth embodiment of the present invention.
  • Figure 5 is a schematic view showing the flow of a method for producing a light olefin from an oxygen-containing compound according to a fifth embodiment of the present invention.
  • Figure 6 is a schematic view showing the flow of a method for producing a low-carbon olefin from an oxygen-containing compound according to a sixth embodiment of the present invention.
  • Figure 7 is a schematic view showing the flow of a method for preparing a low-carbon olefin from an oxygen-containing compound according to a seventh embodiment of the present invention.
  • the present invention may also include other specific embodiments, and is not limited to the above seven.
  • riser reactor 202 riser and distribution plate 203 dense phase bed reactor
  • reaction product line 626 pre-lift line 627 another riser type reactor
  • the C4 + hydrocarbon of the present invention means a hydrocarbon of C4 or higher.
  • lower olefin means ethylene and propylene.
  • yield of lower olefins refers to the single pass yield of lower olefins
  • yield of lower olefins refers to the single pass yield of lower olefins per unit time per unit of reactor.
  • the heavy hourly space velocity refers to the mass of the reactant passing through the unit mass of catalyst per unit time.
  • Yield product yield / sum of yield of hydrocarbon products other than oxygenates ⁇ 100.
  • the hydrocarbons other than the oxygen-containing compound specifically include hydrogen, C1 and C1 or higher non-oxygenated hydrocarbons.
  • a process for producing a low-carbon olefin which produces a low-carbon olefin by continuously bringing an oxygenate raw material into contact with a catalyst to cause a dehydration reaction.
  • the manufacturing method may include the steps of: continuously bringing the oxygenate raw material into contact with the catalyst to cause the dehydration reaction, obtaining a low-carbon olefin-rich oil and gas and a catalyst to be produced, at least a part
  • the spent catalyst is sent to a regeneration reaction to obtain a regenerated catalyst, and at least a portion of the regenerated catalyst is recycled to the dehydration reaction.
  • the manufacturing method may include, for example, continuously contacting the oxygenate feedstock in a reactor (such as a riser type reactor) with the catalyst to perform the dehydration reaction, resulting in a low enrichment Carbon olefin oil and gas and catalyst; the olefin-rich oil and gas and the catalyst to be produced are separated in the oil separation zone, and the separated olefin-rich oil and gas is sent to the product separation and recovery system, and the catalyst to be produced is passed through the reactor.
  • the stripping section in the stripping section is taken out from the reactor and sent to the catalyst receiver to be produced; the catalyst to be produced in the catalyst receiver to be produced is directly sent to the regenerator through the catalyst hopper, or is first sent to the regenerator through the catalyst hopper.
  • the spent catalyst feeder is then sent to the regenerator, and is subjected to charring regeneration in an oxygen-containing atmosphere in the regenerator to obtain a regenerated catalyst; the regenerated catalyst is taken out from the regenerator and sent to the regenerated catalyst receiver, and then passed through the catalyst.
  • the hopper is delivered to the regenerated catalyst feeder, or the regenerated catalyst is delivered directly to the catalyst hopper and then returned to the reactor.
  • regenerator any type conventionally known in the art, such as a fluidized bed regenerator or an ebullated bed regenerator, can be directly used, but is not limited thereto.
  • the manufacturing method may further include: withdrawing a portion of the catalyst to be produced from the reactor or the catalyst receiver to be produced; and returning the portion of the catalyst to be produced to the catalyst immediately or after taking heat to return to the temperature Returning to the reactor in the reactor, or in a catalyst mixer delivered to the lower portion of the reactor, after mixing with the regenerated catalyst; the amount of the portion of the catalyst to be produced that is withdrawn is transported to the regeneration through the catalyst hopper The regenerated catalyst in the catalyst feeder together is sufficient to maintain continuous operation of the catalyst in the reactor.
  • the manufacturing method may further include: extracting a portion of the catalyst to be produced from the reactor or the catalyst receiver to be produced; and transferring the portion of the catalyst to be produced directly or after taking heat to lower the temperature, and then transporting Returning to the reactor after mixing with the regenerated catalyst in the regenerated catalyst feeder; the amount of the portion of the catalyst to be produced that is taken out is together with the regenerated catalyst delivered to the regenerated catalyst feeder through the catalyst hopper To maintain continuous operation of the catalyst in the reactor.
  • the oxygenate starting material is well known to those skilled in the art, may be at least one selected from the group consisting of alcohols, ethers and esters, and may also be other industrial or natural oxygenates, and the invention is not limited. .
  • R 1 and R 2 are the same or different from each other, each independently selected from hydrogen and a C 1-6 branched or linear alkyl group, preferably each independently selected from hydrogen and a C 1-4 branched or linear alkyl group, provided that It is that at most one of R1 and R2 is hydrogen.
  • the oxygenate raw material at least one selected from the group consisting of methanol, ethanol, dimethyl ether, diethyl ether, methyl ethyl ether, dimethyl carbonate, and methyl formate, in particular, methanol is more preferable.
  • the diluent water vapor is generally used, and hydrogen, methane, ethane, nitrogen, carbon monoxide or the like can also be used.
  • the molar ratio of the oxygenate feedstock to the diluent is generally from 40:1 to 0.4:1, preferably from 11:1 to 0.7:1, more preferably from 7:1 to 1.3:1.
  • the catalyst can be of a type well known to those skilled in the art.
  • the catalyst may be a molecular sieve catalyst, and the molecular sieve may be a silicoaluminophosphate-based molecular sieve and/or an aluminosilicate molecular sieve.
  • the silicoaluminophosphate molecular sieve may be selected from one or more of SAPO series, SRM series molecular sieves
  • the aluminosilicate molecular sieve may be selected from one or more of ZSM series and ZRP series molecular sieves. .
  • the molecular sieve may be supported by an alkaline earth metal, K, Mg, Ca, Ba, Zr, Ti, Co, Mo, Ni, Pt, Pd, La, Ce, Cu, Fe, B, Si, P, Sn, One or several elements of Pb, Ga, Cr, V, Sc, Ge, Mn, La, Al, Ni, Fe.
  • the manufacturing method may further comprise the step of separating the low carbon olefin-rich oil and gas to obtain C 4 + hydrocarbons.
  • the present invention may optionally further comprise the step of: continuously the C 4 + hydrocarbons further contact with the catalyst and further reaction occurs, and further to obtain oil to be further enriched in light olefins Producing a catalyst, delivering at least a portion of the further catalyst to be regenerated to the regeneration reaction, obtaining a further regenerated catalyst, and recycling at least a portion of the regenerated catalyst and/or at least a portion of the further regenerated catalyst to the dehydration Reaction and / or the further reaction.
  • the method may comprise a further 4 + hydrocarbons is fed to the reactor via the product recovery system C separation isolated (such as a riser reactor) for the further reaction.
  • the further catalyst and the catalyst may be the same or different and may be of a type well known to those skilled in the art.
  • the further catalyst it may be a molecular sieve catalyst, and the molecular sieve may be a silicoaluminophosphate-based molecular sieve and/or an aluminosilicate molecular sieve.
  • the silicoaluminophosphate molecular sieve may be selected from one or more of SAPO series, SRM series molecular sieves
  • the aluminosilicate molecular sieve may be selected from one or more of ZSM series and ZRP series molecular sieves. .
  • the molecular sieve may be supported by an alkaline earth metal, K, Mg, Ca, Ba, Zr, Ti, Co, Mo, Ni, Pt, Pd, La, Ce, Cu, Fe, B, Si, P, Sn, One or several elements of Pb, Ga, Cr, V, Sc, Ge, Mn, La, Al, Ni, Fe.
  • the manufacturing method may further include: feeding the regenerated catalyst in the regenerated catalyst feeder into the further reactor to contact the C 4 + hydrocarbons and performing the further reaction, The resulting lower olefin-rich oil and gas and further spent catalyst are fed together into the oil separation zone of the reactor.
  • the manufacturing method may further include: feeding the catalyst to be produced in the reactor to the further reactor to contact the C 4 + hydrocarbons and performing the further reaction, resulting in further
  • the low-carbon olefin-rich oil and gas and the further spent catalyst are separated in the further reactor, and the separated further low-carbon olefin-rich oil and gas is sent to the product separation and recovery system, and the separated further The spent catalyst is sent to the spent catalyst receiver.
  • the manufacturing method may further include: feeding the catalyst to be produced in the reactor to the further reactor to contact the C 4 + hydrocarbons and performing the further reaction, resulting in further The low carbon olefin-rich oil and gas and further spent catalyst are fed into the oil separation zone of the reactor.
  • the manufacturing method may further include: feeding the regenerated catalyst in the regenerator directly into the further reactor to contact the C 4 + hydrocarbons and performing the further reaction to obtain a further enrichment a low carbon olefin oil and gas and a further spent catalyst; separating the further low carbon olefin-rich oil and gas and the further spent catalyst in the further reactor, further enriching after separation
  • the oil and gas of the low carbon olefin is fed to the product separation and recovery system, and the further spent catalyst is directly fed into the regenerator for regeneration.
  • the number of the reactors and/or the further reactors is one or more, and is not particularly limited. Additionally, the reactor and/or the further counter The reactors are the same or different from each other, and are each independently selected from a fluidized bed reactor, a dense phase bed reactor, a riser reactor, an ebullated bed reactor, a slurry bed reactor, and two or more of these reactors. The compound form of the species. Preferably, the reactor and/or the further reactors are identical or different from each other, each independently selected from a riser reactor, more preferably each independently selected from an equal diameter riser reactor, etc. Tube reactor, variable diameter riser reactor and riser composite dense phase bed reactor.
  • the riser type reactor may be provided with a pre-lift section, a riser, a chilling medium line, an expanded diameter riser, a reduced diameter, a fast split, a stripping section, a dense phase section, and a settling from the bottom to the top in the vertical direction.
  • a device commonly used in the industry such as a zone, a catalyst mixer, a filter, etc., enables the reactor to be continuously operated; wherein the settling zone, a filter, etc. can constitute the oil separation zone, and the oil separation zone is also Other means for separating the catalyst to be produced from the oil and gas may be included, and the invention is not limited.
  • the dense phase bed of the riser type reactor may not form a dense phase bed, i.e., "zero level".
  • the reactor can be provided with one or more chilled media lines to control the reaction temperature.
  • the chilling medium can be injected into the reactor through one or more chilling medium lines disposed in the middle and downstream of the reactor (relative to the flow direction of the material).
  • the chill medium may be a chiller or a cooled catalyst, and the chiller may be the oxygenate raw material and/or water that is not preheated.
  • the reaction temperature of the dehydration reaction is from 200 to 700 ° C, preferably from 250 to 600 ° C.
  • the reaction pressure P of the dehydration reaction is generally from 0.5 to 10 MPa, preferably from 0.75 to 3.5 MPa, more preferably from 0.8 to 3 MPa, and most preferably from 1 to 2 MPa.
  • the dehydration reaction is generally H WHSV 7-250h -1, preferably 8-150h -1, more preferably 10-100h -1, and more preferably 15-80h -1, most preferably 15-50h -1.
  • f(P) is established.
  • P (unit is MPa) belongs to the interval [0.55, 10.0], preferably belongs to the interval [0.75, 3.5], more preferably belongs to the interval [0.8, 3.0], more preferably belongs to the interval [1.0, 2.0], and H (unit) is h -1) belongs to the interval [7,250], preferably belongs to the interval [8,150], and more preferably belongs to the interval [10,100], and more preferably belongs to the interval [15,80], and most preferably it belongs to the interval [15,50 ].
  • H unit is MPa
  • H (unit) is h -1) belongs to the interval [7,250], preferably belongs to the interval [8,150], and more preferably belongs to the interval [10,100], and more preferably belongs to the interval [15,80], and most preferably it belongs to the interval [15,50 ].
  • the present invention does not particularly limit the manner, magnitude, and the like of the reaction pressure P and the weight hourly space velocity H, as long as the respective values have indeed increased based on the conventional judgment of those skilled in the art, but It can be kept constant or reduced. According to a particular embodiment of the invention, it is preferred that the reaction pressure P increases in proportion to the weight hourly space velocity H or increases according to different or the same amplitude, and may sometimes be an equal increase or a synchronous increase until the expected increase in the yield of the low carbon olefin is achieved. .
  • the weight hourly space velocity H generally also preferably reaches the upper limit of a certain numerical interval specified in the foregoing invention (for example, 50 h). -1 ), but is not limited to this.
  • the reaction pressure P and the weight hourly space velocity H are different within the above-mentioned numerical range or numerical range specified in the present invention, even when the reaction pressure P is increased, the weight hourly space velocity H is also increased. It is also impossible to obtain a large-scale effect of increasing the yield of low-carbon olefins as shown in the present invention (as shown in the examples). This is entirely beyond the expectation of those skilled in the art.
  • the reaction conditions for the further reaction include a reaction temperature of 200 to 700 ° C, preferably 300 to 600 ° C; and a reaction pressure of 0.1 to 6 MPa, preferably 0.8 to 2 MPa.
  • the manufacturing method may further include controlling a ratio of the reaction pressure P in the reactor to a regeneration pressure in the regenerator to be from 3 to 100:1. More specifically, according to the present invention, the reaction pressure P of the dehydration reaction is at least 0.35 MPa higher than the regeneration pressure of the regeneration reaction, preferably at least 0.4 MPa higher, at least 0.5 MPa higher, at least 0.6 MPa higher, and at least 0.7 higher.
  • the reaction pressure P of the dehydration reaction is generally at most 5 MPa higher than the regeneration pressure of the regeneration reaction, preferably at most 4 MPa, at most 3.5 MPa, at most 3.3 MPa, at most 3 MPa, at most 2.5 MPa. At most 2.3 MPa, at most 2 MPa, at most 1.5 MPa, at most 1.3 MPa or at most 1 MPa.
  • the reactor, the regenerator, and the regenerated catalyst feeder can be used.
  • One or more internal heat extractors are provided in the regenerative catalyst receiver.
  • the inner heat extractor may be of a coil type, a bent tube or the like, and the reactor is heated by a liquid such as internal flowing water or carbon tetrachloride.
  • the internal heat extractor commonly used in other industries may also be used in the present invention. application.
  • the manufacturing method may further comprise recycling at least a portion of the catalyst to be produced and/or at least a portion of the further catalyst to be produced to the dehydration reaction or the reactor.
  • a part of the catalyst to be produced may be taken out from the reactor or the catalyst receiver to be produced, and the part of the catalyst to be produced may be directly or after taking heat to reduce the temperature and then returned to the catalyst.
  • the reactor, or the catalyst mixer fed to the lower portion of the reactor, is mixed with the regenerated catalyst and returned to the reactor for reaction.
  • a part of the catalyst to be produced may be taken out from the reactor or the catalyst receiver to be produced, and the part of the catalyst to be produced may be directly or after taking heat to lower the temperature.
  • a portion of the catalyst to be produced which is taken from the reactor or the catalyst receiver to be produced may be subjected to heat extraction by an external heat extractor to lower the temperature.
  • the external heat extractor is well known to those skilled in the art and may be internally
  • a heat take-up device such as a coil, a bend, or the like is provided to reduce the temperature of the catalyst to be produced flowing therethrough.
  • the catalyst mixer may be connected to the reactor, preferably vertically, for input to one or more of the mixed heat regenerated catalyst in the reactor, the regenerated catalyst after heating, and the catalyst to be produced.
  • the temperature of the catalyst mixing zone may be 200 to 600 ° C, preferably 300 to 500 ° C, and the pressure is 0.5 to 10 MPa.
  • the total carbon content of the catalyst entering the reactor (feed zone) and/or the further reactor (feed zone) may be from 3 to 25% by weight, preferably from 6 to 15% by weight.
  • the catalyst entering the reactor or the further reactor may come from
  • the biocatalyst feeder may also be from the catalyst receiver to be produced and/or the reactor, wherein the catalyst from the regenerated catalyst feeder may be a regenerated catalyst or a regenerated catalyst and a catalyst to be produced. Mix the catalyst.
  • the low-carbon olefin-rich oil and gas can be separated by a product separation and recovery system to obtain a part of C 4 + hydrocarbons, and in order to increase the yield of low-carbon olefins,
  • the C 4 + hydrocarbons are fed to the further reactor for the further reaction to crack the C 4 + hydrocarbons into lower olefins.
  • the regenerated catalyst in the regenerated catalyst feeder may be sent to the further reactor to carry out the further reaction with the C 4 + hydrocarbons, and further obtained.
  • the low carbon olefin-rich oil and gas and the further spent catalyst are fed into the oil separation zone of the reactor; wherein the oil separation zone fed to the reactor is further rich in low
  • the oil and gas of the carbon olefin and the further spent catalyst may be separated together with the low-carbon olefin-rich hydrocarbon oil and the spent catalyst produced in the reactor.
  • the catalyst to be produced in the reactor may be fed to the further reactor to contact the C 4 + hydrocarbons and carry out the further reaction, resulting in further enrichment.
  • the low carbon olefin-containing oil and gas and the further spent catalyst may be separated in the further reactor, and the separated further low-carbon olefin-rich oil and gas is sent to the product separation and recovery system, and the separated further The spent catalyst is sent to the spent catalyst receiver.
  • the catalyst to be produced which is stripped by the stripping section of the reactor may be sent to the further reactor and the C 4 + hydrocarbons and the further Reacting, the resulting lower olefin-rich oil and gas and further spent catalyst are fed into the oil separation zone of the reactor; wherein, the reactor is further enriched in light olefins
  • the oil and gas and further spent catalyst can be separated together with the low carbon olefin-rich oil and gas and the spent catalyst produced in the reactor.
  • the regenerated catalyst in the regenerator can be directly fed into the further reactor to contact the C 4 + hydrocarbons and carry out the further reaction to obtain further richness.
  • a low carbon olefin-containing oil and gas and a further spent catalyst separating the further low-carbon olefin-rich oil and gas and the further spent catalyst in the further reactor, the further separation after separation
  • the low-carbon olefin-rich oil and gas is fed to the product separation and recovery system, and the further spent catalyst can be directly fed into the regenerator for regeneration.
  • the reaction of dehydration to olefins in the reactor and the further reaction in the further reactor may employ reaction conditions known to those skilled in the art capable of producing lower olefins, and the two The reaction can be carried out under substantially the same reaction conditions, or different reaction conditions can be employed. Since the reaction feedstock in the further reactor is not exactly the same as the reaction feedstock in the reactor, it is preferred to employ the further step different from the reactor depending on the feedstock conditions in the further reactor.
  • the reaction conditions for the reaction which is skilled in the art can be appreciated, wherein said primary reaction may be a further C 4 + cracking reactions of hydrocarbons.
  • the reaction conditions in the two reactors can be selected, for example, in the range of 200 to 700 ° C, preferably 250 to 600 ° C; and the reaction pressure may be 0.5 to 10 MPa, preferably It is 1-3.5 MPa.
  • the low carbon olefin-rich oil and gas and the catalyst to be produced can be separated by a filter.
  • further low carbon olefin-rich oil and gas and further spent catalyst may be separated by a filter.
  • the (further) regenerated catalyst and the flue gas may be separated by a filter.
  • the filter may be prepared using a porous material, for example, may be selected from a metal sintered porous material and/or a ceramic porous material; the filter may have a filtration precision of 2 ⁇ m particles of 99.9%, preferably, the filtration The 1.2 ⁇ m particle filtration accuracy of the device can reach 99.9%.
  • the filter can be backflushed using backflush to clean the filter cake.
  • the back blowing gas may be one or more selected from the group consisting of a hydrocarbon-containing gas, dry gas, nitrogen gas, and water vapor.
  • the manufacturing method may further comprise: an unreacted complete oxygenate raw material (including various oxygen-containing compounds newly formed in the dehydration reaction, especially two Methyl ether) is recycled to the step of the dehydration reaction, thereby achieving full utilization of the reaction raw materials.
  • an unreacted complete oxygenate raw material including various oxygen-containing compounds newly formed in the dehydration reaction, especially two Methyl ether
  • At least a portion of the catalyst to be produced and/or at least a portion of the further catalyst to be produced may be conveniently delivered to the regeneration reaction by one or more (preferably one or two) catalyst hoppers, and/or Recycling at least a portion of the regenerated catalyst and/or at least a portion of the further regenerated catalyst to the dehydration reaction and/or the further reaction.
  • the catalyst hopper is sometimes referred to as a lock hopper.
  • the catalyst hopper allows the catalyst to be safely and efficiently transferred from the high pressure hydrocarbon environment of the reactor to the low pressure oxygen environment of the regenerator, and from the low pressure oxygen environment of the regenerator to the high pressure hydrocarbon environment of the reactor. . That is to say, by using the catalyst hopper, on the one hand, the reactor hydrocarbon atmosphere can be well isolated from the oxygen-containing atmosphere of the scorch regeneration of the regenerator, ensuring the safety of the process of the invention, and on the other hand being flexible.
  • the operating pressure of the reactor and the regenerator is regulated, in particular, the operating pressure of the reactor can be increased without increasing the operating pressure of the regenerator to increase the throughput of the apparatus.
  • the catalyst hopper of the present invention is such that the same stream can be passed between different atmospheres (e.g., an oxidizing atmosphere and a hydrocarbon atmosphere) and/or a different pressure environment (e.g., from high pressure to low pressure, or vice versa).
  • atmospheres e.g., an oxidizing atmosphere and a hydrocarbon atmosphere
  • a different pressure environment e.g., from high pressure to low pressure, or vice versa.
  • the step of transporting the catalyst from the reactor (high pressure hydrocarbon environment) to the regenerator (low pressure oxygen environment) through the catalyst hopper may include: 1. Purging the residual oxygen in the evacuated catalyst hopper to the regenerator using hot nitrogen gas. 2. Purging nitrogen from the catalyst hopper with dry gas; 3. Pressurizing the vented catalyst hopper with dry gas; 4. Filling the spent catalyst from the catalyst receiver to be emptied 5.
  • the step of circulating the catalyst from the regenerator (low pressure oxygen environment) to the reactor (high pressure hydrocarbon environment) through the catalyst hopper may include: 1. Purging oxygen from the catalyst hopper filled with the regenerated catalyst into the regenerator using hot nitrogen; 2. Purging nitrogen from the catalyst hopper with dry gas; 3. Pressurizing the filled catalyst hopper with dry gas; 4.
  • the regenerated catalyst feeder and the catalyst circulation line to be produced function to transport the catalyst more continuously to the reactor.
  • the inventors of the present invention have found that the output of the catalyst to be produced in the regenerator and the output of the regenerated catalyst can also be batched.
  • the catalyst hopper delivers the catalyst to be regenerated to the regenerator or the regenerator delivers the regenerated catalyst to the catalyst hopper, It is possible to rely on gravity between the regenerator and the catalyst hopper or to create a pressure differential through the lift line without the need to provide a spent catalyst feeder or a regenerated catalyst receiver.
  • the reaction conditions for regeneration are well known to those skilled in the art, for example, the reaction conditions of the regeneration reaction include: a reaction temperature of 450 to 850 ° C, preferably 550 to 700 ° C; and a reaction pressure of 0.1 to 0.5 MPa. Preferably, it is 0.15-0.3 MPa, such as atmospheric pressure; an oxygen-containing atmosphere.
  • the oxygen-containing atmosphere may be air diluted with air or nitrogen or an oxygen-rich gas as a fluidization medium.
  • the present invention in the case where the size and the number of reactors for carrying out the dehydration reaction are maintained; in other words, when the capacity is upgraded based on the existing reactor or reactor scale, by the present invention
  • By increasing the reaction pressure and the weight hourly space velocity of the reactor within a specific range it is possible to greatly increase the treatment amount of the oxygenate raw material of the reactor and correspondingly increase the yield of the low carbon olefin.
  • the yield increase of the low-carbon olefin can reach 50%, preferably 100%, more preferably 150%, 200%, 500% or 790%, and even 1000% or more in the most preferred case of the present invention. .
  • the yield of the oxygenate feedstock in the reactor or the reaction unit by increasing the yield of the oxygenate feedstock in the reactor or the reaction unit, on the basis of maintaining substantially no or slightly higher yields of the lower olefins than in the prior art, Through production, the yield of low-carbon olefins is increased. Therefore, at the expense of sacrificing the yield of low-carbon olefins (such as a reduction of more than 20%), the yield of low-carbon olefins can be increased by simply increasing the amount or throughput of the oxygenate feedstock in the reactor or reactor. In comparison, the yield increase of the low-carbon olefins in the present invention is significantly higher.
  • the yield of the lower olefin can be maintained in a ratio comparable to the prior art. The level is even higher, such as generally 60%-95% or 78%-95%.
  • the size of the reactor or reactor can be significantly reduced and compared to the prior art while ensuring a predetermined low olefin production is achieved.
  • FIG. 1 is a schematic flow chart showing a process for producing a light olefin by using an oxygenate raw material according to a first embodiment of the present invention.
  • the oxygenate feedstock enters the riser reactor 1 of the riser type reactor from the feed line 24, and is contacted with the catalyst from the line 23 lifted by the pre-lift line 28 to carry out dehydration to olefin reaction.
  • the internal heat extractor 2 takes out the excess heat of the dense phase bed reactor 3, and the reaction oil and gas continues to react in the dense phase bed reactor 3, and the resulting oil and gas rich in low carbon olefins is generated.
  • the catalyst enters the settling zone 5, and the settled catalyst is returned to the dense phase bed reactor 3.
  • the low-carbon olefin-rich oil and gas and the carried catalyst fine powder are filtered through the filter 6, and the oil and gas rich in low-carbon olefins are pipelined. 25 is sent to the product separation and recovery system (not shown), and the filtered fine powder of the catalyst to be returned is returned to the dense phase bed reactor for stripping through the stripping section 4, and part of the catalyst to be produced after stripping is sent to the pipeline 16
  • the raw catalyst receiver 8 is charged to the regenerated catalyst feeder 12 by taking heat from the external heat extractor 13.
  • the catalyst to be produced from the catalyst receiver 8 to be fed enters the catalyst hopper 9 via line 17, is depressurized, is sent to the catalyst feeder 10 to be produced via line 18, and is fed to the regenerator 7 via line 19 and the main wind from line 27.
  • the countercurrent contact burns and regenerates, and the excess heat is taken out through the internal heat extractor 15 (the heat can be taken by the hot material flow rate and the inner heat extractor 15 is buried in the dense bed height), and the flue gas is sent to the subsequent energy recovery via the pipeline 26.
  • the regenerated catalyst is sent to the regenerated catalyst receiver 11 via the line 20, and the excess heat of the regenerated catalyst is taken out through the internal heat extractor 14, and the regenerated catalyst after the heat is sent to the catalyst hopper 9 via the line 21, After the pressure is sent to the regenerated catalyst feeder 12 via line 22 and mixed with the catalyst to be produced from the heat extractor 13, it is sent to the pre-lift section of the riser reactor 1 via line 23.
  • FIG. 2 is a schematic flow chart showing a process for producing a light olefin by using an oxygenate raw material according to a second embodiment of the present invention.
  • the catalyst from line 223 is mixed with the catalyst to be produced from line 213 in catalyst mixer 211, and the riser reactor 201 is fed to the riser type reactor after pre-lifting line 219 is lifted.
  • the oxygenate feedstock enters the riser reactor 201 via the feed line 224, contacts the catalyst of the mixer 211, and undergoes a dehydration to olefin reaction. After the reaction, the oil and gas continue to react in the riser and the distribution plate 202, and then enters the dense phase.
  • the bed reactor 203, the chill medium from the chill medium line 220 enters the riser reactor 201 to control the reaction temperature, the unconverted feedstock continues to contact with the catalyst in the dense bed reactor 203, and the excess heat of reaction is internally heated.
  • the oil and gas rich in low-carbon olefins and the catalyst to be produced are settled into the settling zone 205, and the catalyst to be introduced enters the dense-phase bed reactor 203.
  • the oil and gas rich in low-carbon olefins and the fine powder of the catalyst to be produced are filtered through the filter.
  • the oil rich in low-carbon olefin is sent to the product separation and recovery system (not shown) via line 225, and the filtered fine powder of the catalyst to be precipitated is returned to the dense bed reaction. 203.
  • the catalyst to be produced is stripped in the stripping section 204, the partially stripped catalyst is returned to the catalyst mixer 211 via line 213, and the other partially stripped catalyst is sent to the catalyst receiver 208 via line 216.
  • the catalyst to be produced from the catalyst receiver 208 to be activated enters the catalyst hopper 209 via line 217, is depressurized, is sent to the regenerator 207 via line 221, and is counter-currently contacted with the main stream from line 227 for charring regeneration.
  • the flue gas is sent via line 226.
  • Subsequent energy recovery and purification system (not shown), the regenerated catalyst is sent to the regenerated catalyst receiver 210 via line 215, and the excess heat of the regenerated catalyst is taken out through the internal heat extractor 214. After the heat is taken, the regenerated catalyst is sent to the catalyst hopper 209 via line 218. After being boosted, it is sent to the regenerated catalyst feeder 212 via line 222 and sent to the catalyst mixer 211 via line 223.
  • Fig. 3 is a flow chart showing a process for producing a light olefin by using an oxygenate raw material according to a third embodiment of the present invention.
  • the oxygenate feedstock enters the riser reactor 301 of the riser type reactor from the feed line 324, reacts with the catalyst from the line 323, and undergoes a dehydration to olefin reaction, and the reacted oil and gas enters the expanded diameter.
  • the riser 302 continues the reaction, and the reacted oil and gas enters the dense-phase bed reactor 303 for reaction, and the obtained low-carbon olefin-rich oil and gas and the spent catalyst enter the settling zone 305 to settle, and the settled catalyst is returned to the densely packed catalyst.
  • the phase bed reactor, the low-carbon olefin-rich oil and gas and the carried catalyst fine powder are filtered through a filter 306, and the low-carbon olefin-rich oil is sent to the product separation and recovery system (not shown) through the line 325, and filtered.
  • the fine powder of the spent catalyst is returned to the stripping section 304 gas of the dense phase bed reactor, and some of the stripped catalyst after stripping is sent to the raw catalyst receiver 308 via the line 316, and another part of the catalyst to be produced is taken out.
  • the heater 313 takes heat and is sent to the regenerated catalyst feeder 312.
  • the catalyst to be produced from the catalyst receiver 308 to be activated enters the catalyst hopper 309 via line 317, is depressurized, and is sent via line 318 to the catalyst feeder 310 to be fed, and is sent via line 319 to the regenerator 307 and the main stream from line 327.
  • the countercurrent contact burns and regenerates, and the excess heat is taken out by the internal heat extractor 315 (the heat can be taken by the heat take-off flow and the internal heat extractor 315 is buried in the dense bed height), and the flue gas is sent to the subsequent energy recovery via the line 326.
  • the regenerated catalyst is sent to the regenerated catalyst receiver 311 via the line 320, and the excess heat of the regenerated catalyst is taken out through the internal heat extractor 314. After the heat is taken, the regenerated catalyst is sent to the catalyst hopper 309 via the line 321 to boost the pressure. After being fed to the regenerated catalyst feeder 312 via line 322 and mixed with the catalyst to be produced from the external heat extractor 313, it is sent to the riser reactor 301 and the further riser type reactor 330 via lines 323, 332, respectively.
  • the catalyst from line 332 enters the further riser reactor 330 pre-stage section, through from the pre-stage line into the further riser reactor 330 after the pre-stage 328 of the media to enhance, the product separation and recovery system of the separated C 4 +
  • the hydrocarbons are further reacted via the feed line 329 into a further riser reactor 330 in contact with the catalyst, and the resulting low carbon olefin-rich oil is passed via line 331 into a dense bed fluidized bed 303.
  • Fig. 4 is a flow chart showing a process for producing a light olefin by using an oxygenate raw material according to a fourth embodiment of the present invention.
  • the catalyst from line 423 is mixed with the catalyst to be produced from the external heat extractor 413 in the catalyst mixer 411, and is lifted to the riser of the riser type reactor after the pre-lift line 419 is lifted by the pre-lift line 419.
  • the reactor 401, the oxygenate feedstock enters the riser reactor 401 via the feed line 424, reacts with the catalyst of the catalyst mixer 411 to cause dehydration to olefin reaction, and the reacted product and catalyst pass through the inner riser and the distribution plate 402.
  • the chilling medium from the chilling medium line 420 enters the riser reactor 401 to control the reaction temperature, and the unconverted raw material continues to contact with the catalyst in the dense phase bed reactor 403, and the excess reaction heat is contained therein.
  • the heat extractor 415 takes out, and the generated low-carbon olefin-rich oil and gas and the spent catalyst enter the settling zone 405 to settle, and the raw catalyst enters the stripping section 404, and the low-carbon olefin-containing oil and gas and the carried catalyst powder are filtered through the filter.
  • the oil containing low olefins is sent via line 425 to a product separation and recovery system (not shown), and the filtered catalyst fines are settled back to the dense bed reactor 403.
  • the catalyst to be produced is stripped in the stripping section 404, and the partially stripped catalyst is returned to the catalyst mixer 411 via the internal heat extractor 413, and the other portion is sent to the first reaction zone of the further riser type reactor via line 430. 431.
  • the C 4 + hydrocarbons separated by the product separation and recovery system are sent via line 429 to a first reaction zone 431 of a further riser reactor, which is contacted with a catalyst from line 430 for further reaction, resulting in a low enriched olefin
  • the oil and gas and the catalyst enter the second reaction zone 432 to continue the reaction, and are sent to the settling zone 435 via the reduced diameter 433.
  • the low carbon olefin-rich oil and gas and the carried catalyst fine powder are filtered by the filter 436, and the oil is sent through the pipeline 437.
  • the separation recovery system (not shown) is stripped by the stripping section 434 and sent to the catalyst recycle system via lines 438,416.
  • the catalyst to be produced from the catalyst receiver 408 to be activated enters the catalyst hopper 409 via line 417, is depressurized, is sent to the regenerator 407 via line 421, and is countercurrently contacted with the main stream from line 427 for charring regeneration.
  • the flue gas is sent via line 426.
  • Subsequent energy recovery and purification system (not shown), the regenerated catalyst is sent to the regenerated catalyst receiver 410 via line 415, and the excess heat of the regenerated catalyst is taken out through the internal heat extractor 414. After the heat is taken, the regenerated catalyst is sent to the catalyst hopper 409 via line 418. After boosting, it is sent to the regenerated catalyst feeder 412 via line 422 and then to the catalyst mixer 411 via line 423.
  • Fig. 5 is a schematic flow chart showing a process for producing a light olefin by using an oxygenate raw material according to a fifth embodiment of the present invention.
  • the oxygenate feedstock enters the riser reactor 501 of the riser reactor via feed line 524, and reacts with the catalyst from line 523 to carry out the dehydration to olefin reaction.
  • the reacted product and catalyst are passed through.
  • the inner riser and the quick break 502 enter the dense phase bed reactor 503, and the chill medium from the chill medium line 528 enters the riser reactor 501 to control the reaction temperature, and the unconverted feedstock continues to contact the catalyst in the dense bed reactor 503.
  • the obtained low-carbon olefin-rich oil and gas and the catalyst to be produced enter the settling zone 505, and the low-carbon olefin-rich oil and gas and the carried catalyst fine powder are filtered by the filter 506, and the low-carbon olefin-rich oil and gas pipeline is filled.
  • the 525 is fed to a product separation and recovery system (not shown), and the filtered fine powder of the catalyst to be precipitated is settled into the dense phase bed reactor 503.
  • the partially stripped catalyst is sent to the catalyst receiver 508 via line 516, and the partially stripped catalyst is sent to the further riser via line 535.
  • the remaining catalyst to be generated is sent to the external heat extractor 513 for heat and sent to the inner riser and the fast minute 502.
  • the first reaction zone product separation and recovery system was separated C 4 + hydrocarbons is fed further riser reactor via line 533 530 and a reduced diameter and a second reaction zone 531, and the pre-lifting line 532 from line 535
  • the pre-elevation gas is further reacted with the elevated catalyst, and the resulting low-carbon olefin-rich oil and gas and catalyst are sent to the dense bed 503 via line 534.
  • a portion of the catalyst to be produced from the catalyst receiver 508 is heated by the external heat extractor 529 and sent to the regenerated catalyst feeder 512.
  • the other portion enters the catalyst hopper 509 via the line 517, and is stepped down and sent to the reactor via line 518.
  • the catalyst feeder 510 is sent to the regenerator 507 via line 519 and counter-currently contacted with the main air from the line 527 for scorch regeneration.
  • the excess heat is taken out by the internal heat extractor 515, and the flue gas is sent to the subsequent energy recovery and purification system via the line 526. (not shown), the regenerated catalyst is sent to the regenerated catalyst receiver 511 via the line 520, and the excess heat of the regenerated catalyst is taken out by the internal heat extractor 514.
  • the regenerated catalyst is sent to the catalyst hopper 509 via the line 521, and the line is boosted and then passed through the line.
  • 522 is sent to the regenerated catalyst feeder catalyst feeder 512, and the catalyst to be produced from the catalyst 508 to be produced is taken to the regenerated catalyst feeder 512 after being taken up by the heat extractor 529, and the catalyst is mixed with the regenerated catalyst and then passed through the pipeline.
  • 523 is fed to the riser reactor 501.
  • Fig. 6 is a schematic flow chart showing a process for producing a light olefin by using an oxygenate raw material according to a sixth embodiment of the present invention.
  • the catalyst from line 637 is mixed with the catalyst to be produced from the external heat extractor 613 in the catalyst mixer 611, and is first fed to the riser type reactor after being lifted by the pre-lift line 619.
  • the feedstock enters the first reaction zone 601 via the feed line 624, and reacts with the catalyst of the catalyst mixer 611 to cause dehydration to olefins.
  • the reacted product and catalyst enter the second reaction zone 602, and the unconverted raw materials are passed.
  • the catalyst is continuously contacted with the catalyst to obtain a low-carbon olefin-rich oil and gas and a catalyst to be produced, which are reduced in diameter and fast-divided into the settling zone 605, and the low-carbon olefin-rich oil and gas and the carried catalyst are fine.
  • the oil rich in low-carbon olefin is sent to the product separation and recovery system (not shown) via the line 625, and the filtered fine powder of the catalyst to be settled into the stripping section 604.
  • Part of the spent catalyst after stripping through the stripping section is sent to the waiting line via line 616.
  • the catalyst receiver 608 is partially fed to the external heat extractor 613 and taken to the catalyst mixer 611 via line 624. The remaining portion is heated by the external heat extractor 612 and sent to the regenerated catalyst feeder 636.
  • the catalyst to be produced from the catalyst receiver 608 to be fed enters the catalyst hopper 609 via line 617, is depressurized, is sent to the regenerator 607 via line 621, and is counter-currently contacted with the main air from the line 622 to be regenerated, and the flue gas is sent through the line 620.
  • Subsequent energy recovery and purification system (not shown), excess heat is taken out by the internal heat extractor 615, and the regenerated catalyst is sent to the regenerated catalyst receiver 610.
  • the excess heat of the regenerated catalyst is taken out by the internal heat extractor 614, and the regenerated catalyst is taken after the heat is taken.
  • Line 618 is fed to catalyst hopper 609, boosted and sent via line 623 to regenerated catalyst feeder 636 where it is mixed with the spent catalyst from external heat taker 612 and sent to catalyst mixer 611 via line 637.
  • the oil and gas and the catalyst to be produced rich in low-carbon olefins are separated, and the separated oil is sedimented in the sedimentation zone 629, and then filtered through the filter 631.
  • the filtered oil is sent to the subsequent separation system via line 632 (not shown), and the catalyst is discharged.
  • the stripping section 630 is stripped, and the catalyst to be produced after stripping is sent to the regenerator 607 via line 634 for regeneration.
  • Fig. 7 is a flow chart showing a process for producing a light olefin by using an oxygenate raw material according to a seventh embodiment of the present invention.
  • the oxygenate feedstock and diluent enter the fluidized bed reactor 701 from feed line 702, contact with the catalyst from line 723 for dehydration to olefins, and internal heat extractor 713 to remove the fluidized bed reaction.
  • the oil rich in low-carbon olefin is sent to the product separation and recovery system (not shown) via the line 703, and the filtered fine powder of the catalyst to be settled is returned to the fluidized bed reactor, and is partially treated.
  • the biocatalyst is sent via line 16 to the spent catalyst receiver 8 and stripped.
  • the spent catalyst from the spent catalyst receiver 708 enters the catalyst hopper 709 via line 717, is depressurized, and is sent via line 718 to the spent catalyst feeder 710 where it is sent to the regenerator 707 and the main stream from line 724.
  • the countercurrent contact is burnt and regenerated, and the excess heat is taken out through the internal heat extractor 715, and the flue gas is sent to the subsequent energy recovery and purification system via the pipeline 704.
  • the regenerated catalyst is sent to the regenerated catalyst receiver 711 via the line 720, and the excess heat of the regenerated catalyst is taken out by the internal heat extractor 714, and the regenerated catalyst after the heat is sent to the catalyst hopper 709 via the line 721, and is pressurized.
  • Line 722 is sent to regenerated catalyst feeder 712 and stripped and sent to fluidized bed reactor 701 via line 723.
  • Examples 1-6 were carried out according to the process shown in Fig. 1 (the examples of the present invention and the comparative examples were tested according to the process shown in Fig. 1 using the same reactor), and the reaction conditions were basically the same, and only the reaction pressure and the weight hourly space velocity were changed.
  • the starting materials, catalysts, reaction conditions and product yields are shown in Table 1.
  • Examples 7-8 were carried out according to the process shown in Fig. 1. Compared with Example 4, when Examples 7-8 were used to increase the reaction pressure, the weight hourly space velocity was not increased correspondingly, and other operating conditions were substantially equivalent.
  • the reaction materials, catalysts, reaction conditions, and product yields are shown in Table 2.
  • Example 4 Example 7 and Example 8 that, compared with Example 4, when the other reaction conditions are substantially equivalent, if the reaction pressure is simply increased without correspondingly increasing the weight hourly space velocity, the low carbon olefin Both yield and yield decreased, the low carbon olefin yield decreased from 84.3% of Example 4 to 82.9% of Example 8, and the low carbon olefin yield decreased from 3.73 kg/h of Example 4 to Example 8. 2.64kg/h.
  • Example 9 was carried out according to the process shown in Fig. 1. Compared with Example 3, the operating conditions were substantially equivalent, and only the carbon content of the catalyst at the inlet of the reactor was changed (the carbon content of the catalyst at the inlet of the reactor means that the inlet of the reactor was not reacted with the raw material. Pre-catalyst carbon content).
  • the reaction materials, catalysts, reaction conditions, and product yields are shown in Table 2.
  • Example 3 and Example 9 It can be seen from Example 3 and Example 9 that, compared with Example 3, when the reaction pressure and the weight hourly space velocity are substantially equivalent, the reactor inlet catalytic carbon content is reduced from 7.3% to 4.5%, and the light olefin is reduced.
  • the yield decreased from 3.84kg/h to 3.58kg/h; the yield of low-carbon olefins was 84.9%. It is reduced to 84.3%; the mass ratio of ethylene to propylene becomes larger.
  • Example 10 was carried out in accordance with the process shown in Figure 3.
  • the starting materials, catalysts, reaction conditions and product yields are listed in Table 2.
  • the yield of the low carbon olefin was 93.3%, and the yield of the low carbon olefin was 4.03 kg/h.
  • Example 11 was carried out according to the process shown in Fig. 7, and the raw material of the reaction raw material was ethanol.
  • the catalyst, the reaction conditions and the product yield are shown in Table 2.
  • Example 11 It can be seen from Example 11 that the yield of low carbon olefin is 82.8%, and the yield of low carbon olefin is 3.26 kg/h.
  • Example 12 was carried out in accordance with the process shown in Figure 2, and the starting materials, catalysts, reaction conditions and product yields are shown in Table 4.
  • Example 12 is a production scheme for producing a gasoline product while increasing the production of low-carbon olefins. As can be seen from Example 12, the propylene yield was 65.9%, the gasoline yield was 25.3%, the propylene yield was 1.98 kg/h, and the gasoline yield was 0.76 kg/h.
  • Comparative Examples 1-6 The reactor, raw materials, and catalysts used in Comparative Examples 1-6 were identical to those in Examples 1-10, and were carried out in accordance with the process shown in FIG. Compared with Examples 1-10, Comparative Example 1 is a conventional methanol-made low-carbon olefin condition, and the carbon content of the reactor inlet catalyst is significantly lower, and the reaction pressure and weight hourly space velocity are significantly lower than the present invention; Comparative Example 2-6 Only the reaction pressure and the weight hourly space velocity were changed, and other operating conditions were roughly equivalent to those of Examples 1-10. Comparative Examples 1-6 The starting materials, catalysts, reaction conditions and product yields are shown in Table 3.
  • Comparative Example 1 is a conventional methanol-made low-carbon olefin operating condition, and Comparative Example 2 differs only from the carbon content of the reactor inlet catalyst of Comparative Example 1. Compared with Comparative Example 1, the carbon content of the reactor inlet catalyst increased from 1.5% of Comparative Example 1 to 7.2% of Comparative Example 2, and the yield of lower olefins decreased from 0.47 kg/h of Comparative Example 1 to Comparative Example 2 0.43 kg/h; the conversion rate was reduced from 100% of Comparative Example 1 to 80.2% of Comparative Example 2.
  • Comparative Example 2 Compared with Examples 1-8, the carbon content and the like of the reactor inlet catalyst were substantially equal except for the reaction pressure and the weight hourly space velocity. Compared with Comparative Example 2, the yields of the low carbon olefins of Examples 1-8 were substantially equal or increased, and the yield of low carbon olefins was greatly increased, for example, low. The carbon olefin yield increased from 79.2% of Comparative Example 2 to 84.9% of Example 3; the yield of low carbon olefin increased from 0.43 kg/h of Comparative Example 2 to 3.84 kg/h of Example 3, with an increase of up to 793.02%.
  • Comparative Example 3 and Example 3 except for the weight hourly space velocity, the remaining operating conditions were approximately equivalent.
  • the comparative example 3 weight hourly space velocity was significantly higher than that of Example 3.
  • the yield of low carbon olefin increased from 0.46 kg/h of Comparative Example 3 to 3.84 kg/h of Example 3, with an increase of up to 734.78%; the yield of low carbon olefins from the comparative example 82.9% of 3 increased to 84.9% of Example 3.
  • Comparative Example 4 and Example 2 except for the reaction pressure were approximately equivalent.
  • the reaction pressure of Comparative Example 4 was significantly lower than that of Example 2.
  • the yield of light olefins increased from 0.44 kg/h of Comparative Example 4 to 2.54 kg/h of Example 2, with an increase of up to 477.27%; the yield of low carbon olefins was from 4 79.4% increased to 84.4% of Example 2.
  • Comparative Example 5 and Example 2 except for the weight hourly space velocity, the remaining operating conditions were approximately equivalent.
  • the comparative example 5 weight hourly space velocity was significantly higher than that of Example 2.
  • the yield of low carbon olefin increased from 0.43 kg/h of Comparative Example 5 to 2.54 kg/h of Example 2, with an increase of up to 490.70%; the yield of low carbon olefins from the comparative example 75.3% of 5 increased to 84.4% of Example 2.
  • Comparative Example 6 and Example 1 except for the reaction pressure were substantially equivalent.
  • the reaction pressure of Comparative Example 6 was significantly higher than that of Example 1.
  • the use of the present invention increased the yield of light olefins from 0.33 kg/h of Comparative Example 6 to 1.11 kg/h of Example 1, with an increase of up to 236.36%; the yield of low carbon olefins from the comparative example 33.4% of 6 increased to 83.7% of Example 1.
  • Example I was carried out according to the process shown in Figure 1. The starting materials, catalysts, reaction conditions and product yields are listed in Table I.
  • Example II was carried out according to the process shown in Figure 3. The starting materials, catalysts, reaction conditions and product yields are listed in Table I.
  • Example III was carried out in accordance with the procedure shown in Figure 2, and the starting materials, catalysts, reaction conditions, and product yields are listed in Table II.
  • the yield of ethylene and propylene can be higher than the level of the existing industrial process by the method of the present invention; as can be seen from Table II, the yield of propylene and gasoline is 65.9% by the method of the present invention, respectively. And 25.3%, higher than the existing industrial process level, and since the reaction system pressure of the present invention is higher than that of the existing industrial device, the raw material processing amount of the reaction system of the present invention is higher than that of the existing industrial device under the same other operating conditions. .
  • Example III Reaction raw material (oxygen compound raw material mass fraction > 98wt%) Industrial methyl ether Catalyst (Qilu Catalyst) ZSM-5 molecular sieve catalyst Raising the reaction conditions of the tubular reactor Temperature, °C 500 Pressure, MPa 1.5 Scorch regeneration condition of regenerator Regeneration pressure, MPa 0.2 Regeneration temperature, °C 650 Regenerated catalyst quantification, weight % 0.4 Product yield Ethylene, % 4.1 Propylene, % 65.9 gasoline,% 25.3

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Abstract

本发明公开了一种低碳烯烃的制造方法，其中在通过连续地使含氧化合物原料与催化剂接触而发生脱水反应以制造低碳烯烃的方法中，使所述脱水反应的反应压力P为1-2MPa，所述脱水反应的重时空速H为15-50h-1。根据本发明的低碳烯烃制造方法，不仅操作过程简单、连续，而且具有投资省、低碳烯烃增产幅度大和安全性高的特点。

Description

一种低碳烯烃的制造方法 技术领域
本发明涉及一种由含氧化合物原料制造低碳烯烃的方法。更具体而言，本发明涉及一种在由含氧化合物原料制造低碳烯烃的方法中，使低碳烯烃增产的方法。
背景技术
低碳烯烃(C2-C4烯烃)作为基本有机化工原料，在现代石油和化学工业中具有举足轻重的地位。制造低碳烯烃的方法大致可以分为两大类，即传统石油路线和新兴非石油路线。自20世纪10年代以来，世界各国开始致力于研发非石油资源(特别是含氧化合物原料)制造低碳烯烃的路线，取得了一些进展。
含氧化合物脱水反应制造低碳烯烃后除了获得烃类产品外，还副产一定量的水，以甲醇为例，反应后可以获得烃类产品约44％，以乙醇为例，反应后可以获得烃类产品46％。已知的是，含氧化合物原料制造低碳烯烃的反应是分子数量增加的反应，低反应压力有利于化学平衡向生成低碳烯烃方向进行。鉴于此，现有技术在制造低碳烯烃时，为了获得理想的低碳烯烃收率，普遍使用了较低的反应压力。这种低反应压力(一般是0.1-0.3MPa)带来的直接后果是，如果为了增产低碳烯烃而希望提高含氧化合物原料的处理量，现有技术就不得不为此而增加反应器的尺寸或数量，以维持低碳烯烃的收率在可以接受的水平。显然的是，这样会相应增加设备的投资和维护成本。
在现有技术的低碳烯烃制造方法中，为了确保制造方法的连续性，使催化剂在反应器和再生器之间循环流动。为了方便该循环流动，反应器与再生器一般在基本上相同的压力下操作。此时，反应器是烃气氛，再生器是含氧气氛，如果没有将二者很好地隔离，就会存在很大的安全隐患。
另外，现有技术的低碳烯烃生产装置普遍采用了与催化裂化装置相似的旋风分离器，在生产过程中催化剂的自然跑损是无法避免的，尤其是当催化剂中粒度不超过20微米的催化剂细粉增多的时候，这会对后续的产物分离带来不利影响，对催化剂的重复使用也是不利的。
发明内容
本发明的目的是提供一种低碳烯烃的制造方法，该方法克服了现有技术存在的前述缺点，并且能够直接利用现有的反应器简便地实现增产低碳烯烃的目的。
本发明的发明人通过刻苦的研究，出人意料地发现，如果在增加反应压力的同时相应增加含氧化合物原料的重时空速，就可以使低碳烯烃的收率维持在与现有技术相当甚至更高的水平，而不会如现有技术以往所预期的那样降低，其结果是，针对现有的反应器，通过按照本发明的规定增加该反应器的反应压力和重时空速，就可以大幅度地增加该反应器的含氧化合物原料处理量而相应增加低碳烯烃的产量(增产低碳烯烃)。本发明人的这一发现突破了本领域技术人员的常规认识，并基于该发现而完成了本发明。
具体而言，本发明涉及以下方面的内容。
1.一种低碳烯烃的制造方法(或增产方法)，其特征在于，在通过连续地使含氧化合物原料与催化剂接触而发生脱水反应以制造低碳烯烃的方法中，使所述脱水反应的反应压力P为0.5-10MPa，优选0.75-3.5MPa，更优选0.8-3MPa，最优选1-2MPa，所述脱水反应的重时空速H为7-250h-1，优选8-150h-1，更优选10-100h-1，更优选15-80h-1，最优选15-50h-1。
2.根据前述任一方面所述的制造方法，其中在进行所述脱水反应时，严格增函数H＝f(P)成立，其中P(单位是MPa)属于区间[0.55，10.0]，优选属于区间[0.75，3.5]，更优选属于区间[0.8，3.0]，最优选属于区间[1.0，2.0]，H(单位是h-1)属于区间[7，250]，优选属于区间[8，150]，更优选属于区间[10，100]，更优选属于区间[15，80]，最优选属于区间[15，50]。
3.根据前述任一方面所述的制造方法，包括以下步骤：
连续地使所述含氧化合物原料与所述催化剂接触而发生所述脱水反应，获得富含低碳烯烃的油气和待生催化剂，
将至少一部分所述待生催化剂输送至再生反应，获得再生催化剂，和
将至少一部分所述再生催化剂循环至所述脱水反应，
其中所述脱水反应的反应压力P比所述再生反应的再生压力至少高0.35MPa，优选至少高0.4MPa、至少高0.5MPa、至少高0.6MPa、至少高0.7MPa、至少高0.8MPa、至少高0.9MPa、至少高1.0MPa、至少高1.1MPa、至少高1.2MPa、至少高1.3MPa、至少高1.4MPa、至少高1.5MPa、至少高1.6MPa、至少高1.7MPa、至少高1.8MPa、至少高1.9MPa或者至少高2.0MPa。
4.根据前述任一方面所述的制造方法，还包括分离所述富含低碳烯烃的油气而获得C4 +烃类的步骤，并且任选还包括以下步骤：
连续地使所述C4 +烃类与进一步的催化剂接触而发生进一步反应，获得进一步的富含低碳烯烃的油气和进一步的待生催化剂，
将至少一部分所述进一步的待生催化剂输送至所述再生反应，获得进一步的再生催化剂，和
将至少一部分所述再生催化剂和/或至少一部分所述进一步的再生催化剂循环至所述脱水反应和/或所述进一步反应。
5.根据前述任一方面所述的制造方法，其中用于进行所述脱水反应的反应器和/或用于进行所述进一步反应的反应器的数目是一个或多个，并且各自独立地选自流化床反应器、密相床反应器、提升管反应器、沸腾床反应器、浆态床反应器、以及这些反应器中两种或更多种的复合形式，优选选自提升管反应器，更优选各自独立地选自等直径提升管反应器、等线速提升管反应器、变径提升管反应器以及提升管复合密相床反应器。
6.根据前述任一方面所述的制造方法，其中所述含氧化合物原料选自醇、醚和酯中的至少一种，优选选自R1-O-R2、R1-OC(＝O)O-R2、R1-C(＝O)O-R2和R1-C(＝O)-R2(其中，R1和R2彼此相同或不同，各自独立地选自氢和C1-6支链或直链烷基，优选各自独立地选自氢和C1-4支链或直链烷基，前提是R1和R2中的至多一个是氢)中的至少一种，更优选选自甲醇、乙醇、二甲醚、二乙醚、甲乙醚、碳酸二甲酯和甲酸甲酯中的至少一种。
7.根据前述任一方面所述的制造方法，其中所述催化剂和所述进一步的催化剂彼此相同或不同，各自独立地选自分子筛催化剂中的至少一种，优选各自独立地选自硅铝磷酸盐分子筛催化剂和硅铝酸盐分子筛催化剂中的至少一种。
8.根据前述任一方面所述的制造方法，其中所述再生反应的反应条件包括：反应温度450-850℃，优选550-700℃；反应压力0.1-0.5MPa，优选0.15-0.3MPa；含氧气氛，优选空气气氛或者氧气气氛。
9.根据前述任一方面所述的制造方法，其中通过过滤器分离出所述待生催化剂和/或所述进一步的待生催化剂和/或所述再生催化剂和/或所述进一步的再生催化剂。
10.根据前述任一方面所述的制造方法，其中通过一个或多个(优选一个或两个)催化剂料斗(9)实现所述输送和所述循环。
11.根据前述任一方面所述的制造方法，还包括将至少一部分所述待生催化剂和/或至少一部分所述进一步的待生催化剂循环至所述脱水反应和/或所述进一步反应的步骤。
12.根据前述任一方面所述的制造方法，其中所述催化剂和/或所述进一步的催化剂的总含碳量为3-25wt％，最优选6-15wt％。
13.根据前述任一方面所述的制造方法，其中在维持用于进行所述脱水反应的反应器的尺寸和数量不变的情况下，该制造方法能够使低碳烯烃的产量提高50％，优选提高100％，更优选提高150％、200％、500％或790％，最优选提高1000％或更高。
14.根据前述任一方面所述的制造方法，还包括将未反应完全的含氧化合物原料循环至所述脱水反应的步骤。
15.根据前述任一方面所述的制造方法，包括以下步骤：
连续地将所述含氧化合物原料在提升管型反应器中与所述催化剂接触进行所述脱水反应，产生富含低碳烯烃的油气和待生催化剂；
使富含烯烃的油气和待生催化剂在油剂分离区进行分离，将分离后的富含烯烃的油气送入产品分离回收系统，将待生催化剂经提升管型反应器中的汽提段汽提后从所述提升管型反应器引出并输送至待生催化剂接收器；
将待生催化剂接收器中的待生催化剂通过催化剂料斗直接输送至再生器，或先通过催化剂料斗输送至待生催化剂进料器后再输送至再生器，并在再生器中在含氧气氛下进行烧焦再生，得到再生催化剂；
将再生催化剂直接输送到催化剂料斗，或先从再生器引出并输送至再生催化剂接收器，然后再输送至催化剂料斗；
将催化剂料斗内的再生催化剂输送至再生催化剂进料器后返回到 所述提升管型反应器中。
技术效果
与现有技术相比，本发明的低碳烯烃制造方法主要具有以下优点。
根据本发明的低碳烯烃制造方法，在增加反应压力的同时相应增加含氧化合物原料的重时空速，在不改变已有反应器或反应装置的尺寸和数量的情况下，可以使低碳烯烃的收率维持在与现有技术相当甚至更高的水平，并最终大幅度(最高可以超过790％)提高低碳烯烃的产量。鉴于此，本发明的低碳烯烃制造方法属于低碳烯烃增产方法，可以应用于已有的低碳烯烃生产装置的改造或产能升级。
根据本发明的低碳烯烃制造方法，在确保达到预定的低碳烯烃产量的同时，与现有技术相比，可以显著降低反应器或反应装置的尺寸和数量，由此降低整个低碳烯烃生产装置的规模和投资成本。鉴于此，本发明的低碳烯烃制造方法是新一代的高产能低碳烯烃制造方法，可以应用于建设与已有的低碳烯烃生产装置相比，装置规模更小、投资成本更低、并且低碳烯烃产量更高的新一代低碳烯烃生产装置。
根据本发明的低碳烯烃制造方法，在使反应器在较高压力下操作的同时，维持再生器在较低的压力下操作，由此降低整个低碳烯烃制造方法和制造装置的复杂度。
根据本发明的低碳烯烃制造方法，反应器的反应压力明显高于再生器的再生压力，由此通过使用压力切换装置(比如闭锁料斗或催化剂料斗)，能够实现反应器的烃气氛与再生器的含氧气氛的完全隔离，以及催化剂循环，由此确保整个制造方法和制造装置的安全运行。
本发明的其他特征和优点将在随后的具体实施方式部分予以详细说明。
附图说明
附图是用来提供对本发明的进一步理解，并且构成说明书的一部分，与下面的具体实施方式一起用于解释本发明，但并不构成对本发明的限制。在附图中：
图1是按本发明的第一种具体实施方式的含氧化合物制取低碳烯烃方法的流程示意图
图2是按本发明的第二种具体实施方式的含氧化合物制取低碳烯烃方法的流程示意图
图3是按本发明的第三种具体实施方式的含氧化合物制取低碳烯烃方法的流程示意图
图4是按本发明的第四种具体实施方式的含氧化合物制取低碳烯烃方法的流程示意图
图5是按本发明的第五种具体实施方式的含氧化合物制取低碳烯烃方法的流程示意图
图6是按本发明的第六种具体实施方式的含氧化合物制取低碳烯烃方法的流程示意图
图7是按本发明的第七种具体实施方式的含氧化合物制取低碳烯烃方法的流程示意图
本发明还可以包括其它具体实施方式，并不限于上述七种。
附图标记说明
1提升管反应器  2内取热器  3密相床反应器  4汽提段
5沉降区  6过滤器  7再生器  8待生催化剂接收器
9催化剂料斗  10待生催化剂进料器  11再生催化剂接收器
12再生催化剂进料器  13外取热器  14内取热器
15内取热器  16管线  17管线  18管线  19管线
20管线  21管线  22管线  23管线  24进料线
25反应产物线  26烟气线  27管线  28预提升线。
201提升管反应器  202内提升管及分布板  203密相床反应器
204汽提段  205沉降区  206过滤器  207再生器
208待生催化剂接收器  209催化剂料斗  210再生催化剂接收器
211催化剂混合器  212再生催化剂进料器  213管线
214内取热器  215内取热器  216管线  217管线
218管线
219预提升线  220激冷介质线  221管线  222管线
223管线
224进料线  225反应产物线  226烟气线  227管线
301提升管反应器  302扩径提升管  303密相床反应器
304汽提段  305沉降区  306过滤器  307再生器
308待生催化剂接收器  309催化剂料斗  310待生催化剂进料器
311再生催化剂接收器  312再生催化剂进料器  313外取热器
314内取热器  315内取热器  316管线  317管线
318管线  319管线  320管线  321管线
322管线
323管线  324进料线  325反应产物线  326烟气线
327管线  328预提升线  329进料线  330另一提升管型反应器
331管线  332管线
401提升管反应器  402内提升管及分布板  403密相床反应器
404汽提段  405沉降区  406过滤器  407再生器
408待生催化剂接收器  409催化剂料斗  410再生催化剂接收器
411催化剂混合器  412再生催化剂进料器  413外取热器
414内取热器  415内取热器  416管线  417管线
418管线
419预提升线  420激冷介质线  421管线  422管线
423管线  424进料线  425反应产物线  426烟气线
427管线  428预提升线  429进料线
430管线
431第一反应区  432第二反应区  433缩径及快分
434汽提段  435沉降区  436过滤器  437管线  438管线
501提升管反应器  502内提升管及快分  503密相床反应器
504汽提段  505沉降区  506过滤器  507再生器
508待生催化剂接收器  509催化剂料斗  510待生催化剂进料器
511再生催化剂接收器  512催化剂进料器
513外取热器  514内取热器  515内取热器
516管线
517管线  518管线  519管线  520管线  521管线
522管线  523管线  524进料线  525反应产物线
526烟气线  527管线  528激冷介质线
529外取热器
530第一反应区  531第二反应区及缩径  532预提升线
533进料线  534管线  535管线
601第一反应区  602第二反应区  603缩径及快分
604汽提段  605沉降区  606过滤器  607再生器
608待生催化剂接收器  609催化剂料斗  610再生催化剂接收器
611催化剂混合器  612外取热器  613外取热器  614内取热器
615内取热器  616管线  617管线  618管线  619预提升线
620烟气线  621管线  622管线  623管线  624进料线
625反应产物线  626预提升线  627另一提升管型反应器
628快分  629沉降区  630汽提段  631过滤器  632反应产物线
633管线  634管线  635进料线  636催化剂进料器637管线
701流化床反应器  702进料线  703反应产物线  704烟气线
705沉降区  706过滤器  707再生器  708待生催化剂接收器
709催化剂料斗  710待生催化剂进料器  711再生催化剂接收器
712再生催化剂进料器  713内取热器  714内取热器
715外取热器  716管线  717管线  718管线  719管线
720管线  721管线  722管线  723管线  724主风
具体实施方式
以下结合附图对本发明的具体实施方式进行详细说明。应当理解的是，此处所描述的具体实施方式仅用于说明和解释本发明，并不用于限制本发明。
在本说明书的上下文中，术语“反应器”和“进一步的反应器”是指二个独立的反应器。本发明C4+烃类指C4及C4以上烃类。
在本说明书的上下文中，术语“低碳烯烃”是指乙烯和丙烯。
在本说明书的上下文中，“低碳烯烃的收率”指的是低碳烯烃的单程收率，而“低碳烯烃的产量”指的是单位时间内单位反应器的低碳烯烃的单程产量，重时空速是指单位时间内通过单位质量催化剂的反应物的质量。
收率＝产品产量/除含氧化合物外烃类产品产量之和×100。
所述除含氧化合物外烃类具体包括氢气、C1及C1以上非含氧烃类。
根据本发明，提供一种低碳烯烃的制造方法，其通过连续地使含氧化合物原料与催化剂接触而发生脱水反应以制造低碳烯烃。
根据本发明，所述制造方法可以包括以下步骤：连续地使所述含氧化合物原料与所述催化剂接触而发生所述脱水反应，获得富含低碳烯烃的油气和待生催化剂，将至少一部分所述待生催化剂输送至再生反应，获得再生催化剂，和将至少一部分所述再生催化剂循环至所述脱水反应。更为具体而言，所述制造方法比如可以包括：连续地将所述含氧化合物原料在反应器(比如提升管型反应器)中与所述催化剂接触进行所述脱水反应，产生富含低碳烯烃的油气和待生催化剂；使富含烯烃的油气和待生催化剂在油剂分离区进行分离，将分离后的富含烯烃的油气送入产品分离回收系统，将待生催化剂经反应器中的汽提段汽提后从所述反应器引出并输送至待生催化剂接收器；将待生催化剂接收器中的待生催化剂通过催化剂料斗直接输送至再生器，或先通过催化剂料斗输送至待生催化剂进料器然后再输送至再生器，并在再生器中在含氧气氛下进行烧焦再生，得到再生催化剂；将再生催化剂从再生器引出并输送至再生催化剂接收器，然后通过催化剂料斗输送至再生催化剂进料器，或再生催化剂直接输送到催化剂料斗，后返回到所述反应器中。
根据本发明，作为所述再生器，可以直接使用本领域常规已知的任何类型，比如流化床再生器或沸腾床再生器，但并不限于此。
根据本发明，所述制造方法还可以包括：从所述反应器或所述待生催化剂接收器中引出一部分待生催化剂；将引出的该部分待生催化剂直接或经取热降低温度后返回到所述反应器中，或者输送到所述反应器下部的催化剂混合器中与再生催化剂混合后返回到所述反应器中；所述引出的该部分待生催化剂的量与通过催化剂料斗输送至再生催化剂进料器中的再生催化剂一起足以维持所述反应器中催化剂的连续运转。
根据本发明，所述制造方法还可以包括：从所述反应器或所述待生催化剂接收器中引出一部分待生催化剂；将引出的该部分待生催化剂直接或经取热降低温度后，输送至所述再生催化剂进料器中与再生催化剂混合后返回到所述反应器中；所述引出的该部分待生催化剂的量与通过催化剂料斗输送至再生催化剂进料器中的再生催化剂一起足 以维持所述反应器中催化剂的连续运转。
根据本发明，所述含氧化合物原料是本领域技术人员所熟知的，可以是选自醇、醚和酯中的至少一种，也可以是其它工业或天然含氧化合物，本发明并没有限制。作为所述含氧化合物原料，优选选自R1-O-R2、R1-OC(＝O)O-R2、R1-C(＝O)O-R2和R1-C(＝O)-R2中的至少一种。其中，R1和R2彼此相同或不同，各自独立地选自氢和C1-6支链或直链烷基，优选各自独立地选自氢和C1-4支链或直链烷基，前提是R1和R2中的至多一个是氢。作为所述含氧化合物原料，更优选选自甲醇、乙醇、二甲醚、二乙醚、甲乙醚、碳酸二甲酯和甲酸甲酯中的至少一种，尤其是甲醇。
根据本发明，在进行所述脱水反应时，有时还需要使用稀释剂。作为所述稀释剂，一般采用水蒸气，也可以采用氢气、甲烷、乙烷、氮气、一氧化碳等。在使用时，所述含氧化合物原料与所述稀释剂的摩尔比一般为40∶1-0.4∶1，优选为为11∶1-0.7∶1，更优选为7∶1-1.3∶1。
根据本发明，所述催化剂可以使用本领域技术人员所熟知的类型。例如，作为所述催化剂，可以为分子筛催化剂，所述分子筛可以为硅铝磷酸盐类分子筛和/或硅铝酸盐分子筛。在此，所述硅铝磷酸盐分子筛可以选自SAPO系列、SRM系列分子筛中的一种或几种，所述硅铝酸盐分子筛可以选自ZSM系列、ZRP系列分子筛中的一种或几种。另外，所述分子筛可以负载选自碱土金属、K、Mg、Ca、Ba、Zr、Ti、Co、Mo、Ni、Pt、Pd、La、Ce、Cu、Fe、B、Si、P、Sn、Pb、Ga、Cr、V、Sc、Ge、Mn、La、Al、Ni、Fe中的一种或几种元素。
根据本发明，所述制造方法还可以包括分离所述富含低碳烯烃的油气而获得C4 +烃类的步骤。
根据本发明，所述制造方法任选还可以包括以下步骤：连续地使所述C4 +烃类与进一步的催化剂接触而发生进一步反应，获得进一步的富含低碳烯烃的油气和进一步的待生催化剂，将至少一部分所述进一步的待生催化剂输送至所述再生反应，获得进一步的再生催化剂，和将至少一部分所述再生催化剂和/或至少一部分所述进一步的再生催化剂循环至所述脱水反应和/或所述进一步反应。比如，所述制造方法可以包括将经所述产品分离回收系统分离得到的C4 +烃类送入进一步的反应器(比如提升管型反应器)进行所述进一步反应。
根据本发明，所述进一步的催化剂与所述催化剂可以相同或不同，可以使用本领域技术人员所熟知的类型。例如，作为所述进一步的催化剂，可以为分子筛催化剂，所述分子筛可以为硅铝磷酸盐类分子筛和/或硅铝酸盐分子筛。在此，所述硅铝磷酸盐分子筛可以选自SAPO系列、SRM系列分子筛中的一种或几种，所述硅铝酸盐分子筛可以选自ZSM系列、ZRP系列分子筛中的一种或几种。另外，所述分子筛可以负载选自碱土金属、K、Mg、Ca、Ba、Zr、Ti、Co、Mo、Ni、Pt、Pd、La、Ce、Cu、Fe、B、Si、P、Sn、Pb、Ga、Cr、V、Sc、Ge、Mn、La、Al、Ni、Fe中的一种或几种元素。
根据本发明，所述制造方法还可以包括：将所述再生催化剂进料器中的再生催化剂送入所述进一步的反应器内与所述C4 +烃类接触并进行所述进一步反应，将得到的进一步的富含低碳烯烃的油气和进一步的待生催化剂一起送入所述反应器的所述油剂分离区中。
根据本发明，所述制造方法还可以包括：将所述反应器内的待生催化剂送入所述进一步的反应器与所述C4 +烃类接触并进行所述进一步反应，得到的进一步的富含低碳烯烃的油气和进一步的待生催化剂在所述进一步的反应器中分离，将分离后的进一步的富含低碳烯烃的油气送入所述产品分离回收系统，将分离后的进一步的待生催化剂送入待生催化剂接收器内。
根据本发明，所述制造方法还可以包括：将所述反应器内的待生催化剂送入所述进一步的反应器与所述C4 +烃类接触并进行所述进一步反应，得到的进一步的富含低碳烯烃的油气和进一步的待生催化剂一起送入所述反应器的所述油剂分离区中。
根据本发明，所述制造方法还可以包括：将再生器内的再生催化剂直接送入所述进一步的反应器内与所述C4 +烃类接触并进行所述进一步反应，得到进一步的富含低碳烯烃的油气和进一步的待生催化剂；使所述进一步的富含低碳烯烃的油气和所述进一步的待生催化剂在所述进一步的反应器中分离，将分离后的进一步的富含低碳烯烃的油气送入所述产品分离回收系统，将所述进一步的待生催化剂直接送入所述再生器内进行再生。
根据本发明，所述反应器和/或所述进一步的反应器的数目是一个或多个，并没有特别的限制。另外，所述反应器和/或所述进一步的反 应器彼此相同或不同，各自独立地选自流化床反应器、密相床反应器、提升管反应器、沸腾床反应器、浆态床反应器、以及这些反应器中两种或更多种的复合形式。优选的是，所述反应器和/或所述进一步的反应器彼此相同或不同，各自独立地选自提升管反应器，更优选各自独立地选自等直径提升管反应器、等线速提升管反应器、变径提升管反应器以及提升管复合密相床反应器。另外，所述提升管型反应器沿垂直方向从下至上还可以设置预提升段、提升管、激冷介质线、扩径提升管、缩径、快分、汽提段、密相段、沉降区、催化剂混合器、过滤器等工业上常用装置，使所述反应器能够连续运行；其中，所述沉降区、过滤器等装置可以构成所述油剂分离区，所述油剂分离区也可以包括其它使待生催化剂和油气分离的装置，本发明并没有限制。根据本发明，所述提升管型反应器的密相床可以不形成密相床，即“零料位”。
由于含氧化合物原料制造烯烃是放热反应，根据本发明，所述反应器可以设置一个或多个激冷介质线，控制反应温度。根据本发明的一种具体实施方式，可以通过设置在所述反应器的中下游(相对于物料流向)的一个或多个激冷介质线向所述反应器中注入激冷介质。在此，所述激冷介质可以为激冷剂或冷却的催化剂，所述激冷剂可以为未预热的所述含氧化合物原料和/或水。
根据本发明，一般地，所述脱水反应的反应温度为200-700℃，优选250-600℃。特别地，为了实现本发明的低碳烯烃增产目的，所述脱水反应的反应压力P一般为0.5-10MPa，优选0.75-3.5MPa，更优选0.8-3MPa，最优选1-2MPa。另外，所述脱水反应的重时空速H一般为7-250h-1，优选8-150h-1，更优选10-100h-1，更优选15-80h-1，最优选15-50h-1。
根据本发明一个特别优选的实施方式，在进行所述脱水反应(换句话说，基于现有的反应器或反应装置进行改造，拟大幅度提高低碳烯烃的产量)时，严格增函数H＝f(P)成立。在此，P(单位是MPa)属于区间[0.55，10.0]，优选属于区间[0.75，3.5]，更优选属于区间[0.8，3.0]，更优选属于区间[1.0，2.0]，并且H(单位是h-1)属于区间[7，250]，优选属于区间[8，150]，更优选属于区间[10，100]，更优选属于区间[15，80]，最优选属于区间[15，50]。根据该严格增函数，在本发 明所规定的特定数值区间内增加所述脱水反应的反应压力P的同时，必须在本发明所规定的相应特定数值区间内增加所述脱水反应的重时空速H。在此，本发明对所述反应压力P和所述重时空速H的增加方式或幅度等没有特别的限定，只要基于本领域技术人员的常规判断，各自的数值的确已经增加即可，但不可以维持恒定或降低。根据本发明一个特别的实施方式，优选反应压力P与重时空速H成比例增加或按照不同或相同的幅度增加，有时可以是等比例增加或同步调增加，直至达到预期的低碳烯烃增产幅度。在某些情况下，当反应压力P达到本发明前述规定的某一数值区间的上限(比如3MPa)时，重时空速H一般也优选达到本发明前述规定的某一数值区间的上限(比如50h-1)，但并不限定于此。
根据本发明，需要特别指出的是，当反应压力P和重时空速H不同时处于本发明前述规定的数值范围或数值区间内时，即使在增加反应压力P的同时同样增加重时空速H，也无法获得如本发明这样大幅度的低碳烯烃增产效果(如实施例所示)。这完全出乎本领域技术人员的意料之外。
根据本发明，所述进一步反应的反应条件包括：反应温度200-700℃，优选300-600℃；反应压力0.1-6MPa，优选0.8-2MPa。
根据本发明，所述制造方法还可以包括：控制所述反应器中的反应压力P与所述再生器中的再生压力之比为3-100∶1。更为具体而言，根据本发明，所述脱水反应的反应压力P比所述再生反应的再生压力至少高0.35MPa，优选至少高0.4MPa、至少高0.5MPa、至少高0.6MPa、至少高0.7MPa、至少高0.8MPa、至少高0.9MPa、至少高1.0MPa、至少高1.1MPa、至少高1.2MPa、至少高1.3MPa、至少高1.4MPa、至少高1.5MPa、至少高1.6MPa、至少高1.7MPa、至少高1.8MPa、至少高1.9MPa或者至少高2.0MPa。或者，根据本发明，所述脱水反应的反应压力P比所述再生反应的再生压力一般至多高5MPa，优选至多高4MPa、至多高3.5MPa、至多高3.3MPa、至多高3MPa、至多高2.5MPa、至多高2.3MPa、至多高2MPa、至多高1.5MPa、至多高1.3MPa或者至多高1MPa。
根据本发明，由于含氧化合物原料制造低碳烯烃和待生催化剂的再生是都放热反应，可以在所述反应器、再生器、再生催化剂进料器 或再生催化剂接收器中设置一个或多个内取热器。所述内取热器可以是盘管、弯管等类型，通过内部流动的水、四氯化碳等液体，对反应器进行取热，其它工业上常用的内取热器，本发明也可以应用。
本发明的发明人通过研究表明，含氧化合物原料与有一定积碳的催化剂接触有利于反应的快速进行，这是因为一方面，催化剂中的积碳作为活性中心不断与含氧化合物原料反应，引入烷基基团；另一方面，这些积碳也在不断地进行脱烷基化反应，生成乙烯和丙烯等低碳烯烃，即所谓的“烃池”反应。为此，根据本发明，所述制造方法还可以包括将至少一部分所述待生催化剂和/或至少一部分所述进一步的待生催化剂循环至所述脱水反应或所述反应器。
根据本发明的一种具体实施方式，可以从所述反应器或所述待生催化剂接收器中引出的一部分待生催化剂，将引出的该部分待生催化剂直接或经取热降低温度后返回到所述反应器中，或者输送到所述反应器下部的催化剂混合器中与再生催化剂混合后返回到所述反应器中，用于反应。
根据本发明的另一种具体实施方式，也可以从所述反应器或所述待生催化剂接收器中引出的一部分待生催化剂，将引出的该部分待生催化剂直接或经取热降低温度后，输送至所述再生催化剂进料器中与再生催化剂混合后返回到所述反应器中；所述引出的该部分待生催化剂的量与通过催化剂料斗输送至再生催化剂进料器中的再生催化剂一起足以维持所述反应器中催化剂的连续运转。
根据本发明，从所述反应器或待生催化剂接收器引出的一部分待生催化剂可以经过外取热器进行取热降低温度，所述外取热器是本领域技术人员所熟知的，内部可以设置盘管、弯管等取热装置，用于降低流经其中的待生催化剂的温度。
根据本发明，所述催化剂混合器可以与所述反应器相连，优选垂直相连，用于输入所述反应器内的混合热再生催化剂、取热后再生催化剂、待生催化剂中的一种或几种。此时，催化剂混合区温度可以为200-600℃，优选300-500℃，压力为0.5-10MPa。
根据本发明，进入所述反应器(进料区)和/或所述进一步的反应器(进料区)的催化剂的总含碳量可以为3-25wt％，优选6-15wt％。此时，所述进入所述反应器或所述进一步的反应器的催化剂可以来自再 生催化剂进料器，也可以来自待生催化剂接收器和/或所述反应器，其中，来自所述再生催化剂进料器中的催化剂可以是再生催化剂，也可以是再生催化剂和待生催化剂的混合催化剂。
根据本发明，本领域技术人员可以理解的是，所述富含低碳烯烃的油气经产品分离回收系统分离后可以得到一部分C4 +烃类，为了增加低碳烯烃的产率，可以将所述C4 +烃类送入所述进一步的反应器进行所述进一步反应，将C4 +烃类裂化为低碳烯烃。
根据本发明的一种具体实施方式，可以将所述再生催化剂进料器中的再生催化剂送入所述进一步的反应器内与所述C4 +烃类并进行所述进一步反应，得到的进一步的富含低碳烯烃的油气和进一步的待生催化剂一起送入所述反应器的所述油剂分离区中；其中，送入所述反应器所述油剂分离区的进一步的富含低碳烯烃的油气和进一步的待生催化剂可以与所述反应器中产生的富含低碳烯烃的油气和待生催化剂混合后一起分离。
根据本发明的进一步的具体实施方式，可以将所述反应器内的待生催化剂送入所述进一步的反应器与所述C4 +烃类接触并进行所述进一步反应，得到的进一步的富含低碳烯烃的油气和进一步的待生催化剂可以在所述进一步的反应器中分离，将分离后的进一步的富含低碳烯烃的油气送入所述产品分离回收系统，将分离后的进一步的待生催化剂送入待生催化剂接收器内。
根据本发明的进一步的一种具体实施方式，可以将经所述反应器汽提段汽提后的待生催化剂送入所述进一步的反应器与所述C4 +烃类并进行所述进一步反应，得到的进一步的富含低碳烯烃的油气和进一步的待生催化剂一起送入所述反应器的所述油剂分离区中；其中，返回所述反应器的进一步的富含低碳烯烃的油气和进一步的待生催化剂可以与所述反应器中产生的富含低碳烯烃的油气和待生催化剂混合后一起分离。
根据本发明更进一步的一种具体实施方式，可以将再生器内的再生催化剂直接送入所述进一步的反应器内与所述C4 +烃类接触并进行所述进一步反应，得到进一步的富含低碳烯烃的油气和进一步的待生催化剂；使所述进一步的富含低碳烯烃的油气和所述进一步的待生催化剂分离在所述进一步的反应器中，可以将分离后的进一步的富含低 碳烯烃的油气送入所述产品分离回收系统，可以将所述进一步的待生催化剂直接送入所述再生器内进行再生。
根据本发明，所述反应器中的脱水制烯烃的反应和所述进一步的反应器中的所述进一步反应可以采用本领域技术人员所熟知的能够生成低碳烯烃的反应条件，并且该二个反应可以采用基本相同的反应条件，也可以采用不同的反应条件。由于所述进一步的反应器中的反应原料与所述反应器中的反应原料不完全相同，因此优选的是根据所述进一步的反应器中的原料情况采用不同于所述反应器的所述进一步反应的反应条件，这是本领域技术人员可以理解的，其中，所述进一步反应主要可以是C4 +烃类的裂化反应。优选的是，该二个反应器中的反应条件可以在，例如，下述范围内进行选择：反应温度可以为200-700℃，优选为250-600℃；反应压力可以为0.5-10MPa，优选为1-3.5MPa。
为了使所述反应器或所述进一步的反应器中反应后产生的(进一步的)富含低碳烯烃的油气和(进一步的)待生催化剂进行分离，或者使所述再生器中再生后产生的再生催化剂与烟气等进行分离，可以使用常规的旋风分离器，这是本领域技术人员所熟知的，本发明对此不进行详细描述。
根据本发明的一种优选的具体实施方式，可以通过过滤器使富含低碳烯烃的油气和待生催化剂分离。另外，也可以通过过滤器使进一步的富含低碳烯烃的油气和进一步的待生催化剂分离。再者，也可以通过过滤器使(进一步的)再生催化剂和烟气等进行分离。通过使用过滤器来分离催化剂，可以有效地除去油气或烟气等中携带的催化剂粉尘，相比于现有技术常规使用的旋风分离器而言，能够最大限度地降低生产过程中催化剂的自然跑损。这是本发明的一大优势所在。
根据本发明，所述过滤器可以采用多孔材料制备，例如，可以选自金属烧结多孔材料和/或陶瓷多孔材料；所述过滤器的2μm颗粒过滤精度可达到99.9％，优选地，所述过滤器的1.2μm颗粒过滤精度可达到99.9％。另外，可以使用反吹气对所述过滤器进行反吹以清理滤饼。在此，反吹气可以为选自含烃气体、干气、氮气和水蒸气中的一种或几种。
根据本发明，所述制造方法还可以包括将未反应完全的含氧化合物原料(包括在所述脱水反应中新生成的各种含氧化合物，特别是二 甲醚)循环至所述脱水反应的步骤，由此实现反应原料的充分利用。
根据本发明，可以方便地通过一个或多个(优选一个或两个)催化剂料斗，将至少一部分所述待生催化剂和/或至少一部分所述进一步的待生催化剂输送至再生反应，和/或，将至少一部分所述再生催化剂和/或至少一部分所述进一步的再生催化剂循环至所述脱水反应和/或所述进一步反应。在此，所述催化剂料斗有时也称为闭锁料斗。在本说明书的下文中，尤其是附图和实施例，均以该催化剂料斗为例对本发明的技术思想和各种具体实施方式进行说明，但本发明显然并不限于此。
根据本发明，所述催化剂料斗可使催化剂从所述反应器的高压烃环境向再生器的低压氧环境，以及从再生器的低压氧环境向所述反应器的高压烃环境安全和有效地转移。也就是说，通过使用催化剂料斗，一方面可以使所述反应器烃类气氛与再生器的烧焦再生的含氧气氛很好地隔离，确保本发明工艺方法的安全性，另一方面可以灵活地调控所述反应器和再生器的操作压力，尤其是在不提高再生器操作压力的情况下能够提高所述反应器的操作压力从而提高装置的处理量。
本发明所述的催化剂料斗是一种可使同一物料流在不同的气氛(例如氧化气氛和烃类气氛)之间和/或不同的压力环境(例如从高压至低压，或者反之)之间进行切换的任何已知装置。比如，通过催化剂料斗实现催化剂从反应器(高压烃环境)向再生器(低压氧环境)输送的步骤可以包括：1、采用热氮气将已排空的催化剂料斗中残存的氧吹扫到再生器中；2、采用干气将氮气从催化剂料斗吹扫出去；3、采用干气对已排空的催化剂料斗加压；4、将来自待生催化剂接收器的待生催化剂填充到已排空的催化剂料斗中；5、通过排出加压催化剂料斗内的干气，对填充的催化剂料斗减压；6、用热氮气将干气从填充的催化剂料斗吹扫出去；7、将待生催化剂从填充的催化剂料斗排放到待生催化剂进料器。比如，通过催化剂料斗实现催化剂从再生器(低压氧环境)向反应器(高压烃环境)循环的步骤可以包括：1、采用热氮气将氧从填充再生催化剂的催化剂料斗吹扫到再生器中；2、采用干气将氮气从催化剂料斗吹扫出去；3、采用干气对填充的催化剂料斗加压；4、将再生催化剂从填充的催化剂料斗排放到再生催化剂进料器；5、通过排出加压催化剂料斗内的干气，对已排空的催化剂料斗减压；6、 用热氮气将干气从已排空的催化剂料斗吹扫出去；7、将再生催化剂从再生催化剂接收器填充到已排空的催化剂料斗。
由于催化剂料斗是批次输送催化剂，根据本发明，所述再生催化剂进料器和待生催化剂循环线的作用是使催化剂向反应器地输送更加连续。然而本发明的发明人发现，再生器内的待生催化剂输入和再生催化剂的输出也可以是成批次的，当催化剂料斗向再生器输送待生催化剂或再生器向催化剂料斗输送再生催化剂时，可以在再生器和催化剂料斗之间依靠重力，或通过提升线形成压力差进行输送，而无需设置待生催化剂进料器或再生催化剂接收器。
根据本发明，若所述进一步的反应器与所述反应器相连通，本领域技术人员可以理解的是，两个反应器的压力可以是相同的，也就是说，若有需要，催化剂料斗也同样可以完成对所述进一步的反应器中催化剂的输送和循环。
根据本发明，所述再生的反应条件是本领域所属技术人员所熟知的，例如，所述再生反应的反应条件包括：反应温度450-850℃，优选550-700℃；反应压力0.1-0.5MPa，优选0.15-0.3MPa，比如常压；含氧气氛。其中，所述的含氧气氛可以为以空气、氮气稀释的空气或者富氧气体作为流化介质。
根据本发明，在维持用于进行所述脱水反应的反应器的尺寸和数量不变的情况下；换句话说，基于现有的反应器或反应装置规模进行产能升级时，通过按照本发明的规定在特定的范围内增加反应器的反应压力和重时空速，就可以大幅度地增加该反应器的含氧化合物原料处理量而相应增加低碳烯烃的产量。此时，低碳烯烃的增产幅度可以达到50％，优选达到100％，更优选达到150％、200％、500％或790％，在本发明最优选的情况下甚至可以达到1000％或更高。
需要强调的是，根据本发明，在维持低碳烯烃的收率与现有技术相比基本上不变或略有提高的基础上，通过提高反应器或反应装置的含氧化合物原料处理量或通过量来实现低碳烯烃的增产。因此，与以牺牲低碳烯烃的收率(比如降低幅度超过20％)为代价，通过简单地提高反应器或反应装置的含氧化合物原料处理量或通过量来实现低碳烯烃增产的情况相比，本发明中低碳烯烃的增产幅度要显著更高。在此，根据本发明，所述低碳烯烃的收率可以维持在与现有技术相当的 水平甚至更高，比如一般为60％-95％或者78％-95％。
从另一个角度来看，通过按照本发明的前述规定来制造低碳烯烃，在确保达到预定的低碳烯烃产量的同时，与现有技术相比，可以显著降低反应器或反应装置的尺寸和数量，由此降低整个低碳烯烃生产装置的规模和投资成本。
下面结合附图进一步说明本发明所具体实施方式，但本发明并不因此而受到任何限制。为了方便描述起见，以下以提升管型反应器作为反应器的例子，但本发明并不限于此。
第一种具体实施方式
图1是按本发明的第一种具体实施方式的含氧化合物原料制造低碳烯烃方法的流程示意图。
如图1所示，含氧化合物原料自进料线24进入提升管型反应器的提升管反应器1，与经预提升线28提升的来自管线23的催化剂接触进行脱水制烯烃的反应，反应后油气进入密相床反应器3，内取热器2取出密相床反应器3的多余热量，反应油气在密相床反应器3继续反应，产生的富含低碳烯烃的油气和待生催化剂进入沉降区5，沉降后的待生催化剂返回密相床反应器3，富含低碳烯烃的油气及携带的待生催化剂细粉经过滤器6过滤后，富含低碳烯烃的油气经管线25送入产品分离回收系统(未图示)，过滤后的待生催化剂细粉沉降返回密相床反应器经汽提段4汽提，汽提后的部分待生催化剂经管线16送入待生催化剂接收器8，另一部分待生催化剂经外取热器13取热后送入再生催化剂进料器12。
来自待生催化剂接收器8的待生催化剂经管线17进入催化剂料斗9，降压后经管线18送入待生催化剂进料器10，经管线19送入再生器7与来自管线27的主风逆流接触烧焦再生，多余热量经内取热器15取出(取热量可以通过取热物流量以及内取热器15埋入密相床层高度控制)，烟气经管线26送入后续能量回收、净化系统(未图示)，再生催化剂经管线20送入再生催化剂接收器11，再生催化剂多余热量经内取热器14取出，取热后的再生催化剂经管线21送入催化剂料斗9，升压后经管线22送入再生催化剂进料器12与来自取热器13的待生催化剂混合后，经管线23送入提升管反应器1预提升段。
第二种具体实施方式
图2是按本发明的第二种具体实施方式的含氧化合物原料制造低碳烯烃方法的流程示意图。
如图2所示，来自管线223的催化剂与来自管线213的待生催化剂在催化剂混合器211中混合，经预提升线219预提升气提升后送入提升管型反应器的提升管反应器201，含氧化合物原料经进料线224进入提升管反应器201，与来混合器211的催化剂接触并发生脱水制烯烃反应，反应后油气在内提升管及分布板202继续反应，然后进入密相床反应器203，来自激冷介质线220的激冷介质进入提升管反应器201控制反应温度，未转化的原料在密相床反应器203与催化剂继续接触反应，多余反应热由内取热器215取出，得到富含低碳烯烃的油气和待生催化剂进入沉降区205沉降后，待生催化剂进入密相床反应器203，富含低碳烯烃的油气及携带的待生催化剂细粉经过滤器206过滤后，富含低碳烯烃的油气经管线225送入产品分离回收系统(未图示)，过滤后的待生催化剂细粉沉降返回密相床反应器203。待生催化剂在汽提段204汽提后，部分汽提后的待生催化剂经管线213返回催化剂混合器211，另一部分汽提后的待生催化剂经管线216送入待生催化剂接收器208。
来自待生催化剂接收器208的待生催化剂经管线217进入催化剂料斗209，降压后经管线221送入再生器207与来自管线227的主风逆流接触烧焦再生，烟气经管线226送入后续能量回收、净化系统(未图示)，再生催化剂经管线215送入再生催化剂接收器210，再生催化剂多余热量经内取热器214取出，取热后再生催化剂经管线218送入催化剂料斗209，升压后经管线222送入再生催化剂进料器212，经管线223送入催化剂混合器211。
第三种具体实施方式
图3是按本发明的第三种具体实施方式的含氧化合物原料制造低碳烯烃方法的流程示意图。
如图3所示，含氧化合物原料自进料线324进入提升管型反应器的提升管反应器301，与来自管线323的催化剂接触反应并发生脱水制烯烃反应，反应后的油气进入扩径提升管302进行继续反应，继续反应后的油气进入密相床反应器303进行反应，得到的富含低碳烯烃的油气和待生催化剂进入沉降区305沉降，沉降后的待生催化剂返回密 相床反应器，富含低碳烯烃的油气及携带的待生催化剂细粉经过滤器306过滤后，富含低碳烯烃的油气经管线325送入产品分离回收系统(未图示)，过滤后的待生催化剂细粉沉降返回密相床反应器的汽提段304气体，汽提后的部分待生催化剂经管线316送入后待生催化剂接收器308，另一部分待生催化剂送入外取热器313取热后送入再生催化剂进料器312。
来自待生催化剂接收器308的待生催化剂经管线317进入催化剂料斗309，降压后经管线318送入待生催化剂进料器310，经管线319送入再生器307与来自管线327的主风逆流接触烧焦再生，多余热量经内取热器315取出(取热量可以通过取热物流量以及内取热器315埋入密相床层高度控制)，烟气经管线326送入后续能量回收、净化系统(未图示)，再生催化剂经管线320送入再生催化剂接收器311，再生催化剂多余热量经内取热器314取出，取热后再生催化剂经管线321送入催化剂料斗309，升压后经管线322送入再生催化剂进料器312与来自外取热器313的待生催化剂混合后，分别经管线323、332送入提升管反应器301和进一步的提升管型反应器330。
来自管线332的催化剂进入进一步的提升管型反应器330预提升段，经来自预提升线328的预提升介质提升后进入进一步的提升管型反应器330，经产品分离回收系统分离的C4 +烃类经原料进料线329进入进一步的提升管型反应器330与催化剂接触进行进一步反应，生成的富含低碳烯烃的油气经管线331进入密相床流化床303。
第四种具体实施方式
图4是按本发明的第四种具体实施方式的含氧化合物原料制造低碳烯烃方法的流程示意图。
如图4所示，来自管线423的催化剂与来自外取热器413的待生催化剂在催化剂混合器411中混合，经预提升线419预提升气提升后送入提升管型反应器的提升管反应器401，含氧化合物原料经进料线424进入提升管反应器401，与来催化剂混合器411的催化剂接触反应发生脱水制烯烃反应，反应后的产物及催化剂经内提升管及分布板402进入密相床反应器403，来自激冷介质线420的激冷介质进入提升管反应器401控制反应温度，未转化的原料在密相床反应器403与催化剂继续接触反应，多余反应热由内取热器415取出，生成的富含低碳烯 烃的油气和待生催化剂进入沉降区405沉降，待生催化剂进入汽提段404，含低碳烯烃的油气及携带的待生催化剂细粉经过滤器406过滤后，含低碳烯烃的油气经管线425送入产品分离回收系统(未图示)，过滤后的催化剂细粉沉降返回密相床反应器403。待生催化剂在汽提段404汽提，部分汽提后的待生催化剂经内取热器413返回催化剂混合器411，另一部分经管线430送入进一步的提升管型反应器的第一反应区431。经产品分离回收系统分离的C4 +烃类经管线429送入进一步的提升管型反应器的第一反应区431，与来自管线430的催化剂接触进行进一步反应，反应生成的富含低碳烯烃的油气和催化剂进入第二反应区432继续反应，经缩径快分433送入沉降区435，富含低碳烯烃的油气及携带的催化剂细粉经过滤器436过滤后，油气经管线437送入分离回收系统(未图示)，催化剂经汽提段434汽提后经管线438、416送入催化剂循环系统。
来自待生催化剂接收器408的待生催化剂经管线417进入催化剂料斗409，降压后经管线421送入再生器407与来自管线427的主风逆流接触烧焦再生，烟气经管线426送入后续能量回收、净化系统(未图示)，再生催化剂经管线415送入再生催化剂接收器410，再生催化剂多余热量经内取热器414取出，取热后再生催化剂经管线418送入催化剂料斗409，升压后经管线422送入再生催化剂进料器412，然后经管线423送入催化剂混合器411。
第五种具体实施方式
图5是按本发明的第五种具体实施方式的含氧化合物原料制造低碳烯烃方法的流程示意图。
如图5所示，含氧化合物原料经进料线524进入提升管型反应器的提升管反应器501，与来自管线523的催化剂接触反应进行脱水制烯烃的反应，反应后的产物及催化剂经内提升管及快分502进入密相床反应器503，来自激冷介质线528的激冷介质进入提升管反应器501控制反应温度，未转化的原料在密相床反应器503与催化剂继续接触反应，得到的富含低碳烯烃的油气和待生催化剂进入沉降区505，富含低碳烯烃的油气及携带的待生催化剂细粉经过滤器506过滤后，富含低碳烯烃的油气经管线525送入产品分离回收系统(未图示)，过滤后的待生催化剂细粉沉降进入密相床反应器503。待生催化剂经汽提段 504汽提后，部分汽提后的待生催化剂经管线516送入待生催化剂接收器508，部分汽提后的待生催化剂经管线535送入进一步的提升管型反应器的第一反应区530，剩下的待生催化剂送入外取热器513取热后送入内提升管及快分502。经产品分离回收系统分离的C4 +烃类经管线533送入进一步的提升管型反应器的第一反应区530和第二反应区及缩径531，与来自管线535的经预提升线532的预提升气与提升的催化剂接触发生进一步反应，反应生成的富含低碳烯烃的油气和催化剂经管线534送入密相床503。
来自待生催化剂接收器508的部分待生催化剂经外取热器529取热后送入再生催化剂进料器512，另一部分经管线517进入催化剂料斗509，降压后经管线518送入待生催化剂进料器510，经管线519送入再生器507与来自管线527的主风逆流接触烧焦再生，多余热量经内取热器515取出，烟气经管线526送入后续能量回收、净化系统(未图示)，再生催化剂经管线520送入再生催化剂接收器511，再生催化剂多余热量经内取热器514取出，取热后再生催化剂经管线521送入催化剂料斗509，升压后经管线522送入再生催化剂进料器催化剂进料器512，来自待生催化剂508的待生催化剂经取热器529取热后送入再生催化剂进料器512，待生催化剂和再生催化剂混合后经管线523送入提升管反应器501。
第六种具体实施方式
图6是按本发明的第六种具体实施方式的含氧化合物原料制造低碳烯烃方法的流程示意图。
如图6所示，来自管线637的催化剂与来自外取热器613的待生催化剂在催化剂混合器611中混合，经预提升线619预提升气提升后送入提升管型反应器的第一反应区601，原料经进料线624进入第一反应区601，与来催化剂混合器611的催化剂接触发生脱水制烯烃的反应，反应后的产物及催化剂进入第二反应区602，未转化的原料在第二反应区602与催化剂继续接触反应，得到富含低碳烯烃的油气和待生催化剂经缩径及快分603进入沉降区605，富含低碳烯烃的油气及携带的待生催化剂细粉经过滤器606过滤后，富含低碳烯烃的油气经管线625送入产品分离回收系统(未图示)，过滤后的待生催化剂细粉沉降进入汽提段604。经汽提段汽提后的部分待生催化剂经管线616送入待生 催化剂接收器608，部分送入外取热器613取热后经管线624送入催化剂混合器611，剩下部分经外取热器612取热后送入再生催化剂进料器636。
来自待生催化剂接收器608的待生催化剂经管线617进入催化剂料斗609，降压后经管线621送入再生器607与来自管线622的主风逆流接触烧焦再生，烟气经管线620送入后续能量回收、净化系统(未图示)，多余热量由内取热器615取出，再生催化剂送入再生催化剂接收器610，再生催化剂多余热量经内取热器614取出，取热后再生催化剂经管线618送入催化剂料斗609，升压后经管线623送入再生催化剂进料器636与来自外取热器612的待生催化剂混合后经管线637送入催化剂混合器611。
来自管线633的再生催化剂经预提升气626预提升后，与来自进料线635的C4 +烯烃在进一步的提升管型反应器627中进行进一步反应，反应油气及催化剂经快分628分离出富含低碳烯烃的油气和待生催化剂，分离后的油气经沉降区629沉降，再经过滤器631过滤，过滤后的油气经管线632送入后续分离系统(未图示)，待生催化剂经汽提段630汽提，汽提后待生催化剂经管线634送入再生器607再生。
第七种具体实施方式
图7是按本发明的第七种具体实施方式的含氧化合物原料制造低碳烯烃方法的流程示意图。
如图7所示，含氧化合物原料和稀释剂自进料线702进入流化床反应器701，与来自管线723的催化剂接触进行脱水制烯烃的反应，内取热器713取出流化床反应器701的多余热量，产生的富含低碳烯烃的油气和部分待生催化剂进入沉降区5，沉降后的待生催化剂返回流化床反应器701，富含低碳烯烃的油气及携带的待生催化剂细粉经过滤器706过滤后，富含低碳烯烃的油气经管线703送入产品分离回收系统(未图示)，过滤后的待生催化剂细粉沉降返回流化床反应器，部分待生催化剂经管线16送入待生催化剂接收器8并进行汽提。
来自待生催化剂接收器708的待生催化剂经管线717进入催化剂料斗709，降压后经管线718送入待生催化剂进料器710，经管线719送入再生器707与来自管线724的主风逆流接触烧焦再生，多余热量经内取热器715取出，烟气经管线704送入后续能量回收、净化系统 (未图示)，再生催化剂经管线720送入再生催化剂接收器711，再生催化剂多余热量经内取热器714取出，取热后的再生催化剂经管线721送入催化剂料斗709，升压后经管线722送入再生催化剂进料器712并进行汽提，经管线723送入流化床反应器701。
实施例
以下实施例用于举例说明本发明，而不是用于限定本发明。
实施例1-6
实施例1-6按图1所示工艺进行(本发明实施例和对比例按图1所示工艺进行试验时，采用相同反应器)，反应条件基本一致，仅改变反应压力和重时空速。反应原料、催化剂、反应条件以及产品收率列于表1。
从实施例1-6可以看出，采用本发明实施方式，在其它反应条件大致相当时，提高反应压力的同时，相应提高重时空速，低碳烯烃收率最高可以达到84.9％。
实施例7-8
实施例7-8按图1所示工艺进行，与实施例4相比，实施例7-8提高反应压力时，没有相应提高重时空速，其它操作条件大致相当。反应原料、催化剂、反应条件以及产品收率列于表2。
从实施例4、实施例7和实施例8中可以看出，与实施例4相比，在其它反应条件大致相当时，如果只单纯提高反应压力，而不相应提高重时空速，低碳烯烃收率和收率均会下降，低碳烯烃收率从实施例4的84.3％最低降低到实施例8的82.9％；低碳烯烃产量从实施例4的3.73kg/h降低到实施例8的2.64kg/h。
实施例9
实施例9按图1所示工艺进行，与实施例3相比，操作条件大致相当，仅改变反应器入口催化剂含碳量(反应器入口催化剂含碳量是指反应器入口未与原料接触反应前催化剂含碳量)。反应原料、催化剂、反应条件以及产品收率列于表2。
从实施例3和实施例9中可以看出，与实施例3相比，在反应压力和重时空速大致相当时，反应器入口催化含碳量从7.3％降低到4.5％时，低碳烯烃产量从3.84kg/h降低到3.58kg/h；低碳烯烃收率从84.9％ 降低到84.3％；乙烯与丙烯的质量比变大。
实施例10
实施例10按图3所示工艺进行，反应原料、催化剂、反应条件以及产品收率列于表2。
从实施例10中可以看出，低碳烯烃收率为93.3％，低碳烯烃产量为4.03kg/h。
实施例11
实施例11按图7所示工艺进行，反应原料原料为乙醇，催化剂、反应条件以及产品收率列于表2。
从实施例11中可以看出，低碳烯烃收率为82.8％，低碳烯烃产量为3.26kg/h.
实施例12
实施例12按图2所示工艺进行，反应原料、催化剂、反应条件以及产品收率列于表4。
实施例12为增产低碳烯烃的同时生产汽油产品生产方案。从实施例12中可以看出，丙烯收率为65.9％，汽油收率为25.3％，丙烯产量为1.98kg/h，汽油产量为0.76kg/h。
对比例1-6
对比例1-6所采用的反应器、原料、催化剂与实施例1-10一致，按图1所示工艺进行。与实施例1-10相比，对比例1为常规甲醇制低碳烯烃条件，反应器入口催化剂含碳量明显相对较低，反应压力和重时空速明显低于本发明；对比例2-6仅改变反应压力和重时空速，其它操作条件与实施例1-10大致相当。对比例1-6反应原料、催化剂、反应条件以及产品收率列于表3。
对比例1为常规甲醇制低碳烯烃操作条件，对比例2仅与对比例1反应器入口催化剂含碳量不同。与对比例1相比，反应器入口催化剂含碳量从对比例1的1.5％提高到对比例2的7.2％后，低碳烯烃产量从对比例1的0.47kg/h降低到对比例2的0.43kg/h；转化率从对比例1的100％降低到对比例2的80.2％。
对比例2与实施例1-8相比，除反应压力和重时空速不同外，反应器入口催化剂含碳量等均大致相当。与对比例2相比，实施例1-8低碳烯烃收率大致相当或有所增加，低碳烯烃产量大幅度提高，例如，低 碳烯烃收率由对比例2的79.2％提高到实施例3的84.9％；低碳烯烃产量由对比例2的0.43kg/h提高到实施例3的3.84kg/h，提高幅度高达793.02％。
对比例3和实施例3除重时空速外，其余操作条件大致相当。对比例3重时空速明显高于实施例3。与对比例3相比，采用本发明后，低碳烯烃产量从对比例3的0.46kg/h提高到实施例3的3.84kg/h，提高幅度高达734.78％；低碳烯烃收率从对比例3的82.9％提高到实施例3的84.9％。
对比例4和实施例2除反应压力外，其余操作条件大致相当。对比例4反应压力明显低于实施例2。与对比例4相比，采用本发明后，低碳烯烃产量从对比例4的0.44kg/h提高到实施例2的2.54kg/h，提高幅度高达477.27％；低碳烯烃收率从比例4的79.4％提高到实施例2的84.4％。
对比例5和实施例2除重时空速外，其余操作条件大致相当。对比例5重时空速明显高低于实施例2。与对比例5相比，采用本发明后，低碳烯烃产量从对比例5的0.43kg/h提高到实施例2的2.54kg/h，提高幅度高达490.70％；低碳烯烃收率从对比例5的75.3％提高到实施例2的84.4％。
对比例6和实施例1除反应压力外，其余操作条件大致相当。对比例6反应压力明显高于实施例1。与对比例6相比，采用本发明后，低碳烯烃产量从对比例6的0.33kg/h提高到实施例1的1.11kg/h，提高幅度高达236.36％；低碳烯烃收率从对比例6的33.4％提高到实施例1的83.7％。
表1
Figure PCTCN2015000705-appb-000001
*收率＝产品产量/除含氧化合物外烃类产品产量之和×100
表2
Figure PCTCN2015000705-appb-000002
*收率＝产品产量/除含氧化合物外烃类产品产量之和×100
表3
Figure PCTCN2015000705-appb-000003
*收率＝产品产量/除含氧化合物外烃类产品产量之和×100
表4
Figure PCTCN2015000705-appb-000004
*收率＝产品产量/除含氧化合物外烃类产品产量之和×100
实施例I
实施例I按图1所示工艺进行，反应原料、催化剂、反应条件以及产品收率列于表I。
实施例II
实施例II按图3所示工艺进行，反应原料、催化剂、反应条件以及产品收率列于表I。
实施例III
实施例III按图2所示工艺进行，反应原料、催化剂、反应条件以及产品收率列于表II。
从表I可以看出，采用本发明的方法，乙烯和丙烯的收率能够高于现有工业工艺的水平；从表II可以看出，采用本发明方法，丙烯和汽油收率分别为65.9％和25.3％，高于现有工业工艺水平，并且由于本发明的反应系统压力高于现有工业装置，故在其它操作条件相同情况下，本发明反应系统的原料处理量高于现有工业装置。
表I
Figure PCTCN2015000705-appb-000005
表II
  实施例III
反应原料(含氧化合物原料质量分数＞98wt％) 工业甲醚
催化剂(齐鲁催化剂公司) ZSM-5分子筛催化剂
提升管型反应器的反应条件  
温度，℃ 500
压力，MPa 1.5
再生器的烧焦再生条件  
再生压力，MPa 0.2
再生温度，℃ 650
再生催化剂定量，重％ 0.4
产品收率  
乙烯，％ 4.1
丙烯，％ 65.9
汽油，％ 25.3

Claims (15)

  1. 一种低碳烯烃的制造方法，其特征在于，在通过连续地使含氧化合物原料与催化剂接触而发生脱水反应以制造低碳烯烃的方法中，使所述脱水反应的反应压力P为0.5-10MPa，优选0.75-3.5MPa，更优选0.8-3MPa，最优选1-2MPa，所述脱水反应的重时空速H为7-250h-1，优选8-150h-1，更优选10-100h-1，更优选15-80h-1，最优选15-50h-1。
  2. 根据权利要求1所述的制造方法，其中在进行所述脱水反应时，严格增函数H＝f(P)成立，其中P(单位是MPa)属于区间[0.55，10.0]，优选属于区间[0.75，3.5]，更优选属于区间[0.8，3.0]，最优选属于区间[1.0，2.0]，H(单位是h-1)属于区间[7，250]，优选属于区间[8，150]，更优选属于区间[10，100]，更优选属于区间[15，80]，最优选属于区间[15，50]。
  3. 根据权利要求1所述的制造方法，包括以下步骤：
    连续地使所述含氧化合物原料与所述催化剂接触而发生所述脱水反应，获得富含低碳烯烃的油气和待生催化剂，
    将至少一部分所述待生催化剂输送至再生反应，获得再生催化剂，和
    将至少一部分所述再生催化剂循环至所述脱水反应，
    其中所述脱水反应的反应压力P比所述再生反应的再生压力至少高0.35MPa，优选至少高0.4MPa、至少高0.5MPa、至少高0.6MPa、至少高0.7MPa、至少高0.8MPa、至少高0.9MPa、至少高1.0MPa、至少高1.1MPa、至少高1.2MPa、至少高1.3MPa、至少高1.4MPa、至少高1.5MPa、至少高1.6MPa、至少高1.7MPa、至少高1.8MPa、至少高1.9MPa或者至少高2.0MPa。
  4. 根据权利要求3所述的制造方法，还包括分离所述富含低碳烯烃的油气而获得C4 +烃类的步骤，并且任选还包括以下步骤：
    连续地使所述C4 +烃类与进一步的催化剂接触而发生进一步反应，获得进一步的富含低碳烯烃的油气和进一步的待生催化剂，
    将至少一部分所述进一步的待生催化剂输送至所述再生反应，获得进一步的再生催化剂，和
    将至少一部分所述再生催化剂和/或至少一部分所述进一步的再生 催化剂循环至所述脱水反应和/或所述进一步反应。
  5. 根据权利要求1或4所述的制造方法，其中用于进行所述脱水反应的反应器和/或用于进行所述进一步反应的反应器的数目是一个或多个，并且各自独立地选自流化床反应器、密相床反应器、提升管反应器、沸腾床反应器、浆态床反应器、以及这些反应器中两种或更多种的复合形式，优选选自提升管反应器，更优选各自独立地选自等直径提升管反应器、等线速提升管反应器、变径提升管反应器以及提升管复合密相床反应器。
  6. 根据权利要求1所述的制造方法，其中所述含氧化合物原料选自醇、醚和酯中的至少一种，优选选自R1-O-R2、R1-OC(＝O)O-R2、R1-C(＝O)O-R2和R1-C(＝O)-R2(其中，R1和R2彼此相同或不同，各自独立地选自氢和C1-6支链或直链烷基，优选各自独立地选自氢和C1-4支链或直链烷基，前提是R1和R2中的至多一个是氢)中的至少一种，更优选选自甲醇、乙醇、二甲醚、二乙醚、甲乙醚、碳酸二甲酯和甲酸甲酯中的至少一种。
  7. 根据权利要求1或4所述的制造方法，其中所述催化剂和所述进一步的催化剂彼此相同或不同，各自独立地选自分子筛催化剂中的至少一种，优选各自独立地选自硅铝磷酸盐分子筛催化剂和硅铝酸盐分子筛催化剂中的至少一种。
  8. 根据权利要求3所述的制造方法，其中所述再生反应的反应条件包括：反应温度450-850℃，优选550-700℃；反应压力0.1-0.5MPa，优选0.15-0.3MPa；含氧气氛，优选空气气氛或者氧气气氛。
  9. 根据权利要求3或4所述的制造方法，其中通过过滤器分离出所述待生催化剂和/或所述进一步的待生催化剂和/或所述再生催化剂和/或所述进一步的再生催化剂。
  10. 根据权利要求3或4所述的制造方法，其中通过一个或多个(优选一个或两个)催化剂料斗(9)实现所述输送和所述循环。
  11. 根据权利要求3或4所述的制造方法，还包括将至少一部分所述待生催化剂和/或至少一部分所述进一步的待生催化剂循环至所述脱水反应和/或所述进一步反应的步骤。
  12. 根据权利要求1或4所述的制造方法，其中所述催化剂和/或所述进一步的催化剂的总含碳量为3-25wt％，最优选6-15wt％。
  13. 根据权利要求1所述的制造方法，其中在维持用于进行所述脱水反应的反应器的尺寸和数量不变的情况下，该制造方法能够使低碳烯烃的产量提高50％，优选提高100％，更优选提高150％、200％、500％或790％，最优选提高1000％或更高。
  14. 根据权利要求1所述的制造方法，还包括将未反应完全的含氧化合物原料循环至所述脱水反应的步骤。
  15. 根据权利要求1所述的制造方法，包括以下步骤：
    连续地将所述含氧化合物原料在提升管型反应器中与所述催化剂接触进行所述脱水反应，产生富含低碳烯烃的油气和待生催化剂；
    使富含烯烃的油气和待生催化剂在油剂分离区进行分离，将分离后的富含烯烃的油气送入产品分离回收系统，将待生催化剂经提升管型反应器中的汽提段汽提后从所述提升管型反应器引出并输送至待生催化剂接收器；
    将待生催化剂接收器中的待生催化剂通过催化剂料斗直接输送至再生器，或先通过催化剂料斗输送至待生催化剂进料器后再输送至再生器，并在再生器中在含氧气氛下进行烧焦再生，得到再生催化剂；
    将再生催化剂直接输送到催化剂料斗，或先从再生器引出并输送至再生催化剂接收器，然后再输送至催化剂料斗；
    将催化剂料斗内的再生催化剂输送至再生催化剂进料器后返回到所述提升管型反应器中。
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Cited By (1)

* 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

Families Citing this family (13)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN107266279B (zh) * 2016-04-06 2020-05-19 中国石油化工股份有限公司 一种由含氧化合物制取低碳烯烃的方法
CN107469798B (zh) * 2016-06-07 2020-12-04 中国石油化工股份有限公司 一种硫吸附剂的再生方法
CN107400532A (zh) * 2017-06-13 2017-11-28 神华集团有限责任公司 甲醇制低碳烯烃装置
CN107982967A (zh) * 2017-12-29 2018-05-04 逸盛大化石化有限公司 一种降低产品pt酸的压力过滤机反吹气系统
CN110951501B (zh) * 2018-09-27 2022-04-12 中国石油化工股份有限公司 一种低生焦原料的催化转化方法
CN111659322A (zh) * 2019-03-06 2020-09-15 浙江佳汇新材料有限公司 一种制备1,1,1,3-四氯丙烷的装置及工艺
US20200354636A1 (en) * 2019-05-08 2020-11-12 Exxonmobil Research And Engineering Company Removal of catalyst fines from fluidized bed effluent in the conversion of oxygenate feedstock
WO2020227849A1 (zh) * 2019-05-10 2020-11-19 中国科学院大连化学物理研究所 甲醇和/或二甲醚制烯烃催化剂的部分再生方法及甲醇和/或二甲醚制烯烃方法
CN110437023A (zh) * 2019-07-12 2019-11-12 中国神华煤制油化工有限公司 制备烯烃的方法和烯烃生产系统
CN112745895B (zh) * 2019-10-31 2022-10-21 中国石油化工股份有限公司 临氢催化裂解fcc循环油连续生产芳烃的方法和装置
CN112774735A (zh) * 2019-11-09 2021-05-11 洛阳维达石化工程有限公司 一种催化剂预烃池化的方法及其设备
CN111330517B (zh) * 2020-03-18 2022-04-29 山东凯泰科技股份有限公司 一种催化剂添加及过滤的自动控制系统
CN113087584A (zh) * 2021-03-15 2021-07-09 中石化洛阳工程有限公司 一种含氧化合物生产低碳烯烃的方法

Citations (2)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
DE3524890A1 (de) * 1984-07-14 1986-01-23 The Agency Of Industrial Science And Technology, Tokio/Tokyo Verfahren zur herstellung niedermolekularer olefine
CN101348404A (zh) * 2007-07-18 2009-01-21 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

Family Cites Families (15)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
FI97595C (fi) * 1994-01-04 1997-01-10 Nokia Telecommunications Oy Matkaviestinjärjestelmä ja matkaviestinjärjestelmän tukiasema
AU2198595A (en) * 1994-04-01 1995-10-23 Minnesota Mining And Manufacturing Company Clothing bearing retroreflective appliques
US6531639B1 (en) 2000-02-18 2003-03-11 Exxonmobil Chemical Patents, Inc. Catalytic production of olefins at high methanol partial pressures
US7038102B2 (en) * 2003-07-30 2006-05-02 Exxonmobil Chemical Patents Inc. Liquid contacting of post-quench effluent vapor streams from oxygenate to olefins conversion to capture catalyst fines
US7138558B2 (en) * 2004-11-01 2006-11-21 Uop Llc Direct return of oxygenate recycle stream in olefin production process
US20050203326A1 (en) * 2004-12-16 2005-09-15 Miller Lawrence W. Process for containment of catalyst particles in a oxygenate-to-olefin process
US7439414B2 (en) * 2005-10-13 2008-10-21 Uop Llc Oxygenate conversion catalyst processing
CN101274872B (zh) * 2007-03-30 2012-06-27 中国石油化工股份有限公司 一种乙醇脱水制乙烯的流化装置和工艺方法
CN101279877B (zh) * 2007-04-04 2011-07-20 中国石油化工股份有限公司 含氧化合物转化过程中提高乙烯、丙烯收率的方法
US9061954B2 (en) 2008-02-07 2015-06-23 Total Research & Technology Feluy Dehydration of alcohols on crystalline silicates
CN101260013B (zh) * 2008-04-24 2011-07-20 中国石油化工股份有限公司 含氧化合物制备低碳烯烃的方法
KR101544257B1 (ko) 2010-03-15 2015-08-12 토탈 리서치 앤드 테크놀로지 펠루이 산 촉매 상에서의 이소부탄올의 동시적인 탈수 및 골격 이성질체화
CN102814151B (zh) 2011-06-08 2014-02-26 富德（北京）能源化工有限公司 由含氧化合物制烯烃的流化床反应器和方法
CN103157502B (zh) 2011-12-19 2015-04-15 中国科学院大连化学物理研究所 甲醇和/或二甲醚制取乙烯丙烯的催化剂、其制法及应用
WO2014207134A1 (en) * 2013-06-27 2014-12-31 Shell Internationale Research Maatschappij B.V. An olefin cracking catalyst

Patent Citations (2)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
DE3524890A1 (de) * 1984-07-14 1986-01-23 The Agency Of Industrial Science And Technology, Tokio/Tokyo Verfahren zur herstellung niedermolekularer olefine
CN101348404A (zh) * 2007-07-18 2009-01-21 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

Cited By (2)

* 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

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