WO2016061905A1 - 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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WO2016061905A1
WO2016061905A1 PCT/CN2015/000704 CN2015000704W WO2016061905A1 WO 2016061905 A1 WO2016061905 A1 WO 2016061905A1 CN 2015000704 W CN2015000704 W CN 2015000704W WO 2016061905 A1 WO2016061905 A1 WO 2016061905A1
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catalyst
reaction
mpa
reactor
dehydrogenation
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PCT/CN2015/000704
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English (en)
French (fr)
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王新
于敬川
李明罡
龚剑洪
宗保宁
许友好
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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 claimed from CN201410557916.7A external-priority patent/CN105585409B/zh
Priority claimed from CN201410557715.7A external-priority patent/CN105585400B/zh
Application filed by 中国石油化工股份有限公司, 中国石油化工股份有限公司石油化工科学研究院 filed Critical 中国石油化工股份有限公司
Priority to JP2017521190A priority Critical patent/JP6693952B2/ja
Priority to SG11201703275TA priority patent/SG11201703275TA/en
Priority to US15/520,721 priority patent/US10144680B2/en
Priority to CN201580011667.4A priority patent/CN106068253B/zh
Publication of WO2016061905A1 publication Critical patent/WO2016061905A1/zh

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    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C5/00—Preparation of hydrocarbons from hydrocarbons containing the same number of carbon atoms
    • C07C5/32—Preparation of hydrocarbons from hydrocarbons containing the same number of carbon atoms by dehydrogenation with formation of free hydrogen
    • C07C5/327—Formation of non-aromatic carbon-to-carbon double bonds only
    • C07C5/333—Catalytic processes
    • C07C5/3335—Catalytic processes with metals
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C5/00—Preparation of hydrocarbons from hydrocarbons containing the same number of carbon atoms
    • C07C5/32—Preparation of hydrocarbons from hydrocarbons containing the same number of carbon atoms by dehydrogenation with formation of free hydrogen
    • C07C5/327—Formation of non-aromatic carbon-to-carbon double bonds only
    • C07C5/333—Catalytic processes
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C5/00—Preparation of hydrocarbons from hydrocarbons containing the same number of carbon atoms
    • C07C5/32—Preparation of hydrocarbons from hydrocarbons containing the same number of carbon atoms by dehydrogenation with formation of free hydrogen
    • C07C5/327—Formation of non-aromatic carbon-to-carbon double bonds only
    • C07C5/333—Catalytic processes
    • C07C5/3335—Catalytic processes with metals
    • C07C5/3337—Catalytic processes with metals of the platinum group
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C11/00—Aliphatic unsaturated hydrocarbons
    • C07C11/02—Alkenes
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C2521/00—Catalysts comprising the elements, oxides or hydroxides of magnesium, boron, aluminium, carbon, silicon, titanium, zirconium or hafnium
    • C07C2521/02—Boron or aluminium; Oxides or hydroxides thereof
    • C07C2521/04—Alumina
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C2523/00—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group C07C2521/00
    • C07C2523/38—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group C07C2521/00 of noble metals
    • C07C2523/54—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group C07C2521/00 of noble metals combined with metals, oxides or hydroxides provided for in groups C07C2523/02 - C07C2523/36
    • C07C2523/56—Platinum group metals
    • C07C2523/62—Platinum group metals with gallium, indium, thallium, germanium, tin or lead
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C2523/00—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group C07C2521/00
    • C07C2523/70—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group C07C2521/00 of the iron group metals or copper
    • C07C2523/76—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group C07C2521/00 of the iron group metals or copper combined with metals, oxides or hydroxides provided for in groups C07C2523/02 - C07C2523/36
    • C07C2523/84—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group C07C2521/00 of the iron group metals or copper combined with metals, oxides or hydroxides provided for in groups C07C2523/02 - C07C2523/36 with arsenic, antimony, bismuth, vanadium, niobium, tantalum, polonium, chromium, molybdenum, tungsten, manganese, technetium or rhenium
    • C07C2523/85—Chromium, molybdenum or tungsten
    • C07C2523/86—Chromium
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C2529/00—Catalysts comprising molecular sieves
    • 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

Definitions

  • This invention relates to a process for the manufacture of lower olefins from alkane 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 alkane feedstock.
  • low-carbon olefins (C 2 -C 4 olefins) play an important role in the modern petroleum and chemical industries.
  • a technique for producing a low-carbon olefin using an alkane raw material is increasingly favored.
  • the reactors employed in the manufacture of lower olefins from alkane feedstocks are primarily fixed bed reactors and fluidized bed reactors.
  • the fixed bed reactor has the disadvantages of poor heat transfer effect of the bed, relatively troublesome replacement and regeneration of the catalyst, and difficulty in continuous reaction, but has the advantage that a large amount of alkane raw material can be processed.
  • fluidized bed reactors can address the aforementioned disadvantages of fixed bed reactors, but their alkane feedstock throughput is much less than that of fixed bed reactors of the same size.
  • 7,235, 706 B2 disclose a process for the production of light olefins by dehydrogenation of an alkane feedstock using a fluidized bed reactor and a regenerator wherein the reaction temperature is 450-800 ° C and the reaction pressure is 0.01- 0.3MPa, volumetric space velocity is 100-1000h -1 .
  • 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 hydrogen atmosphere (reduction atmosphere)
  • the regenerator is an oxygen-containing atmosphere. If the two are not well separated, There will be great security risks.
  • 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 found that if the volumetric space velocity of the alkane feedstock is increased correspondingly while increasing the reaction pressure, the yield of the lower olefin can be maintained at a level higher or higher than that of the prior art. Level, which does not decrease as previously expected in the prior art, as a result of which, for existing reactors, by increasing the reaction pressure and volumetric space velocity of the reactor in accordance with the teachings of the present invention, Increasing the alkane feedstock throughput of the reactor increases the yield of lower olefins (increasing production of lower 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 process for producing a low-carbon olefin (or a method for increasing production), characterized in that the dehydrogenation reaction is carried out in a process for producing a light olefin by continuously bringing an alkane raw material into contact with a catalyst to cause a dehydrogenation reaction.
  • the reaction pressure P is 0.4-6 MPa, preferably 0.4-3 MPa, more preferably 0.5-2 MPa, most preferably 0.6-2 MPa
  • the volume space velocity H of the dehydrogenation reaction is 100-5000 h -1 , preferably 200-2000 h -1 . Most preferred is 500-1000 h -1 .
  • the reaction pressure P of the dehydrogenation reaction is at least 0.3 MPa higher than the regeneration pressure of the regeneration reaction, preferably at least 0.5 MPa higher, at least 0.7 MPa higher, at least 0.9 MPa higher, at least 1.2 MPa higher, or at least 2.0 MPa higher.
  • the number of reactors used to carry out the dehydrogenation reaction is one or more, and each is independently selected from a fluidized bed reactor, a dense phase bed reactor a riser reactor, an ebullated bed reactor, and a composite form of two or more of these reactors, preferably selected from a fluidized bed reactor, more preferably selected from a bubbling fluidized bed reactor or turbulent fluidization Bed reactor.
  • alkane feedstock is selected from at least one of C 2-12 linear or branched alkanes (preferably selected from C 2-5 straight or branched alkanes) At least one (more preferably selected from at least one of propane and isobutane) or a mixture selected from C 3-12 hydrocarbons, or selected from natural gas condensate, natural gas liquid, catalytic cracking liquid gas, oil field gas At least one of condensate, shale gas condensate, straight run naphtha, shale oil light component, hydrogenated naphtha, coker gasoline, and cracked gasoline.
  • the alkane feedstock is selected from at least one of C 2-12 linear or branched alkanes (preferably selected from C 2-5 straight or branched alkanes) At least one (more preferably selected from at least one of propane and isobutane) or a mixture selected from C 3-12 hydrocarbons, or selected from natural gas condensate, natural gas liquid, catalytic cracking liquid gas, oil field gas At least
  • the catalyst is selected from at least one of a dehydrogenation catalyst, a cracking catalyst, and a dehydrogenation/cracking composite catalyst.
  • reaction conditions of the regeneration reaction comprise: a reaction temperature of 550-750 ° C, preferably 600-700 ° C; a reaction pressure of 0.1-0.5 MPa, preferably 0.1-0.3 MPa;
  • the biocatalyst residence time is 5 to 60 minutes, preferably 6 to 20 minutes; an oxygen-containing 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 dehydrogenation reaction. Preferably, it is increased by 100%, more preferably by 150%, 200%, 500% or 800%, most preferably by 1000% or more.
  • the raw catalyst drawn from the reactor is sent to the catalyst receiver to be produced, and then sent to the catalyst feed tank to be produced through the lock hopper, and then sent from the catalyst feed tank to the regenerator, and in the regenerator. Charring regeneration under an oxygen-containing atmosphere to obtain a regenerated catalyst;
  • the regenerated catalyst is continuously withdrawn from the regenerator to the regenerated catalyst receiver, and then sent to the regenerated catalyst feed tank through the lock hopper and continuously returned to the reactor from the regenerated catalyst feed tank.
  • the method for producing a low carbon olefin of the present invention has the following advantages in comparison with the prior art.
  • the volumetric space velocity of the alkane raw material is increased correspondingly while increasing the reaction pressure, and the low-carbon olefin can be obtained without changing the size and the number of the existing reactor or the reaction device.
  • the rate is maintained at a level that is comparable to or even higher than the prior art, and ultimately a large (up to 1000%) 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 flow chart of a method for producing a low carbon olefin according to an embodiment of the present invention
  • FIG. 2 is a schematic flow chart of a method for producing a low carbon olefin according to a further embodiment of the present invention
  • FIG. 3 is a top view (A) and a front view (B) of one embodiment of the built-in baffle (ie, a plate grid) of FIG. 2.
  • Control valve 19 Control valve 20 Control valve 21 Line 22 Line
  • 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.
  • a process for producing a low-carbon olefin which produces a low-carbon olefin by continuously bringing an alkane raw material into contact with a catalyst to cause a dehydrogenation reaction.
  • the manufacturing method may include the steps of: continuously contacting the alkane feedstock with the catalyst to cause the dehydrogenation reaction, obtaining a low-carbon olefin-rich oil and gas and a catalyst to be produced, at least a part of which The raw catalyst is delivered to the regeneration reaction to obtain a regenerated catalyst, and at least a portion of the regenerated catalyst is recycled to the dehydrogenation reaction.
  • the manufacturing method may further comprise: continuously contacting the preheated alkane feedstock with the catalyst in the reactor and dehydrogenating under dehydrogenation conditions to produce oil and gas rich in low carbon olefins and a carbon-based catalyst; separating the oil and gas from the catalyst to be produced, feeding the separated oil and gas into a product separation and recovery system, continuously withdrawing the catalyst to be produced from the reactor; and conveying the catalyst to be produced from the reactor to be treated
  • the catalyst receiver is produced, it is transported to the catalyst feed tank through the lock hopper, and then transported or stripped from the feed catalyst feed tank to the regenerator, and is subjected to charring regeneration in an oxygen atmosphere under the regenerator.
  • the regenerated catalyst is obtained; the regenerated catalyst is continuously withdrawn from the regenerator to the regenerated catalyst receiver, and then sent to the regenerated catalyst feed tank through the lock hopper, and continuously returned to the reactor from the regenerated catalyst feed tank.
  • 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 active catalyst may be partially oxidized after the charred regeneration of the catalyst through the regenerator, the dehydrogenation reaction of the alkane generates hydrogen, and the regenerated catalyst after charring Even if it has not been subjected to reduction treatment, it can be subjected to a dehydrogenation reaction while being returned to the reactor.
  • the method of the present invention may further comprise: after the regenerated catalyst drawn from the regenerator is sent to the regenerated catalyst feed tank through the lock hopper, and then subjected to a reduction treatment under a reducing atmosphere to obtain a reduction catalyst, and then the reduction catalyst It is continuously returned to the reactor.
  • the alkane feedstock may be a C 2-12 linear or branched alkane (preferably a C 2-5 linear or branched alkane) or a mixture thereof, for example, may be selected from the group consisting of ethane, propane, and iso
  • One or more of butane, n-butane and isopentane may also be selected from the group consisting of natural gas condensate, natural gas liquid, catalytic cracking liquid gas, oil field gas condensate and shale gas condensate. At least one may also be an industrial or natural alkane feedstock monomer or mixture of other sources.
  • the alkane feedstock may also be a mixture of small molecule hydrocarbons.
  • the small molecular hydrocarbon mixture it may be a mixture of C3 to C12 hydrocarbons, for example, may be selected from the group consisting of straight run naphtha, oil field condensate, shale oil light component, hydrogenated naphtha, coking
  • One or more of gasoline and cracked gasoline may also be a mixture of industrial or natural small molecule hydrocarbons from other sources.
  • the number of reactors is one or more and may be of a type well known to those skilled in the art.
  • the reactor it may be a fluidized bed reactor, a dense phase bed reactor, a riser reactor, an ebullated bed reactor or a composite form of two or more of these reactors, wherein a fluidized bed reaction is preferred Device.
  • a fluidized bed reactor a bubbling fluidized bed reactor or a turbulent fluidized bed reactor is more preferable, and a bubbling fluidized bed reactor is more preferable.
  • the fluidized bed reactor may be internally provided with a layered inner baffle to prevent uneven mixing of oil and gas and/or catalyst, so that the oil and gas and/or catalyst are The flat push state is reacted through the reactor to increase the conversion of the alkane feedstock and the selectivity of the desired low carbon olefin;
  • the built-in baffle can be a plate grid, and the plate grid can be installed every 20 to 150 cm
  • One layer is preferably provided with a layer of 50-100 cm, and the distance from the bottom surface of the lowermost panel type grating to the top surface of the uppermost panel type grating may be 5% to 80% of the total height of the internal space of the reactor.
  • the material of the plate grid may be selected from the materials used in the catalytic cracking regenerator gas distributor or the macroporous distribution plate, and the grid shape may be wave Shaped and shaped, the grid has evenly arranged small holes or large holes for the catalyst and gas to pass regularly.
  • the low carbon olefin-rich oil and gas and the catalyst to be produced can be separated by a filter.
  • the regenerated catalyst, the flue gas, or the like may be separated by a filter.
  • the filter may be a metal sintered filter.
  • the metal sintered filter is a well-known porous material which can effectively separate solid particles or powders from gas components and is durable.
  • the type and structure of the metal sintered filter of the present invention is not particularly limited as long as it can effectively separate the oil and gas from the catalyst to be produced, and thus will not be described.
  • the catalyst may be a dehydrogenation catalyst, a cracking catalyst, a dehydrogenation/cracking composite catalyst or a mixture thereof, and each may be of a conventional type well known to those skilled in the art, and the present invention is not particularly limited thereto.
  • the shape of the catalyst is generally microspherical.
  • the dehydrogenation catalyst generally comprises an active component and a carrier.
  • the active component may be metal platinum or chromium oxide
  • the support may be alumina; the alumina is preferably ⁇ -Al 2 O 3 and ⁇ -Al 2 O 3 or a mixture of the two;
  • the metal platinum may be contained in an amount of 0.01% by weight to 1.0% by weight, preferably 0.05% by weight to 0.2% by weight based on the total weight of the catalyst, or when the active component is chromium oxide
  • the content of the chromium oxide may be from 1.0% by weight to 30% by weight, preferably from 8.0% by weight to 20% by weight, and the content of the carrier is an equilibrium amount (i.e., the total weight is 100%).
  • the dehydrogenation catalyst may or may not contain iron oxide and/or tin oxide, and may or may not contain an alkali metal oxide or an alkaline earth metal oxide;
  • the content of iron oxide and/or tin oxide may be 0% by weight to 5.0% by weight, preferably 0.2% by weight to 2% by weight; the content of alkali metal oxide or alkaline earth metal oxide may be 0% by weight to 5.0% by weight.
  • % preferably from 0.5% by weight to 2% by weight
  • the alkali metal oxide may be, for example, potassium oxide
  • the alkaline earth metal oxide may be, for example, magnesium oxide.
  • the dehydrogenation/cracking composite catalyst is well known to those skilled in the art and may include an active component, a cocatalyst component, and a carrier. Due to the preparation of the composite catalyst The dehydrogenation reaction is carried out while the cracking reaction is carried out, so that the active component may include a dehydrogenation functional metal component and a cracking functional molecular sieve.
  • the dehydrogenation functional metal component may be one or more of metals or oxides of Cr, Fe, Pt, Sn, Zn, V and Cu, preferably Cr or Pt.
  • the dehydrogenation functional metal component may be contained in an amount of 0.1 to 30% by weight based on the total weight of the catalyst; and the cracking functional molecular sieve may be at least one of a ZSM type, a Y type molecular sieve and a ? type zeolite, preferably a ZRP zeolite.
  • the weight content may be from 5 to 50% by weight, preferably from 20 to 30% by weight, based on the total weight of the catalyst;
  • the cocatalyst component may be an alkali metal and/or alkaline earth metal oxide, preferably potassium oxide and/or magnesium oxide.
  • the weight content may be 0.1 to 5% by weight based on the total weight of the catalyst;
  • the carrier may be selected from inorganic oxides, for example, may be at least one selected from the group consisting of alumina, silica, and aluminum silicate, preferably crystalline.
  • the aluminum silicate may have a weight content of 15% by weight to 94.8% by weight based on the total weight of the catalyst.
  • the composite catalyst is generally microspherical in shape and can be prepared by industrially common methods such as spray drying, rolling into balls, and the like.
  • the reaction temperature of the dehydrogenation reaction is from 500 to 700 ° C, preferably from 530 to 600 ° C.
  • the reaction pressure P of the dehydrogenation reaction is 0.4 to 6 MPa, preferably 0.4 to 3 MPa, more preferably 0.5 to 2 MPa, most preferably 0.6 to 2 MPa, and the dehydrogenation
  • the volumetric space velocity H of the reaction is from 100 to 5000 h -1 , preferably from 200 to 2000 h -1 , most preferably from 500 to 1000 h -1 .
  • the stringent increase function H is carried out when the dehydrogenation reaction is carried out (in other words, based on an existing reactor or reaction unit, which is intended to substantially increase the yield of light olefins).
  • f(P) f(P) is established.
  • P (unit is MPa) belongs to the interval [0.4, 6.0], preferably belongs to the interval [0.4, 3.0], more preferably belongs to the interval [0.5, 2.0], most preferably belongs to the interval [0.6, 2.0]
  • H (the unit is h -1 ) belongs to the interval [100, 5000], preferably belongs to the interval [200, 2000], and most preferably belongs to the interval [500, 1000].
  • the reaction pressure P increases in proportion to the volumetric space velocity H or increases in accordance with different or the same amplitude, and may sometimes be an equal increase or a synchronous increase until the desired increase in the yield of the low-carbon olefin is achieved.
  • the volume space velocity H generally also preferably reaches the upper limit of a certain numerical interval specified in the foregoing invention (for example, 1000 h). -1 ), but is not limited to this.
  • the reaction pressure P and the volume space velocity H are different within the numerical range or numerical range defined by the present invention, even if the reaction pressure P is increased, the volume air 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 conditions of the regeneration reaction can be well known to those skilled in the art, and the present invention is not particularly limited thereto.
  • the reaction conditions of the regeneration reaction may be: a temperature of 550 to 750 ° C, preferably 600 to 700 ° C, a pressure of 0.1 to 0.5 MPa, preferably 0.1 to 0.3 MPa, and a catalyst residence time of 5 to 60 minutes, preferably 6 to 20 Minute; oxygen-containing atmosphere.
  • the oxygen-containing atmosphere may be air, air diluted with nitrogen, or oxygen-enriched gas as a fluidization medium.
  • the preferred regenerator fluidization medium is air or air diluted with nitrogen, and if necessary, may be supplemented with fuel gas such as refining.
  • the plant is dry to increase the temperature of the catalyst bed in the regenerator.
  • the manufacturing method may further include: transporting the regenerated catalyst drawn from the regenerator to the regenerated catalyst feed tank through the lock hopper, and performing a reduction treatment under a reducing atmosphere to obtain a reduction catalyst to oxidize the catalyst into The high valence metal oxide is reduced to the active dehydrogenation component in a lower state and the reduction catalyst is then continuously returned to the reactor.
  • the conditions of the reduction treatment can be determined depending on the conditions of the catalyst used, which are well known and understood by those skilled in the art, and the present invention need not be described in detail.
  • the reduction treatment may be carried out under the conditions of a temperature of 500 to 600 ° C, a pressure of 0.4 to 2.0 MPa, and a catalyst residence time of 1 to 10 minutes; and the reducing atmosphere may be a fluidized stream containing hydrogen. Medium; the reduced stream may be substantially free of oxygen and contain from 50 to 100% by volume of hydrogen, and may contain from 0 to 50% by volume of refinery dry gas.
  • a catalyst having platinum as an active component when used, a catalyst for a long time after the reaction is subjected to regeneration and charring may require a chlorination renewal process to redistribute the platinum active center, at which time the regenerated catalyst may be fed.
  • the can is used as a chlorination processor.
  • the reaction pressure in the reactor is controlled to be higher than that in the regenerator
  • the regeneration pressure is at least 0.3 MPa high.
  • the reaction pressure P of the dehydrogenation reaction is at least 0.3 MPa higher than the regeneration pressure of the regeneration reaction, preferably at least 0.5 MPa higher, at least 0.7 MPa higher, at least 0.9 MPa higher, at least 1.2 MPa higher, or at least 2.0 higher. MPa.
  • the reaction pressure P of the dehydrogenation reaction is at most 5 MPa higher than the regeneration pressure of the regeneration reaction, at most 3.5 MPa, at most 3 MPa, at most 2.5 MPa, at most 2 MPa, at most 1.5 MPa or Up to 1 MPa high.
  • At least a portion of said catalyst to be produced is delivered to the regeneration reaction by one or more (preferably one or two) lock hoppers, and/or at least a portion of said regenerated catalyst is recycled to said dehydrogenation reaction.
  • the lock hopper allows the catalyst to be safely and efficiently transferred from the high pressure hydrocarbon or hydrogen 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 or hydrogen environment of the reactor. .
  • the reducing atmosphere (hydrogen atmosphere) of the reactor and the regenerated catalyst feed tank for reductive catalyst reduction can be well isolated from the oxygen-containing atmosphere of the scorch regeneration of the regenerator,
  • the operating pressure of the reactor and the regenerator can be flexibly adjusted, in particular, the operating pressure of the reactor can be increased without increasing the operating pressure of the regenerator, thereby increasing the throughput of the device.
  • the lock hopper of the present invention is a type that allows the same material stream to be switched between different atmospheres (e.g., an oxidizing atmosphere and a reducing atmosphere) and/or a different pressure environment (e.g., from high pressure to low pressure, or vice versa). Any device.
  • the step of transporting the catalyst from the reactor (high pressure hydrocarbon environment) to the regenerator (low pressure oxygen environment) by the lock hopper may include: 1.
  • the step of circulating the catalyst from the regenerator (low pressure oxygen environment) to the reactor (high pressure hydrocarbon environment) by the lock hopper may include: 1.
  • the lock hopper can be used only by one, that is, the catalyst to be produced and the regenerated catalyst are transported by using the same lock hopper, or the different catalysts can be separately used to carry out the catalyst as needed. And the transport of the regenerated catalyst, all of which are within the scope of protection of the present invention.
  • the catalyst to be produced which is taken out from the reactor can be continuously conveyed to the catalyst receiver, the regenerated catalyst receiver, the catalyst feed tank and the regenerated catalyst feed tank.
  • the catalyst receiver is sent to the raw catalyst feed tank through the lock hopper, and then continuously sent from the raw catalyst feed tank to the regenerator, and the regenerated catalyst extracted from the regenerator can be continuously sent to the regenerated catalyst.
  • the receiver is then sent to the regenerated catalyst feed tank through the lock hopper, and then continuously sent to the reactor from the regenerated catalyst feed tank, thereby achieving continuous progress of the reaction process and the regeneration process; wherein the regenerated catalyst feed tank can be used as
  • the feed tank can also be used as a reducer for regenerating the catalyst.
  • the oil contained in the catalyst stream to be produced may be stripped to the reactor by hydrogen to avoid loss of material; in the regenerative catalyst receiver, nitrogen or other non-oxygen gas may be used to make The catalyst in the receiver remains fluidized, and on the other hand, the oxygen contained in the regenerated catalyst stream is stripped into the regenerator; likewise, in the feed tank of the catalyst to be produced, air or nitrogen can be used as a boosting catalyst. Gas to keep the catalyst in the feed tank fluidized.
  • the heat required for carrying out the dehydrogenation reaction is mainly provided by a high-temperature regenerated catalyst, and if necessary, a heating means for the raw materials and/or catalysts entering the reactor may be additionally provided.
  • the oil and gas contained in the catalyst stream to be produced in the catalyst receiver to be produced is stripped with hydrogen to the reactor.
  • the manufacturing method may further include returning the unreacted alkane raw material separated by the product separation and recovery system as a raw material to the reactor.
  • the alkane feedstock throughput of the reactor can be greatly increased to correspondingly increase the yield of low-carbon olefins.
  • the yield increase of low-carbon olefins can reach 50%, preferably up to 100%, more preferably up to 150%, 200%, 500% or 800%, even in the most preferred case of the invention up to 1000% or higher.
  • the amount or throughput of the alkane feedstock in the reactor or reactor is increased by maintaining the yield of the lower olefins substantially unchanged or slightly improved compared to the prior art.
  • the yield increase of the low carbon olefins in the present invention is significantly higher.
  • the yield of the lower olefin can be maintained even higher than the level comparable to the prior art, such as generally 38-55%, preferably 43-50%.
  • 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.
  • the preheated feedstock is passed through line 7 through a feed distributor into fluidized bed reactor 1, contacted with a regenerated active regenerated catalyst from line 28, vaporized and reacted and sent to the top of reactor 1 .
  • the reaction oil and a small amount of catalyst particles are separated by a gas-solid separation device, the catalyst particles are returned to the reactor bed, and the separated dehydrogenation product is passed through a line 8 to a subsequent separation system for product separation.
  • the catalyst to be produced in the upper portion of the reactor enters the catalyst receiver 3 to be produced via line 21.
  • the catalyst in the spent catalyst receiver 3 is vaporized by the hydrogen from the line 11 to carry the reacted oil and gas, and then flows through the line 22 and the control valve 15 to the lock hopper 4, and the stripped oil is sent to the reactor 1 via the line 31.
  • the catalyst After the catalyst is subjected to a series of processes of purging, boosting, filling and depressurizing in the lock hopper 4, it flows into the catalyst feed tank 5 via the line 23 and the control valve 18 in sequence, and then passes through the line 24 and the control valve in sequence. After mixing with the air from line 12, it is upgraded via line 25 to the upper middle portion of regenerator 2 (such as a fluidized bed regenerator).
  • regenerator 2 such as a fluidized bed regenerator.
  • the spent catalyst is contacted with an oxygen-containing gas from line 9 in a regenerator 2 and a charring reaction occurs to restore catalyst activity.
  • the regenerated flue gas is discharged from the top of the regenerator 2 via line 10 and vented after heat exchange and catalyst dust recovery systems.
  • the regenerated catalyst is mixed with nitrogen from line 13 via control valve 20, and is then upgraded via line 26 to regenerated catalyst receiver 6.
  • the catalyst in regenerated catalyst receiver 6 is fluidized by nitrogen from line 14 and stripped of oxygen carried by the catalyst.
  • the tank 27 and the control valve 17 are sequentially flowed into the lock hopper 4.
  • the regenerated catalyst undergoes a series of processes of purging, depressurizing, filling and boosting in the lock hopper 4, and then flows into the regenerated catalyst feed tank 40 through the control valve 16 and the line 28, and then flows into the reactor 1 through the line 42. Contact and reaction with the feed from line 7.
  • FIG. 2 is a further embodiment of the present invention, the flow of which is based on FIG. 1, after the regenerated catalyst is discharged from the lock hopper 4, and then flows into the regenerated catalyst feed tank 40 through the control valve 16 and the line 28, respectively. After being reduced by the hydrogen-containing gas from the line 41, it is again introduced into the reactor 1 through the line 42 to be in contact with the raw material. The raw material and the catalyst are contacted and reacted in the reactor 1 in which the plate grid 50 is disposed.
  • the apparatus used in the examples are pressurized fluidized bed apparatus having similar embodiments to those described in the figures to achieve similar reaction and regeneration effects.
  • the catalysts used in Examples 1-12 and Comparative Examples 1-10 were prepared catalysts, which were Cr-Fe-K/Al 2 O 3 catalysts.
  • the preparation process of Cr-Fe-K/Al 2 O 3 catalyst (hereinafter referred to as chromium catalyst) is as follows: First, 780 g of chromium nitrate (analytical grade), 100 g of ferric nitrate (analytical grade), and 80 g of potassium nitrate (analytical grade) solid feed The mixture was stirred for 1 hour in a vertical stirred tank containing 3000 g of distilled water; then, pre-dried 2000 g of ⁇ -Al 2 O 3 was charged into the above vertical stirred tank, thoroughly stirred and immersed for 2 hours; and the slurry in the stirred tank was transferred.
  • the excess clear water is filtered out into the filter tank, and then the catalyst is placed in a drying oven at 200 ° C for drying, which requires at least 2 h; the dried catalyst is placed in a muffle furnace at 520 ° C for 6 h to obtain activation.
  • the Cr-Fe-K/Al 2 O 3 dehydrogenation catalyst was placed in a desiccator for use.
  • Examples 1 to 6 were carried out in accordance with the process shown in Fig. 1, and the raw material used was propane (purity of 99.5% or more), and the above-mentioned Cr-Fe-K/Al2O3 dehydrogenation catalyst was used.
  • the experimental conditions, feedstock feed, low carbon olefin yield and yield data are listed in Table 1.
  • Comparative Example 1 to Comparative Example 4 were carried out in accordance with the process shown in Fig. 1, and the raw material used was propane, and the catalyst used was the same as in Examples 1 to 6.
  • the experimental conditions, feedstock feed, low carbon olefin yield and yield data are also listed in Table 1.
  • Examples 7 to 12 were carried out in accordance with the process shown in Fig. 1, and the raw materials used were isobutane (purity of 99.5% or more), and the above-mentioned Cr-Fe-K/Al 2 O 3 dehydrogenation catalyst was used.
  • the experimental conditions, feedstock feed, low carbon olefin yield and yield data are listed in Table 2.
  • Comparative Example 5 to Comparative Example 10 were carried out in accordance with the procedure shown in Fig. 1, and the starting material used was isobutane, and the catalyst used was the same as in Examples 7 to 12.
  • the experimental conditions, feedstock feed, low carbon olefin yield and yield data are also listed in Table 2.
  • the catalysts used in Examples 13-18 and Comparative Examples 11-15 were prepared as dehydrogenation/cracking composite catalysts.
  • the preparation method was as follows: a certain amount of industrial catalytic cracking catalyst CIP-2 (produced by Sinopec Catalyst Qilu Branch), molecular sieve was weighed.
  • the active component is ZRP molecular sieve, the content is 25% by weight, and the rest is aluminum silicate; then the dehydrogenation active component is immersed on the cracking catalyst by dipping method, and heated under a water bath of 60-70 ° C, using H 2 PtCl 6 (analytically pure), mixed with SnCl 2 (analytical grade) and MgCl 2 (analytical grade), dried at 120 ° C for 12 h, calcined at 550 ° C for 4 h, and dechlorinated with water vapor for 2 h to obtain Pt-Sn-Mg. /ZRP catalyst.
  • the Pt content was 0.2%
  • the Sn content was 1%
  • the Mg content was 0.5%
  • the balance was a ZRP catalyst.
  • Examples 13 to 18 were carried out in accordance with the process shown in Fig. 2, and the raw materials used were straight run naphtha (see Table 3 for the properties), and the above Pt-Sn-Mg/ZRP catalyst was used.
  • the experimental conditions, feedstock feed, low carbon olefin yield and yield data are listed in Table 4.
  • Comparative Example 11 to Comparative Example 15 were carried out in accordance with the procedure shown in Fig. 2, and the raw materials used were straight run naphtha, and the catalysts used were the same as in Examples 13 to 18.
  • the experimental conditions, feedstock feed, low carbon olefin yield and yield data are also listed in Table 4.
  • Example I The raw materials used in Example I were purchased gas sources, which were propane (purity of 99.5% or more), isobutane (purity of 99.5% or more), and a mixture of propane and isobutane (mass ratio of 1:1).
  • the catalyst used in Example I was a catalyst prepared, which was a Cr-Fe-K/Al 2 O 3 catalyst and a Pt-Sn-K/Al 2 O 3 catalyst, respectively.
  • chromium catalyst The preparation process of Cr-Fe-K/Al 2 O 3 catalyst (hereinafter referred to as chromium catalyst) is as follows: First, 780 g of chromium nitrate (analytical grade), 100 g of ferric nitrate (analytical grade), and 80 g of potassium nitrate (analytical grade) solid feed The mixture was stirred for 1 hour in a vertical stirred tank containing 3000 g of distilled water; then, pre-dried 2000 g of ⁇ -Al 2 O 3 was charged into the above vertical stirred tank, thoroughly stirred and immersed for 2 hours; and the slurry in the stirred tank was transferred.
  • the excess clear water is filtered out into the filter tank, and then the catalyst is placed in a drying oven at 200 ° C for drying, which requires at least 2 h; the dried catalyst is placed in a muffle furnace at 520 ° C for 6 h to obtain activation.
  • the Cr-Fe-K/Al 2 O 3 dehydrogenation catalyst was placed in a desiccator for use.
  • Pt-Sn-K/Al 2 O 3 catalyst (hereinafter referred to as platinum catalyst) is as follows: first, 20 g of chloroplatinic acid (analytical grade), 120 g of tin nitrate (analytical grade), and 90 g of potassium nitrate (analytical grade) solid The mixture was poured into a vertical stirred tank containing 2400 g of distilled water and stirred for 1 h; then, pre-dried 2000 g of ⁇ -Al 2 O 3 was charged into the above vertical stirred tank, thoroughly stirred and immersed for 2 h; the slurry in the stirred tank was Transfer to a filter tank to filter out excess clear water, then place the catalyst in a 180 ° C drying oven for drying, this process requires at least 2 h; the dried catalyst is placed in a muffle furnace at 500 ° C for 4 h to obtain activation The Pt-Sn-K/Al 2 O 3 dehydrogenation catalyst was placed in a desic
  • Example I-1 was carried out in accordance with the process shown in Fig. 1, and the raw material used was propane, and the prepared chromium-based catalyst and platinum-based catalyst were used, respectively. Experimental conditions, raw material conversion rates, and product selectivity data are listed in Table I-1.
  • Example I-2 was carried out in accordance with the process shown in Fig. 2, and the raw material used was isobutane, and the prepared chromium-based catalyst and platinum-based catalyst were used, respectively. Experimental conditions, feedstock conversion rates, and product selectivity data are listed in Table I-2.
  • Example I-3 was carried out in accordance with the procedure shown in Fig. 2, and the raw material used was a mixture of propane and isobutane, using the prepared chromium-based catalyst and platinum-based catalyst, respectively. Experimental conditions, feedstock conversion rates, and product selectivity data are listed in Table I-3.
  • the conversion of the raw materials and the target olefins are at a lower reaction temperature and a lower regeneration temperature.
  • the yield can reach the level of the existing industrial dehydrogenation process, and due to the pressure of the reaction system
  • the force is higher than that of the existing industrial device, so 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.
  • feedstocks used in Examples II-1, II-2, and II-3 were hydrogenated naphtha, cracked gasoline, and straight run naphtha, respectively, and the properties are shown in Table II-1.
  • the dehydrogenation/cracking composite catalyst is prepared in the laboratory.
  • the preparation method is as follows: a certain amount of industrial catalytic cracking catalyst CIP-2 (produced by Sinopec Catalyst Qilu Branch) is weighed, and the molecular sieve active component is ZRP molecular sieve, and the content is 25 wt%.
  • the rest is aluminum silicate; then the dehydrogenation active component is immersed on the cracking catalyst by dipping, and heated under a water bath of 60-70 ° C, using H 2 PtCl 6 (analytical grade), SnCl 2 (analytical grade) and The mixture of MgCl 2 (analytical grade) was immersed, dried at 120 ° C for 12 h, calcined at 550 ° C for 4 h, and dechlorinated with water vapor for 2 h to obtain a Pt-Sn-Mg/ZRP catalyst.
  • the Pt content was 0.2%
  • the Sn content was 1%
  • the Mg content was 0.5%
  • the balance was a ZRP catalyst.
  • Example II-1 was carried out according to the process shown in Figure 1.
  • the raw materials used were hydrogenated naphtha.
  • the experimental conditions, raw material conversion rate and product selectivity data are listed in Table II-2.
  • Example II-2 was carried out according to the process shown in Fig. 2, and the raw material used was cracked gasoline.
  • the experimental conditions, raw material conversion rate and product selectivity data are shown in Table II-2.
  • Example II-3 was carried out according to the process shown in Figure 2, and the raw materials used were straight run naphtha.
  • the experimental conditions, raw material conversion rates, and product selectivity data are listed in Table II-2.
  • Example II-1 Example II-2

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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烷烃原料与催化剂接触而发生脱氢反应以制造低碳烯烃的方法中，使所述脱氢反应的反应压力P为0.6-2MPa，所述脱氢反应的体积空速H为500-1000h-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
本发明涉及一种由烷烃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烷烃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。作为其制造方法，利用烷烃原料来制造低碳烯烃的技术越来越受到青睐。
目前，由烷烃原料制造低碳烯烃的技术所采用的反应器主要有固定床反应器和流化床反应器。固定床反应器存在着床层传热效果较差、催化剂的更换和再生相对麻烦、并且反应难以连续进行等缺点，但其优点是，可以实现大的烷烃原料处理量。另外，流化床反应器可以解决固定床反应器存在的前述缺点，但其烷烃原料处理量相比于相同尺寸的固定床反应器而言则要小得多。比如，EP0894781 A1和US7235706 B2公开了一种利用烷烃原料脱氢制造低碳烯烃的方法，该方法采用了流化床反应器和再生器，其中反应温度为450-800℃，反应压力为0.01-0.3MPa，体积空速为100-1000h-1。
已知的是，烷烃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烷烃原料的处理量(比如达到与现有技术的固定床相当的原料处理量)，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，反应器是氢气气氛(还原气氛)，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烷烃原料的体积空速，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体积空速，就可以大幅度地增加该反应器的烷烃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，在通过连续地使烷烃原料与催化剂接触而发生脱氢反应以制造低碳烯烃的方法中，使所述脱氢反应的反应压力P为0.4-6MPa，优选0.4-3MPa，更优选0.5-2MPa，最优选0.6-2MPa，所述脱氢反应的体积空速H为100-5000h-1，优选200-2000h-1，最优选500-1000h-1。
2.根据前述任一方面所述的制造方法，其中在进行所述脱氢反应时，严格增函数H＝f(P)成立，其中P(单位是MPa)属于区间[0.4，6.0]，优选属于区间[0.4，3.0]，更优选属于区间[0.5，2.0]，最优选属于区间[0.6，2.0]，H(单位是h-1)属于区间[100，5000]，优选属于区间[200，2000]，最优选属于区间[500，1000]。
3.根据前述任一方面所述的制造方法，包括以下步骤：
连续地使所述烷烃原料与所述催化剂接触而发生所述脱氢反应，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氢反应，
其中所述脱氢反应的反应压力P比所述再生反应的再生压力至少高0.3MPa，优选至少高0.5MPa、至少高0.7MPa、至少高0.9MPa、至少高1.2MPa或者至少高2.0MPa。
4.根据前述任一方面所述的制造方法，其中用于进行所述脱氢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、以及这些反应器中两种或更多种的复合形式，优选选自流化床反应器，更优选选自鼓泡流化床反应器或者湍流流化床反应器。
5.根据前述任一方面所述的制造方法，其中所述烷烃原料选自C2-12直链或支链烷烃中的至少一种(优选选自C2-5直链或支链烷烃中的至少一种(更优选选自丙烷和异丁烷中的至少一种)或者选自C3-12烃的混合物)，或者选自天然气凝析油、天然气液、催化裂化液化气、油田气凝析液、页岩气凝析液、直馏石脑油、页岩油轻组分、加氢石脑油、焦化汽油和裂化汽油中的至少一种。
6.根据前述任一方面所述的制造方法，其中所述催化剂选自脱氢催化剂、裂化催化剂和脱氢/裂化复合催化剂中的至少一种。
7.根据前述任一方面所述的制造方法，其中所述再生反应的反应条件包括：反应温度550-750℃，优选600-700℃；反应压力0.1-0.5MPa，优选0.1-0.3MPa；待生催化剂停留时间5-60分钟，优选6-20分钟；含氧气氛，优选空气气氛或者氧气气氛。
8.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/或所述再生催化剂。
9.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)闭锁料斗(4)实现所述输送和所述循环。
10.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氢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800％，最优选提高1000％或更高。
11.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烷 烃原料循环至所述脱氢反应的步骤。
12.根据前述任一方面所述的制造方法，包括以下步骤：
连续地将预热后的所述烷烃原料在反应器中与所述催化剂接触并在脱氢条件下发生所述脱氢反应，产生富含低碳烯烃的油气和积碳的待生催化剂；
使油气和待生催化剂分离，将分离后的油气送入产品分离回收系统，将待生催化剂从反应器连续地引出；
将从反应器引出的待生催化剂输送至待生催化剂接收器后，再通过闭锁料斗输送至待生催化剂进料罐，然后从待生催化剂进料罐输送至再生器，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连续地引出到再生催化剂接收器后，再通过闭锁料斗输送至再生催化剂进料罐，并从再生催化剂进料罐连续地返回到所述反应器中。
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烷烃原料的体积空速，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(最高可达1000％)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，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是按照本发明的一种具体实施方式的低碳烯烃制造方法的流程示意图；
图2是按照本发明的一种进一步的具体实施方式的低碳烯烃制造方法的流程示意图；
图3是图2中的内置挡板(即板式格栅)的一种实施方式的俯视图(A)和主视图(B)。
附图标记说明
1 反应器         2 再生器       3 待生催化剂接收器
4 闭锁料斗         5 待生催化剂进料罐      6 再生催化剂接收器
7 管线    8 管线     9 管线     10 管线    11 管线    12管线
13 管线     14 管线    15 控制阀     16 控制阀     17 控制阀
18 控制阀    19 控制阀    20 控制阀     21 管线      22管线
23 管线    24 管线      25 管线      26 管线       27管线
28 管线     29 管线      30 管线       31 管线
40 再生催化剂进料罐      41 管线    42 管线      50 板式 格栅
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，“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。
根据本发明，提供一种低碳烯烃的制造方法，其通过连续地使烷烃原料与催化剂接触而发生脱氢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：连续地使所述烷烃原料与所述催化剂接触而发生所述脱氢反应，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氢反应。
根据本发明，所述制造方法还可以包括：连续地将预热后的烷烃原料在反应器中与催化剂接触并在脱氢条件下发生脱氢反应，产生富含低碳烯烃的油气和(积碳的)待生催化剂；使油气和待生催化剂分离，将分离后的油气送入产品分离回收系统，将待生催化剂从反应器连续地引出；将从反应器引出的待生催化剂输送至待生催化剂接收器后，再通过闭锁料斗输送至待生催化剂进料罐，然后从待生催化剂进料罐输送或汽提输送至再生器，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连续地引出到再生催化剂接收器后，再通过闭锁料斗输送至再生催化剂进料罐，并从再生催化剂进料罐连续地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。
本领域技术人员可以理解的是，虽然待生催化剂经过再生器的烧焦再生后可能存在其中的活性组分被部分氧化的情况，但是由于烷烃脱氢反应会产生氢气，烧焦后的再生催化剂即使未经过还原处理，其返回到反应器后仍可以边被还原边进行脱氢反应。但是，为了更好地 提高催化剂的活性，本发明的方法还可以包括：将从再生器引出的再生催化剂通过闭锁料斗输送至再生催化剂进料罐后，在还原气氛下进行还原处理，得到还原催化剂，然后将该还原催化剂连续地返回到所述反应器中。
根据本发明，所述烷烃原料可以是C2-12直链或支链烷烃(优选C2-5直链或支链烷烃)或者是它们的混合物，例如可以是选自乙烷、丙烷、异丁烷、正丁烷和异戊烷中的一种或多种，也可以是选自天然气凝析油、天然气液、催化裂化液化气、油田气凝析液和页岩气凝析液中的至少一种，还可以是其它来源的工业或天然烷烃原料单体或混合物。
根据本发明，所述烷烃原料还可以是小分子烃类混合物。作为所述小分子烃类混合物，可以是C3～C12烃类的混合物，例如，可以是选自直馏石脑油、油田凝析液、页岩油轻组分、加氢石脑油、焦化汽油和裂化汽油中的一种或多种，也可以是其它来源的工业或天然小分子烃类混合物。
根据本发明，所述的反应器的数目是一个或多个，并且可以是本领域所属技术人员所熟知的类型。作为所述反应器，可以是流化床反应器、密相床反应器、提升管反应器、沸腾床反应器或者这些反应器中两种或更多种的复合形式，其中优选流化床反应器。作为所述流化床反应器，更优选的是鼓泡流化床反应器或者湍流流化床反应器，更优选的是鼓泡流化床反应器。
根据本发明的一种具体实施方式，所述流化床反应器内部可以设置分层布置的内置挡板，用来阻止油气和/或催化剂的不均匀的混流，使油气和/或催化剂以一种平推流式的状态通过反应器进行反应，以提高烷烃原料的转化率和所需低碳烯烃的选择性；所述内置挡板可为板式格栅，板式格栅可每20～150cm装设一层，优选为50～100cm装设一层，从最下面板式格栅的底面至最上面板式格栅的顶面之间的距离可为反应器内部空间总高度的5％～80％，优选为20％～70％，进一步优选为30％～50％；所述的板式格栅的材质可选自催化裂化再生器气体分布器或大孔分布板所使用材料，格栅形状可以为波浪形等形状，格栅上有均匀布置的供催化剂和气体规则通过的小孔或大孔。
为了使反应器中反应后产生的油气和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，所述过滤器可以是金属烧结过滤器。在此，金属烧结过滤器是一种公知的多孔材料，可以有效地将固体颗粒或粉末与气体组分进行分离，而且坚固耐用。本发明对所述金属烧结过滤器的种类和结构没有特别的限制，只要其能够有效地将所述油气与待生催化剂进行分离即可，因而不进行赘述。
根据本发明，所述的催化剂可以是脱氢催化剂、裂化催化剂、脱氢/裂化复合催化剂或其混合物，并且各自可以使用本领域所属技术人员所熟知的常规类型，本发明对其没有特别的限制。为了满足反应器和再生器的操作要求，所述催化剂的形状一般为微球形。
根据本发明，所述脱氢催化剂一般含有活性组分和载体。按照本发明的一种具体实施方式，例如，所述活性组分可以为金属铂或者氧化铬，所述载体可以为氧化铝；所述氧化铝优选为γ-Al2O3和θ-Al2O3或二者的混合物；以催化剂的总重量为基准，所述金属铂的含量可以为0.01重％～1.0重％，优选为0.05重％～0.2重％，或者当活性组分为氧化铬时所述氧化铬的含量可以为1.0重％～30重％，优选为8.0重％～20重％，所述载体的含量为平衡量(即总重量为100％)。根据本发明的一种具体实施方式，所述脱氢催化剂可以含或不含氧化铁和/或氧化锡，并且可以含或不含碱金属氧化物或碱土金属氧化物；以催化剂的总重量为基准，氧化铁和/或氧化锡的含量可以为0重％～5.0重％，优选为0.2重％～2重％；碱金属氧化物或碱土金属氧化物的含量可以为0重％～5.0重％，优选为0.5重％～2重％，所述碱金属氧化物例如可以为氧化钾，所述碱土金属氧化物例如可以为氧化镁。
根据本发明，所述脱氢/裂化复合催化剂是本领域技术人所熟知的，可以包括活性组分、助催化剂组分和载体。由于所述复合催化剂的作 用是进行脱氢反应，同时进行裂化反应，故所述活性组分可以包括脱氢功能金属组分和裂化功能分子筛。根据本发明一种优选的实施方式，所述脱氢功能金属组分可以是Cr、Fe、Pt、Sn、Zn、V和Cu的金属或氧化物的一种或多种，优选为Cr或Pt及其氧化物，脱氢功能金属组分重量含量可以是催化剂总重量的0.1～30重％；裂化功能分子筛可以是ZSM型、Y型分子筛和β型沸石中的至少一种，优选为ZRP沸石，重量含量可以是催化剂总重量的5～50重％，优选为20～30重％；所述助催化剂组分可以是碱金属和/或碱土金属氧化物，优选自氧化钾和/或氧化镁，重量含量可以是催化剂总重量的0.1～5重％；所述载体可以选自无机氧化物，例如，可以是选自氧化铝、氧化硅和硅酸铝中的至少一种，优选为结晶型硅酸铝，重量含量可以是催化剂总重量的15重％～94.8重％。为了满足流化床反应器和再生器的操作要求，所述复合催化剂的形状一般为微球形，可以采用喷雾干燥、滚动成球等工业上常用方法制备。
根据本发明，一般地，所述脱氢反应的反应温度为500～700℃，优选530～600℃。特别地，为了实现本发明的低碳烯烃增产目的，所述脱氢反应的反应压力P为0.4-6MPa，优选0.4-3MPa，更优选0.5-2MPa，最优选0.6-2MPa，并且所述脱氢反应的体积空速H为100-5000h-1，优选200-2000h-1，最优选500-1000h-1。
根据本发明一个特别优选的实施方式，在进行所述脱氢反应(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[0.4，6.0]，优选属于区间[0.4，3.0]，更优选属于区间[0.5，2.0]，最优选属于区间[0.6，2.0]，H(单位是h-1)属于区间[100，5000]，优选属于区间[200，2000]，最优选属于区间[500，1000]。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氢反应的反应压力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氢反应的体积空速H。在此，本发明对所述反应压力P和所述体积空速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与体积空速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2MPa)时，体积空速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1000h-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和体积空速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的同时同样增加体积空速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，所述再生反应的条件可以采用本领域技术人员所熟知的，本发明对其没有特别的限制。例如，所述再生反应的反应条件可以为：温度550～750℃，优选600-700℃，压力为0.1～0.5MPa，优选0.1-0.3MPa，催化剂停留时间为5～60分钟，优选6-20分钟；含氧气氛。所述的含氧气氛可以为以空气、以氮气稀释的空气、或者富氧气体作为流化介质，优选的再生器流化介质为空气或者以氮气稀释的空气，必要时可以补充燃料气例如炼厂干气以提高再生器中催化剂床层的温度。
根据本发明，所述制造方法还可以包括：将从再生器引出的再生催化剂通过闭锁料斗输送至再生催化剂进料罐，在还原气氛下进行还原处理，得到还原催化剂，以使催化剂中被氧化成的高价态金属氧化物还原为低价态的活性脱氢组分，然后将该还原催化剂连续地返回到所述反应器中。所述还原处理的条件可以根据所使用的催化剂的情况来确定，这是本领域技术人员所熟知和理解的，本发明对此无需进行详细描述。例如，所述还原处理的条件可以是：温度为500～600℃，压力为0.4～2.0MPa，催化剂停留时间为1～10分钟；所述的还原气氛可以是以含氢气的还原物流作为流化介质；该还原物流可以基本不含氧气并且含有50～100体积％的氢气，并且可以含有0～50体积％的炼油厂干气。另外，当使用以铂为活性组分的催化剂时，长时间反应后的催化剂经过再生烧焦后可能需要有氯化更新过程，以重新分配铂活性中心，此时可以将所述再生催化剂进料罐用作氯化处理器来使用。
根据本发明的一种优选的实施方式，在根据本发明提供的由烷烃原料制造低碳烯烃的方法中，控制反应器中的反应压力比再生器中的 再生压力至少高0.3MPa。具体而言，所述脱氢反应的反应压力P比所述再生反应的再生压力至少高0.3MPa，优选至少高0.5MPa、至少高0.7MPa、至少高0.9MPa、至少高1.2MPa或者至少高2.0MPa。或者，根据本发明，所述脱氢反应的反应压力P比所述再生反应的再生压力至多高5MPa、至多高3.5MPa、至多高3MPa、至多高2.5MPa、至多高2MPa、至多高1.5MPa或者至多高1MPa。
根据本发明，通过一个或多个(优选一个或两个)闭锁料斗，将至少一部分所述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循环至所述脱氢反应。在此，所述的闭锁料斗可使催化剂从反应器的高压烃或氢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反应器的高压烃或氢环境安全和有效地转移。也就是说，通过使用闭锁料斗，一方面可以使反应器以及用于再生催化剂还原的再生催化剂进料罐的还原气氛(氢气气氛)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反应器的操作压力从而提高装置的处理量。
本发明所述的闭锁料斗是一种可使同一物料流在不同的气氛(例如氧化气氛和还原气氛)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装置。通过闭锁料斗可以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闭锁料斗中残存的氧吹扫到再生器中；2、采用氢气将氮气从闭锁料斗吹扫出去；3、采用氢气对已排空的闭锁料斗加压；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闭锁料斗中；5、通过排出加压闭锁料斗内的氢气，对填充的闭锁料斗减压；6、用热氮气将氢气从填充的闭锁料斗吹扫出去；7、将待生催化剂从填充的闭锁料斗排放到待生催化剂进料罐。通过闭锁料斗可以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闭锁料斗吹扫到再生器中；2、采用氢气将氮气从闭锁料斗吹扫出去；3、采用氢气对填充的闭锁料斗加压；4、将再生催化剂从填充的闭锁料斗排放到再生催化剂进料罐；5、通过排出加压闭锁料斗内的氢气，对已排空的闭锁料斗减压；6、用热氮气将氢气从已排空的 闭锁料斗吹扫出去；7、将再生催化剂从再生器接收器填充到已排空的闭锁料斗。
根据本发明的一种具体实施方式，所述闭锁料斗可以只使用一个，即待生催化剂和再生催化剂使用同一个闭锁料斗进行输送，也可以根据需要使用不同的闭锁料斗分别进行所述待生催化剂和所述再生催化剂的输送，该等变化均属于本发明的保护范围。
根据本发明的一种具体实施方式，通过设置待生催化剂接收器、再生催化剂接收器、待生催化剂进料罐及再生催化剂进料罐，可将从反应器引出的待生催化剂连续地输送至待生催化剂接收器后再通过闭锁料斗输送至待生催化剂进料罐，然后从待生催化剂进料罐连续地输送至再生器，以及可将从再生器引出的再生催化剂连续地输送至再生催化剂接收器后再通过闭锁料斗输送至再生催化剂进料罐，然后从再生催化剂进料罐连续地输送至反应器，从而实现反应过程和再生过程的连续进行；其中的再生催化剂进料罐既可当进料罐使用，也可当再生催化剂的还原器使用。在待生催化剂接收器中，可用氢气将待生催化剂物流所含的油气汽提至所述反应器中以避免物料的损失；在再生催化剂接收器中，可用氮气或其它非氧气体一方面使接收器内催化剂保持流化，另一方面将再生催化剂物流所含的氧气汽提至所述再生器中；同样地，在待生催化剂进料罐中，可以用空气或氮气作为提升催化剂的提升气，以保持进料罐内催化剂处于流化状态。
在本发明中，进行脱氢反应所需的热量主要由高温的再生催化剂提供，如果需要，也可另外设置针对进入反应器的原料和/或催化剂的加热装置。
根据本发明的一个实施方式，将待生催化剂接收器中的待生催化剂物流所含的油气用氢气汽提至所述反应器。
根据本发明，所述制造方法还可以包括：将经所述产品分离回收系统分离得到的未反应的烷烃原料作为原料返回至所述反应器中。
根据本发明，在维持用于进行所述脱氢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体积空速，就可以大幅度地增加该反应器的烷烃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800％，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烷烃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烷烃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38-55％，优选43-50％。
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，以下以流化床反应器作为反应器的例子，但本发明并不限于此。
图1提供的烷烃原料制造低碳烯烃的方法的流程如下：
如图1所示，预热后的原料经管线7经过原料分配器进入流化床反应器1，与来自管线28的恢复活性的再生催化剂接触、气化和反应后输送至反应器的1顶部。在反应器1顶部，反应油气和少量催化剂颗粒经气固分离设备分离，催化剂颗粒返回到反应器床层，分离后的脱氢产物经管线8进入后续分离系统进行产品分离。反应器上部的待生催化剂经管线21进入待生催化剂接收器3。待生催化剂接收器3中的催化剂经来自管线11的氢气汽提出携带的反应油气后，依次经过管线22和控制阀15流入闭锁料斗4，汽提出的油气经管线31送入反应器1。
待生催化剂在闭锁料斗4中经历一系列吹扫、升压、填充和降压等过程后，依次经管线23和控制阀18流入待生催化剂进料罐5，随后依次经管线24和控制阀19与来自管线12的空气混合后，经管线25提升至再生器2(比如流化床再生器)的中上部。待生催化剂在再生器2中与来自管线9的含氧气体接触并发生烧焦反应，以恢复催化剂活性。 再生烟气经管线10由再生器2顶部排出并经换热和催化剂粉尘回收系统后放空。再生催化剂经控制阀20与来自管线13的氮气混合后，经管线26提升至再生催化剂接收器6，再生催化剂接收器6中的催化剂经来自管线14的氮气流化并汽提出催化剂携带的氧气后，依次经管线27和控制阀17流入闭锁料斗4。
再生催化剂在闭锁料斗4中经历一系列吹扫、降压、填充和升压等过程后，依次经控制阀16和管线28先流入再生催化剂进料罐40，再通过管线42流入反应器1中，与来自管线7的原料接触和反应。
图2是本发明的一种进一步的具体实施方式，其流程是在图1的基础上，再生催化剂由闭锁料斗4排出后，依次经控制阀16和管线28先流入再生催化剂进料罐40，由来自管线41的含氢气体还原后，再通过管线42流入反应器1中与原料接触。原料和催化剂在布置有板式格栅50的反应器1中接触和反应。
实施例
以下的实施例将结合附图对发明的具体实施方式进行说明。
实施例所使用的装置均为加压流化床装置，具有与附图所述装置相似的实施方式，以达到类似的反应和再生效果。
实施例1-12和对比例1-10所用催化剂为制备的催化剂，为Cr-Fe-K/Al2O3催化剂。Cr-Fe-K/Al2O3催化剂(以下简称铬系催化剂)制备过程如下：首先，将780g硝酸铬(分析纯)、100g硝酸铁(分析纯)、80g硝酸钾(分析纯)固体投料到盛有3000g蒸馏水的立式搅拌罐中，搅拌1h；然后，将预先干燥好的2000gγ-Al2O3投料到上述立式搅拌罐中，充分搅拌和浸渍2h；将搅拌罐中的浆液转移到过滤罐中过滤掉多余明水，然后将催化剂放置到200℃的干燥箱中烘干，此过程需要至少2h；将干燥好的催化剂放置到520℃的马弗炉中焙烧6h，制得活化的Cr-Fe-K/Al2O3脱氢催化剂，放到干燥器中备用。
实施例1～实施例6
实施例1～实施例6按图1所示工艺进行，所用原料为丙烷(纯度99.5％以上)，使用上述Cr-Fe-K/Al2O3脱氢催化剂。实验条件、原料进料量、低碳烯烃收率和产量数据列于表1。
对比例1～对比例4
对比例1～对比例4按图1所示工艺进行，所用原料为丙烷，所使用催化剂与实施例1～6相同。实验条件、原料进料量、低碳烯烃收率和产量数据同样列于表1。
实施例7～实施例12
实施例7～实施例12按图1所示工艺进行，所用原料为异丁烷(纯度99.5％以上)，使用上述Cr-Fe-K/Al2O3脱氢催化剂。实验条件、原料进料量、低碳烯烃收率和产量数据列于表2。
对比例5～对比例10
对比例5～对比例10按图1所示工艺进行，所用原料为异丁烷，所使用催化剂与实施例7～12相同。实验条件、原料进料量、低碳烯烃收率和产量数据同样列于表2。
表1
Figure PCTCN2015000704-appb-000001
表2
Figure PCTCN2015000704-appb-000002
实施例13-18和对比例11-15所用催化剂为制备的脱氢/裂化复合催化剂，制备方法如下：称取一定量工业用催化裂解催化剂CIP-2(中国石化催化剂齐鲁分公司生产)，分子筛活性组分为ZRP分子筛，含量为25重％，其余为硅酸铝；然后采用浸渍法将脱氢活性组分浸渍在裂解催化剂上，在60～70℃的水浴加热下，用H2PtCl6(分析纯)、SnCl2(分析纯)和MgCl2(分析纯)的混合液浸渍，在120℃温度下干燥12h，550℃焙烧4h，通水蒸气除氯2h，制得Pt-Sn-Mg/ZRP催化剂。其中Pt含量为0.2％，Sn含量为1％，Mg含量为0.5％，其余为ZRP催化剂。
实施例13～实施例18
实施例13～实施例18按图2所示工艺进行，所用原料为直馏石脑油(性质见表3)，使用上述Pt-Sn-Mg/ZRP催化剂。实验条件、原料进料量、低碳烯烃收率和产量数据列于表4。
对比例11～对比例15
对比例11～对比例15按图2所示工艺进行，所用原料为直馏石脑油，所使用催化剂与实施例13～18相同。实验条件、原料进料量、低碳烯烃收率和产量数据同样列于表4。
表3
Figure PCTCN2015000704-appb-000003
表4
Figure PCTCN2015000704-appb-000004
实施例I
实施例I所用原料为购买的气源，分别为丙烷(纯度99.5％以上)、异丁烷(纯度99.5％以上)、丙烷与异丁烷混合物(质量比1∶1)。
实施例I所用催化剂为制备的催化剂，分别为Cr-Fe-K/Al2O3催化剂和Pt-Sn-K/Al2O3催化剂。
Cr-Fe-K/Al2O3催化剂(以下简称铬系催化剂)制备过程如下：首先，将780g硝酸铬(分析纯)、100g硝酸铁(分析纯)、80g硝 酸钾(分析纯)固体投料到盛有3000g蒸馏水的立式搅拌罐中，搅拌1h；然后，将预先干燥好的2000gγ-Al2O3投料到上述立式搅拌罐中，充分搅拌和浸渍2h；将搅拌罐中的浆液转移到过滤罐中过滤掉多余明水，然后将催化剂放置到200℃的干燥箱中烘干，此过程需要至少2h；将干燥好的催化剂放置到520℃的马弗炉中焙烧6h，制得活化的Cr-Fe-K/Al2O3脱氢催化剂，放到干燥器中备用。
Pt-Sn-K/Al2O3催化剂(以下简称铂系催化剂)制备过程如下：首先，将20g氯铂酸(分析纯)、120g硝酸锡(分析纯)、90g硝酸钾(分析纯)固体投料到盛有2400g蒸馏水的立式搅拌罐中，搅拌1h；然后，将预先干燥好的2000gγ-Al2O3投料到上述立式搅拌罐中，充分搅拌和浸渍2h；将搅拌罐中的浆液转移到过滤罐中过滤掉多余明水，然后将催化剂放置到180℃干燥箱中烘干，此过程需要至少2h；将干燥好的催化剂放置到500℃的马弗炉中焙烧4h，制得活化的Pt-Sn-K/Al2O3脱氢催化剂，放到干燥器中备用。
实施例I-1
实施例I-1按图1所示工艺进行，所用原料为丙烷，分别使用制备的铬系催化剂和铂系催化剂。实验条件、原料转化率以及产品选择性数据列于表I-1。
实施例I-2
实施例I-2按图2所示工艺进行，所用原料为异丁烷，分别使用制备的铬系催化剂和铂系催化剂。实验条件、原料转化率以及产品选择性数据列于表I-2。
实施例I-3
实施例I-3按图2所示工艺进行，所用原料为丙烷和异丁烷混合物，分别使用制备的铬系催化剂和铂系催化剂。实验条件、原料转化率以及产品选择性数据列于表I-3。
从表I-1、表I-2和表I-3可以看出，采用本发明的流化床反应-再生系统，在反应温度和再生温度较低的条件下，原料转化率和目标烯烃的收率能够达到现有工业脱氢工艺的水平，并且由于反应系统的压 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-1
Figure PCTCN2015000704-appb-000005
表I-2
Figure PCTCN2015000704-appb-000006
表I-3
Figure PCTCN2015000704-appb-000007
实施例II
实施例II-1、II-2、II-3所用原料油分别为加氢石脑油、裂化汽油和直馏石脑油，性质如表II-1所示。
脱氢/裂化复合催化剂为实验室制备，制备方法如下：称取一定量工业用催化裂化催化剂CIP-2(中国石化催化剂齐鲁分公司生产)，分子筛活性组分为ZRP分子筛，含量为25重％，其余为硅酸铝；然后采用浸渍法将脱氢活性组分浸渍在裂化催化剂上，在60～70℃的水浴加热下，用H2PtCl6(分析纯)、SnCl2(分析纯)和MgCl2(分析纯)的混合液浸渍，在120℃温度下干燥12h，550℃焙烧4h，通水蒸气除氯2h，制得Pt-Sn-Mg/ZRP催化剂。其中Pt含量为0.2％，Sn含量为1％，Mg含量为0.5％，其余为ZRP催化剂。
实施例II-1
实施例II-1按图1所示工艺进行，所用原料为加氢石脑油，实验条件、原料转化率以及产品选择性数据列于表II-2。
实施例II-2
实施例II-2按图2所示工艺进行，所用原料为裂化汽油，实验条件、原料转化率以及产品选择性数据列于表II-2。
实施例II-3
实施例II-3按图2所示工艺进行，所用原料为直馏石脑油，实验条件、原料转化率以及产品选择性数据列于表II-2。
从表II-2可以看出，采用本发明的流化床反应-再生系统，在反应温度和再生温度都较低的条件下，裂化气收率和(C2 ＝+C3 ＝)收率能够达到现有工业脱氢工艺的水平，并且由于反应系统的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II-1
原料油 加氢石脑油 裂化汽油 直馏石脑油
密度(20℃)，kg/m3 715 720 679
初馏点，℃ 40 40 39
终馏点，℃ 188 199 186
饱和蒸汽压(20℃)，KPa 53 54 50
烷烃，wt％ 58.2 33.5 57.6
环烷烃，wt％ 34.2 12.6 33.2
芳烃，wt％ 7.6 23.6 9.2
表II-2
项目 实施例II-1 实施例II-2 实施例II-3
反应温度，℃ 565 575 570
反应压力，MPa 0.5 0.9 0.7
原料油体积空速，h-1 300 700 500
再生温度，℃ 660 670 660
再生压力，MPa 0.2 0.2 0.2
裂化气收率，wt％ 76.5 55.3 67.6
裂化气中(C2 ＝+C3 ＝)收率，wt％ 49.6 38.9 46.3

Claims (12)

  1. 一种低碳烯烃的制造方法，其特征在于，在通过连续地使烷烃原料与催化剂接触而发生脱氢反应以制造低碳烯烃的方法中，使所述脱氢反应的反应压力P为0.4-6MPa，优选0.4-3MPa，更优选0.5-2MPa，最优选0.6-2MPa，所述脱氢反应的体积空速H为100-5000h-1，优选200-2000h-1，最优选500-1000h-1。
  2. 根据权利要求1所述的制造方法，其中在进行所述脱氢反应时，严格增函数H＝f(P)成立，其中P(单位是MPa)属于区间[0.4，6.0]，优选属于区间[0.4，3.0]，更优选属于区间[0.5，2.0]，最优选属于区间[0.6，2.0]，H(单位是h-1)属于区间[100，5000]，优选属于区间[200，2000]，最优选属于区间[500，1000]。
  3. 根据权利要求1所述的制造方法，包括以下步骤：
    连续地使所述烷烃原料与所述催化剂接触而发生所述脱氢反应，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氢反应，
    其中所述脱氢反应的反应压力P比所述再生反应的再生压力至少高0.3MPa，优选至少高0.5MPa、至少高0.7MPa、至少高0.9MPa、至少高1.2MPa或者至少高2.0MPa。
  4. 根据权利要求1所述的制造方法，其中用于进行所述脱氢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、以及这些反应器中两种或更多种的复合形式，优选选自流化床反应器，更优选选自鼓泡流化床反应器或者湍流流化床反应器。
  5. 根据权利要求1所述的制造方法，其中所述烷烃原料选自C2-12直链或支链烷烃中的至少一种(优选选自C2-5直链或支链烷烃中的至少一种(更优选选自丙烷和异丁烷中的至少一种)或者是C3-12烃的混合物)，或者选自天然气凝析油、天然气液、催化裂化液化气、油田气凝析液、页岩气凝析液、直馏石脑油、页岩油轻组分、加氢石脑油、焦化汽油和裂化汽油中的至少一种。
  6. 根据权利要求1所述的制造方法，其中所述催化剂选自脱氢催化剂、裂化催化剂和脱氢/裂化复合催化剂中的至少一种。
  7. 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550-750℃，优选600-700℃；反应压力0.1-0.5MPa，优选0.1-0.3MPa；待生催化剂停留时间5-60分钟，优选6-20分钟；含氧气氛，优选空气气氛或者氧气气氛。
  8. 根据权利要求3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/或所述再生催化剂。
  9. 根据权利要求3所述的制造方法，其中通过一个或多个(优选一个或两个)闭锁料斗(4)实现所述输送和所述循环。
  10. 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氢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800％，最优选提高1000％或更高。
  11. 根据权利要求1所述的制造方法，还包括将未反应完全的烷烃原料循环至所述脱氢反应的步骤。
  12. 根据权利要求1所述的制造方法，包括以下步骤：
    连续地将预热后的所述烷烃原料在反应器中与所述催化剂接触并在脱氢条件下发生所述脱氢反应，产生富含低碳烯烃的油气和积碳的待生催化剂；
    使油气和待生催化剂分离，将分离后的油气送入产品分离回收系统，将待生催化剂从反应器连续地引出；
    将从反应器引出的待生催化剂输送至待生催化剂接收器后，再通过闭锁料斗输送至待生催化剂进料罐，然后从待生催化剂进料罐输送至再生器，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连续地引出到再生催化剂接收器后，再通过闭锁料斗输送至再生催化剂进料罐，并从再生催化剂进料罐连续地返回到所述反应器中。
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