WO2020015602A1 - 一种烃油催化裂解方法和系统 - Google Patents

一种烃油催化裂解方法和系统 Download PDF

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WO2020015602A1
WO2020015602A1 PCT/CN2019/095950 CN2019095950W WO2020015602A1 WO 2020015602 A1 WO2020015602 A1 WO 2020015602A1 CN 2019095950 W CN2019095950 W CN 2019095950W WO 2020015602 A1 WO2020015602 A1 WO 2020015602A1
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Prior art keywords
reaction
catalyst
oil
fluidized bed
catalytic cracking
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PCT/CN2019/095950
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English (en)
French (fr)
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龚剑洪
张执刚
魏晓丽
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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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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 US17/261,496 priority Critical patent/US11891578B2/en
Priority to SG11202100254VA priority patent/SG11202100254VA/en
Priority to JP2021502767A priority patent/JP7354228B2/ja
Priority to KR1020217004628A priority patent/KR20210031742A/ko
Publication of WO2020015602A1 publication Critical patent/WO2020015602A1/zh
Priority to US18/545,735 priority patent/US20240117255A1/en

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    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G11/00—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
    • C10G11/14—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts
    • C10G11/18—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts according to the "fluidised-bed" technique
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G11/00—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
    • C10G11/14—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts
    • C10G11/18—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts according to the "fluidised-bed" technique
    • C10G11/182—Regeneration
    • 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
    • B01J8/20—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles with liquid as a fluidising medium
    • B01J8/22—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles with liquid as a fluidising medium gas being introduced into the liquid
    • B01J8/224—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles with liquid as a fluidising medium gas being introduced into the liquid the particles being subject to a circulatory movement
    • B01J8/228—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles with liquid as a fluidising medium gas being introduced into the liquid the particles being subject to a circulatory movement externally, i.e. the particles leaving the vessel and subsequently re-entering it
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    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
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    • C10G11/00—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
    • C10G11/14—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts
    • C10G11/18—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts according to the "fluidised-bed" technique
    • C10G11/187—Controlling or regulating
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
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    • C10G51/00—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes only
    • C10G51/02—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes only plural serial stages only
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
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    • C10G51/00—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes only
    • C10G51/02—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes only plural serial stages only
    • C10G51/026—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes only plural serial stages only only catalytic cracking steps
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J2208/00—Processes carried out in the presence of solid particles; Reactors therefor
    • B01J2208/00796—Details of the reactor or of the particulate material
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/10—Feedstock materials
    • C10G2300/107—Atmospheric residues having a boiling point of at least about 538 °C
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    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/10—Feedstock materials
    • C10G2300/1077—Vacuum residues
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/20—Characteristics of the feedstock or the products
    • C10G2300/201—Impurities
    • C10G2300/205—Metal content
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/20—Characteristics of the feedstock or the products
    • C10G2300/30—Physical properties of feedstocks or products
    • C10G2300/308—Gravity, density, e.g. API
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/40—Characteristics of the process deviating from typical ways of processing
    • C10G2300/4006—Temperature
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/40—Characteristics of the process deviating from typical ways of processing
    • C10G2300/4012—Pressure
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/40—Characteristics of the process deviating from typical ways of processing
    • C10G2300/4018—Spatial velocity, e.g. LHSV, WHSV
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    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/40—Characteristics of the process deviating from typical ways of processing
    • C10G2300/4081—Recycling aspects
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/70—Catalyst aspects
    • C10G2300/701—Use of spent catalysts
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
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    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/80—Additives
    • C10G2300/802—Diluents
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
    • C10G2400/02—Gasoline
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
    • C10G2400/04—Diesel oil
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
    • C10G2400/20—C2-C4 olefins
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
    • C10G2400/28—Propane and butane

Definitions

  • the present application relates to the technical field of catalytic cracking, and in particular, to a method and system for catalytic cracking of a hydrocarbon oil.
  • Low-carbon olefins represented by ethylene and propylene are the most basic raw materials for the chemical industry. Natural gas or light petroleum distillates are used as raw materials at home and abroad. Low-carbon olefins are produced by steam cracking in an ethylene unit. Benzene, toluene, and xylene (BTX) are important basic chemical raw materials, of which paraxylene (PX) accounts for about 45% of total BTX consumption. With the development of China's polyester and other industries, the demand for BTX is expected to continue to grow rapidly. About 90% of ethylene, about 70% of propylene, 90% of butadiene, and 30% of aromatic hydrocarbons are from steam cracking by-products.
  • Chinese Patent Application Publication CN1234426A discloses a method for simultaneously producing low-carbon olefins and high-aromatic gasoline from heavy oil, which comprises making heavy petroleum hydrocarbons and water vapor in a composite reactor composed of a riser and a dense-phase fluidized bed Catalytic cracking is performed to increase the yield of low-carbon olefins, especially propylene, and increase the aromatic content in gasoline to about 80% by weight.
  • Chinese Patent Application Publication CN1393510A discloses a method for catalytic conversion of heavy petroleum hydrocarbons to increase production of ethylene and propylene, which comprises contacting a hydrocarbon oil feedstock with a catalyst containing a five-membered ring high silica zeolite in a riser or fluidized bed reactor. And reaction, the method can not only improve the yield of ethylene and propylene, but also alleviate the hydrothermal deactivation of the catalyst to a certain extent.
  • Chinese Patent Application Publication CN1721510A discloses a method for producing low-carbon olefins and aromatics by using double reaction zone catalytic cracking, which uses different weight hourly space velocities in the two reaction zones to achieve the maximum production of low carbon such as propylene and ethylene from heavy raw materials.
  • U.S. Patent Application Publications US2002003103A and US2002189973A disclose FCC units that use double risers to increase the production of propylene.
  • the gasoline (60-300 ° F / 15-150 ° C) produced by the cracking reaction is fed to a second riser for further reaction. It is a mixture of USY molecular sieve and ZSM-5 molecular sieve catalyst.
  • U.S. patent application publication US2002195373A and international patent application publication WO2017223310A disclose methods using a down-flow reactor at high temperature (1020 to 1200 ° F / 550-650 ° C), short contact time ( ⁇ 0.5 seconds), and large agent oil ratio ( 15 to 25).
  • the main catalyst (Y-type faujasite) has low hydrogen transfer activity, and its formulation is formulated in combination with operating conditions to maximize the yield of light olefins.
  • the high-efficiency separator separates the product from the catalyst in 0.1 seconds, minimizing secondary reactions and coke formation.
  • LCO was used to quench the separated gas product to about 930 ° F / 500 ° C and prevent further cracking.
  • the method disclosed in US patent US6538169A and US patent application publication US2003121825A also constitutes a reaction-regeneration system using two reaction zones and a common regenerator.
  • a high temperature and a high agent-to-oil ratio are used to crack heavy raw materials into light olefins or intermediate products capable of being converted into light olefins.
  • the second reaction zone consists of a second riser, where operating conditions are more demanding and more light components are produced from gasoline products.
  • shape-selective molecular sieves, such as ZSM-5 assists the conversion of gasoline to light olefins.
  • Suitable raw materials include VGO, HVGO and hydrogenated gas oil.
  • Chinese Patent Application Publication CN1403540A discloses a catalytic conversion method for preparing ethylene and propylene, in which a riser and a dense-phase fluidized-bed reactor are used to inject light raw materials into the riser, and the reaction is carried out under a relatively severe scale. The product and the coking catalyst enter the fluidized bed and continue to react under relatively mild conditions. This method has a higher yield of ethylene + propylene + butene.
  • Chinese Patent Application Publication CN102051213A discloses a catalytic cracking method comprising contacting heavy raw materials with a catalyst in a first riser reactor including at least two reaction zones for cracking reaction, and separating light raw materials and cracked heavy oil in a second A step in a riser reactor and a fluidized bed reactor in contact with a catalyst to perform a cracking reaction.
  • the method is used for catalytic cracking of heavy oil.
  • the conversion of heavy oil and propylene yield are high, and the dry gas and coke yields are low.
  • Advanced catalytic cracking technology is a reactor that uses double risers or risers in series dense-phase beds to achieve the goal of producing low-carbon olefins and / or light aromatics under relatively severe reaction conditions.
  • Such reactors When processing slag-containing heavy oil, the problems of high dry gas and high coke yield are unavoidable. When a down-going reactor is used, the reduction of coke yield can be achieved, but the reaction conversion rate is relatively low and a special catalyst is required. With the heavy-weighting of raw materials, the requirements for blending residues in catalytic cracking units are increasing. In order to efficiently use inferior heavy oil resources and meet the growing demand for chemical raw materials such as low-carbon olefins and aromatics, it is necessary to develop Catalytic cracking method for converting inferior heavy oil raw materials into high value-added products.
  • the purpose of the present application is to provide a new method and system for catalytic cracking of hydrocarbon oil, which is particularly suitable for producing low-carbon olefins, such as ethylene and propylene, from catalytic cracking of hydrocarbon oil feedstocks, especially heavy feedstock oils.
  • the method and system of the present application for catalytic cracking have low yields of dry gas and coke and good product distribution.
  • the present application provides a method for catalytic cracking of a hydrocarbon oil, which comprises: A contact reaction step in the reactor, wherein in the rapid fluidized bed, the axial solid fraction ⁇ of the catalyst is controlled in a range of about 0.1 to about 0.2.
  • the present application provides a system suitable for the catalytic cracking of a hydrocarbon oil, especially a heavy feedstock oil.
  • the system includes a catalytic cracking reactor, an oil separation device, an optional reaction product separation device, and regeneration.
  • Device for the catalytic cracking of a hydrocarbon oil, especially a heavy feedstock oil.
  • the catalytic cracking reactor includes a dilute phase transport bed and a rapid fluidized bed in series. According to the flow direction of the reaction material, the dilute phase transport bed is in fluid communication with the fast fluidized bed and the dilute phase transport bed is located in the rapid flow. Upstream of the chemical bed
  • the dilute phase transport bed is provided with a catalyst inlet at the bottom and a first reaction raw material inlet at the lower portion, and the fast fluidized bed is provided with an oil agent outlet at the top and an optional second reaction material inlet at the bottom.
  • the separation equipment is provided with an oil inlet, a catalyst outlet, and a reaction product outlet, and the optional reaction product separation equipment is provided with a reaction product inlet, a dry gas outlet, a liquefied gas outlet, a cracked gasoline outlet, a cracked diesel outlet, and a cracked heavy oil outlet.
  • the regenerator is provided with a catalyst inlet and a catalyst outlet,
  • the catalyst inlet of the dilute phase transport bed is in fluid communication with the catalyst outlet of the regenerator
  • the oil agent outlet of the rapid fluidized bed is in fluid communication with the oil agent inlet of the oil agent separation device
  • the oil agent separation device An outlet of the reaction product is in fluid communication with a reaction product inlet of the optional reaction product separation device
  • a catalyst outlet of the oil separation device is in fluid communication with a catalyst inlet of the regenerator.
  • the present application can effectively increase the catalyst density in the rapid fluidized bed by controlling the axial solid fraction ⁇ of the catalyst in the rapid fluidized bed in a range of about 0.1 to about 0.2, thereby greatly improving the catalyst and hydrocarbon oil feedstock at the instant of the reaction. Ratio and control of relatively long reaction time, so that the catalyst can fully react with hydrocarbon oil feedstock, especially inferior heavy oil. This can not only improve the reaction conversion rate, but also improve the yield of low-carbon olefins and light aromatics. At the same time, it can effectively reduce the generation of dry gas and coke, and improve product distribution and product quality.
  • the ratio of the agent to the oil can be adjusted in a larger range, and more active centers can be provided for the cracking reaction.
  • the introduction of supplementary catalyst enhances the flexibility of reaction temperature adjustment, and can effectively adjust the gradient of temperature and catalyst activity in the rapid fluidized bed.
  • petrochemical enterprises can maximize production of high-value-added chemical raw materials from cheap and inferior heavy oil, and help promote the refining and integration process of oil refining enterprises. Improved economic and social benefits of the petrochemical industry.
  • FIG. 1 is a schematic diagram of a preferred embodiment of the present application.
  • FIG. 2 is a schematic diagram of another preferred embodiment of the present application.
  • any specific numerical value (including the end of the numerical range) disclosed herein is not limited to the exact value of the value, but should be understood to also encompass values close to the exact value, such as within the range of ⁇ 5% of the exact value All possible values. And, for the disclosed numerical range, one or more new ones can be obtained by arbitrarily combining between the endpoint values of the range, between the endpoint values and the specific point values within the range, and between the specific point values. Numerical ranges, these new numerical ranges should also be considered as specifically disclosed herein.
  • the dilute phase transport bed and the rapid fluidized bed constitute two reaction zones connected in series to the reactor, and therefore may also be referred to as a dilute phase transport bed reaction zone and a rapid fluidized bed reaction zone, respectively.
  • dilute phase conveying bed has the meaning well known to those skilled in the art, and specifically refers to the catalyst particles therein forming a dilute phase in a suspended state in a fluid and entrained together by the fluid from the fluidized bed. Fluidized bed form.
  • the term “rapid fluidized bed” has a meaning well known to those skilled in the art, and specifically refers to a fluidized bed form in which catalyst particles are in a rapid fluidized state, where rapid fluidized is a kind of
  • the gas-solid contact fluidization of air bubbles is an important feature that solid particles tend to move in groups.
  • the axial solid fraction ⁇ of the catalyst in the fluidized bed is usually in the range of about 0.05 to about 0.4.
  • the catalyst is usually distributed in a dilute and concentrated manner.
  • the axial solids fraction ⁇ of the upper catalyst may be in the range of about 0.05 to about 0.1
  • the axial solids fraction of the lower catalyst may be in the range.
  • ⁇ may be in the range of about 0.3 to about 0.4.
  • the catalyst in the rapid fluidized bed when the axial solid fraction ⁇ of the catalyst is always controlled in a range of about 0.1 to about 0.2 from bottom to top (that is, in the reaction zone, the upper, When the axial solid fraction ⁇ of the catalyst measured in the middle and lower three parts is greater than or equal to about 0.1 and less than or equal to about 0.2), the catalyst in the rapid fluidized bed exhibits a pseudo-uniformly dense phase distribution. Accordingly, the fast fluidization reaction zone in which the catalyst is present in such a fully concentrated phase distribution may be referred to as a "fully concentrated phase reaction zone".
  • water-to-oil weight ratio refers to the ratio of the weight of the total steam injected into the reactor to the weight of the feedstock.
  • upstream and downstream are both based on the flow direction of the reaction material. For example, when the reactant stream flows from bottom to top, “upstream” indicates a position located below, and “downstream” indicates a position located above.
  • any matter or matter not mentioned applies directly to those known in the art without any change.
  • any embodiment described herein can be freely combined with one or more other embodiments described herein, and the technical solutions or technical ideas formed thereby are regarded as part of the original disclosure or original record of the present invention, and should not be It is considered to be something new that has not been disclosed or anticipated herein unless the person skilled in the art believes that the combination is obviously unreasonable.
  • the present application provides a method for catalytic cracking of a hydrocarbon oil, which comprises bringing a hydrocarbon oil feedstock, especially a heavy feedstock oil, and a catalytic cracking catalyst in a reactor including a series of dilute phase transport beds and a rapid fluidized bed.
  • a contact reaction step wherein in the rapid fluidized bed, the axial solid fraction ⁇ of the catalyst is controlled in a range of about 0.1 to about 0.2.
  • the method according to the present application is used to produce low-carbon olefins, such as ethylene and propylene, from heavy feedstock oils, and further includes the following steps:
  • step ii) subjecting the reaction effluent of step i) and optionally a second reaction feedstock including a light feedstock and / or a heavy feedstock oil to a second catalytic cracking reaction in a rapid fluidized bed,
  • the light feedstock is selected from a C4 hydrocarbon fraction, a C5-C6 light gasoline fraction, and any combination thereof, and at least one of the first and second reaction feedstocks includes the heavy feedstock oil;
  • the axial solid fraction ⁇ of the catalyst is controlled in a range of about 0.1 to about 0.2.
  • the catalyst is prevented from becoming dilute and concentrated in the rapid fluidized bed.
  • the distribution is based on the formula, so that the actual agent-oil ratio above and below the fast fluidized bed is kept consistent, thereby reducing the yield of dry gas and coke and increasing the yield of the target product.
  • the method of the present application further includes the following steps:
  • step ii) subjecting the reaction effluent of step i) and the optional second reaction feedstock to a second catalytic cracking reaction in a rapid fluidized bed under conditions effective to produce lower olefins, wherein in the rapid fluidized bed
  • the axial solid fraction ⁇ of the catalyst is controlled in a range of about 0.1 to about 0.2;
  • the "effective generation of low-carbon olefins” means that at least a portion of the reaction raw materials undergo effective cracking, such as deep cracking, in the rapid fluidized bed to produce low-carbon olefin products such as ethylene and propylene, so that the resulting product mixture Rich in low-carbon olefins.
  • the term “rich olefins” means that the total content of low olefins (such as ethylene and propylene) in the reaction product or product mixture is higher than about 10% by weight of the reaction product or product mixture, It is preferably higher than about 15% by weight, and more preferably higher than about 20% by weight.
  • the method of the present application may further include one or more additional reaction steps, for example, in an additional fluidized bed reaction zone such as Reaction steps such as catalytic cracking and / or catalytic isomerization performed in a dilute phase transport bed, a dense phase fluidized bed, a conventional rapid fluidized bed, and the like.
  • additional fluidized bed reaction zone such as Reaction steps such as catalytic cracking and / or catalytic isomerization performed in a dilute phase transport bed, a dense phase fluidized bed, a conventional rapid fluidized bed, and the like.
  • the method of the present application does not include an additional reaction step before the step i) and after the step ii).
  • the method of the present application further includes the following steps:
  • the first reaction raw material including the preheated inferior heavy oil is introduced from the lower part of the dilute-phase transport bed into the dilute-phase transport bed to contact the catalytic cracking catalyst, and is carried out through the dilute-phase transport bed from bottom to top.
  • a first catalytic cracking reaction to obtain a reaction effluent comprising a first reaction product and a semi-standby catalyst;
  • a second catalytic cracking reaction is performed in the process of a fluidized bed to obtain a reaction effluent including a second reaction product and a catalyst to be produced, wherein the axial solid fraction ⁇ of the catalyst in the rapid fluidized bed is controlled to satisfy: 0.1 ⁇ 0.2;
  • step iv) sending the catalyst to be regenerated into a regenerator for scorch regeneration, and returning at least a part of the obtained regenerated catalyst to step i) as the catalytic cracking catalyst;
  • the obtained second reaction product is separated to obtain dry gas, liquefied gas, cracked gasoline, cracked diesel oil and cracked heavy oil.
  • the method of the present application further includes the following steps:
  • step ii) introducing the reaction effluent of step i) into the bottom of the rapid fluidized bed, and passing the rapid reaction fluid from the bottom of the rapid fluidized bed including the preheated inferior heavy oil through the rapid
  • a second catalytic cracking reaction is performed in the process of a fluidized bed to obtain a reaction effluent including a second reaction product and a catalyst to be produced, wherein the axial solid fraction ⁇ of the catalyst in the rapid fluidized bed is controlled to satisfy: 0.1 ⁇ 0.2;
  • step iv) sending the catalyst to be regenerated into a regenerator for scorch regeneration, and returning at least a part of the obtained regenerated catalyst to step i) as the catalytic cracking catalyst;
  • the obtained second reaction product is separated to obtain dry gas, liquefied gas, cracked gasoline, cracked diesel oil and cracked heavy oil.
  • the method of the present application further comprises: introducing one or more supplementary catalysts into the rapid fluidized bed, and bringing it into contact with the materials in the rapid fluidized bed to perform a catalytic cracking reaction.
  • the carbon content of the one or more supplementary catalysts can be independently about 0-1.0% by weight, for example, the one or more supplementary catalysts can be independently selected from regeneration catalysts and catalysts to be produced. And semi-regenerated catalysts, that is, regenerated, ready-to-use and semi-regenerated catalytic cracking catalysts.
  • the total amount of the one or more supplementary catalysts may account for about 0-50% by weight, preferably about 5-30% by weight, of the reactor catalyst circulation.
  • the distance from the introduction position of the one or more supplementary catalysts to the bottom of the rapid fluidized bed is each independently about 0% to about 90% of the total height of the rapid fluidized bed.
  • the introduction positions of the one or more supplementary catalysts are each independently located at a height of about 20% to about 80%, more preferably at a height of about 30% to about 75% of the rapid fluidized bed.
  • the introduction position may be at the bottom of the rapid fluidized bed, or at about 1/3 of the total height of the rapid fluidized bed.
  • the temperature of the supplementary catalyst can be adjusted according to the required reaction temperature, for example, cold and / or hot regenerated catalyst can be introduced, and cold and / or hot stand-by catalyst can also be introduced.
  • the introduction of a supplementary catalyst in the rapid fluidized bed can adjust the ratio of the agent to the oil in a wide range, and provide more active centers for the cracking reaction.
  • the introduction of supplementary catalysts enhances the flexibility of the reaction temperature adjustment, which can effectively adjust the temperature in the rapid fluidized bed and the gradient of catalyst activity.
  • the introduction of supplementary catalyst in a rapid fluidized bed can maintain the uniformity of catalyst density in the fluidized bed as much as possible, effectively adjust the distribution of catalyst density, and ensure that the cracking reaction proceeds fully and effectively, thereby improving the choice of target products. Sex.
  • the catalyst distribution in the rapid fluidized bed can be further adjusted by adjusting the linear velocity of the gas in the rapid fluidized bed, and / or a catalyst distribution plate is provided in the rapid fluidized bed, thereby making the catalyst look like Uniform full dense phase distribution.
  • the hydrocarbon oil feedstock such as a heavy feedstock oil, especially a poor-quality heavy oil
  • the reactor including a dilute phase transport bed and a rapid fluidized bed at one or more locations.
  • the hydrocarbon oil feedstock can be all introduced into the dilute phase transport bed at one feed location, or all introduced into the fast fluidized bed at one feed location.
  • the hydrocarbon oil feedstock may be introduced into the dilute phase transport bed and / or rapid fluidized bed from two or more feed locations in the same or different proportions, for example, a portion of the hydrocarbon oil feedstock from one A feed position is introduced into the dilute phase transport bed, and another portion of the hydrocarbon oil feed is introduced into the rapid fluidized bed from another feed position, or the hydrocarbon oil feed is introduced from two or more feed positions Is introduced into the dilute phase transport bed, or the hydrocarbon oil feedstock is introduced into the rapid fluidized bed from two or more feed locations.
  • inferior heavy oil refers to heavy oil that is more unsuitable for catalytic cracking processing than conventional heavy oil.
  • the properties of the inferior heavy oil may satisfy at least one of the following indicators, such as one, two, three, or four: the density at 20 ° C. is about 900-1000 kg / m 3 , preferably about 910- 940 kg / m3 ; residual carbon is about 2-10% by weight, preferably about 3-8% by weight; the total content of nickel and vanadium is about 2-30ppm, preferably about 5-20ppm; the characteristic factor K value is less than about 12.1, preferably less than about 12.0.
  • the residual carbon in inferior heavy oil was measured by ASTMD-189 Kang's residual carbon experimental method.
  • the inferior heavy oil may be a heavy petroleum hydrocarbon and / or other mineral oil;
  • the heavy petroleum hydrocarbon may be selected from vacuum residue (VR), inferior atmospheric residue (AR), inferior fuel oil Hydrogen residue oil, coking gas oil, deasphalted oil, vacuum wax oil, high acid crude oil, high metal crude oil and any combination thereof;
  • the other mineral oil may be selected from coal liquefaction oil, oil sand oil, shale oil And any combination of them.
  • catalytic cracking catalyst used in this application.
  • it can be various catalytic cracking catalysts that are well known to those skilled in the art and are suitable for producing low-carbon olefins from hydrocarbon oil feedstocks, such as heavy feedstock oils.
  • the catalytic cracking catalyst comprises about 1-50% by weight, preferably about 5-45% by weight, and more preferably about 10-40 Weight percent zeolite, about 5 to 99 weight percent, preferably about 10 to 80 weight percent, more preferably about 20 to 70 weight percent inorganic oxide, and about 0 to 70 weight percent, preferably about 5 to 60 weight percent, more preferably About 10-50% by weight of clay;
  • the zeolite as an active component may include a mesoporous zeolite and optionally a macroporous zeolite.
  • the mesoporous zeolite comprises about 0-50% by weight, preferably about 0-20% by weight, on a dry basis.
  • the mesoporous zeolite and the macroporous zeolite follow the conventional definition in the art, that is, the average pore diameter of the mesoporous zeolite is about 0.5-0.6 nm, and the average pore diameter of the macroporous zeolite is about 0.7-1.0 nm.
  • the macroporous zeolite may be selected from one or more of rare earth Y (REY) zeolite, rare earth hydrogen Y (REHY) zeolite, ultra-stable Y zeolite and high silicon Y zeolite obtained by different methods. Species.
  • the mesoporous zeolite may be selected from zeolites having an MFI structure, such as ZSM series zeolites and / or ZRP zeolites.
  • the above-mentioned mesoporous zeolite can be modified with non-metal elements such as phosphorus and / or transition metal elements such as iron, cobalt, and nickel.
  • ZRP zeolites A more detailed description of ZRP zeolites can be found in US Patent No.
  • the ZSM series zeolite is preferably a mixture of one or more selected from the group consisting of ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM-38, ZSM-48 and other zeolites of similar structure. .
  • ZSM-5 A more detailed description of ZSM-5 can be found in US Patent No. 3,702,886A.
  • the inorganic oxide is preferably silicon dioxide (SiO 2 ) and / or alumina (Al 2 O 3 ).
  • the clay serves as a matrix (ie, a support), and is preferably kaolin and / or polykaolin.
  • the present application has no strict restrictions on the conditions of the catalytic cracking reaction.
  • it may be a catalytic cracking reaction condition that is well known to those skilled in the art and suitable for producing low-carbon olefins from hydrocarbon oil feedstocks, such as heavy feedstock oils.
  • the conditions for the first catalytic cracking reaction may include: a reaction temperature of about 500-600 ° C., a reaction time of about 0.05-5 seconds, and a weight ratio of agent to oil of about 1: 1 to about 50: 1.
  • Water-oil weight ratio is about 0.03: 1 to about 0.5: 1
  • catalyst density is about 20-100 kg / m3
  • gas line speed is about 4-18 m / s
  • reaction pressure is about 130-450 kPa
  • the catalyst mass flow rate G s is about 180-500 kg / (m 2 ⁇ second).
  • the conditions for the first catalytic cracking reaction include: a reaction temperature of about 520-580 ° C., a reaction time of about 1-3 seconds, a weight ratio of agent to oil of about 5: 1 to about 25: 1, water and oil The weight ratio is from about 0.05: 1 to about 0.3: 1.
  • the conditions of the second catalytic cracking reaction may include: a reaction temperature of about 510-650 ° C., a reaction time of about 1-20 seconds, and a weight ratio of agent to oil of about 3: 1 to about 50: 1.
  • Water-oil weight ratio is about 0.03: 1 to about 0.8: 1
  • catalyst density is about 120-290 kg / m 3
  • gas linear velocity is about 0.8-2.5 m / s
  • reaction pressure is about 130-450 kPa
  • the catalyst mass flow rate G s is about 15-150 kg / (m 2 ⁇ second).
  • the conditions of the second catalytic cracking reaction include: a reaction temperature of about 550-620 ° C., a reaction time of about 3-15 seconds, a weight ratio of agent to oil of about 10: 1 to about 30: 1, water and oil The weight ratio is about 0.05: 1 to about 0.5: 1, the catalyst density is about 150-250 kg / m 3 , the gas linear velocity is about 1-1.8 m / s, the reaction pressure is about 130-450 kPa, and the catalyst mass flow The rate G s is about 20-130 kg / (m 2 ⁇ second).
  • the separation of the reaction product from the catalyst to be produced can be performed in a manner well known to those skilled in the art, for example, a cyclone can be used in the settler.
  • the method for further separating the reaction product to obtain dry gas, liquefied gas, cracked gasoline, cracked diesel oil and cracked heavy oil is also well known to those skilled in the art.
  • the dry gas and the liquefied gas can be further separated by conventional separation means in the art to obtain objective products such as ethylene and propylene.
  • the method according to the present application further comprises: using a C4 hydrocarbon fraction and / or a C5-C6 light gasoline fraction as one or more light raw materials in the first and / or second reaction raw materials. It is introduced into the rapid fluidized bed and / or dilute phase transport bed for catalytic cracking reaction.
  • at least one of the first reaction feed and the second reaction feed comprises a light feed selected from a C4 hydrocarbon fraction, a C5-C6 light gasoline fraction, and any combination thereof.
  • the first reaction raw material includes a light raw material and a heavy raw material oil, and at least a part of the light raw material is introduced into the dilute phase transport bed when the heavy raw material oil is introduced into the dilute phase transport bed. Upstream of the site is introduced into the dilute phase transport bed.
  • the first reaction raw material includes a heavy raw material oil, such as a poor quality heavy oil
  • the second reaction raw material includes the light raw material.
  • the first reaction raw material includes the light raw material
  • the second reaction raw material includes a heavy raw material oil, such as a bad heavy oil.
  • C4 hydrocarbon fraction refers to low-molecular-weight hydrocarbons that exist in the form of gas at normal temperature and pressure with the C4 fraction as the main component, including alkanes, alkenes, and alkynes having 4 carbon atoms in the molecule. hydrocarbon. It may include gaseous hydrocarbon products (such as liquefied gas) rich in C4 hydrocarbon fractions produced by the method of the present application, or gaseous hydrocarbons rich in C4 fractions produced by other devices, preferably C4 hydrocarbons produced by the method of the present application. Fractions.
  • the C4 hydrocarbon fraction is preferably an olefin-rich C4 hydrocarbon fraction, wherein the content of the C4 olefin may be greater than about 50% by weight, preferably greater than about 60% by weight, and more preferably about 70% by weight or more.
  • the "C5-C6 light gasoline fraction” refers to a component having a carbon number of C5 to C6 in gasoline, which may include cracked gasoline produced by the method of the present application, or may include gasoline fractions produced by other devices.
  • it may be a C5-C6 fraction selected from at least one of catalytic cracked gasoline, catalytic cracked gasoline, straight run gasoline, coking gasoline, thermal cracked gasoline, thermal cracked gasoline, and hydrogenated gasoline.
  • the catalyst to be produced can be regenerated in a manner well known to those skilled in the art, for example, it can be burned and regenerated in a regenerator.
  • an oxygen-containing gas such as air can be introduced into the regenerator to be in contact with the catalyst to be produced.
  • the flue gas obtained from the scorch regeneration can be separated from the catalyst in the regenerator and sent to a subsequent energy recovery system.
  • the regenerated catalyst after the coke regeneration by the regenerator may be cooled to about 600-680 ° C through a catalyst cooler, and then returned to the reactor.
  • the hot regenerated catalyst is cooled and returned to the reactor, which helps to reduce the contact temperature of the oil agent, improve the contact state between the feed oil and the catalyst, and then improve the selectivity to dry gas and coke.
  • the rapid fluidized bed includes a fully concentrated phase reaction zone and a transition section in order from bottom to top.
  • the fully concentrated phase reaction zone is substantially circular in cross section, and has a bottom end and a top end.
  • the dilute phase transport bed is in communication with the bottom end of the fully concentrated phase reaction zone, and the top of the fully concentrated phase reaction zone is in communication with the transition section through
  • the outlet section of the reactor is connected, and optionally one or more inlets for feeding the second reaction raw material are optionally provided at the bottom of the fully concentrated phase reaction zone,
  • the cross-sectional diameter at the bottom end of the fully concentrated reaction zone is greater than or equal to the diameter of the dilute phase transport bed, and the cross-sectional diameter at the top is greater than the diameter of the outlet section, and the
  • the bottom or side wall is provided with one or more supplemental catalyst inlets, and the positions of the one or more supplemental catalyst inlets are each independently located at a height of about 0% to about 90% of the total height of the rapid fluidized bed, preferably At a height of about 20% to about 80%, more preferably at a height of about 30% to about 75%.
  • the all-concentrated reaction zone may be of various types of hollow cylinders of equal diameter or variable diameter, such as a substantially circular cross-section, open at the bottom and top ends, such as hollow cylinders of equal diameter, Or a type of hollow cylinder with a continuous or discontinuous increase in diameter from bottom to top.
  • the “continuously increasing diameter” means that the diameter continuously increases in a linear or non-linear manner.
  • an inverted hollow truncated cone may be mentioned.
  • the diameter increases discontinuously means that the diameter increases in a discontinuous manner, such as in a stepped manner.
  • a hollow cylinder composed of two or more cylindrical cylinders having an increasing diameter can be cited.
  • the fully concentrated phase reaction zone may be a hollow cylinder, an inverted hollow truncated cone, a hollow cylinder composed of two or more cylindrical cylinders with increasing diameter, and two or more diameters.
  • a catalyst distribution plate is provided at the bottom of the fully concentrated reaction zone.
  • one or more second reaction raw material inlets are provided at the bottom of the fully concentrated phase reaction zone, and preferably, a gas distributor is provided at a position of the second reaction raw material inlet.
  • the ratio of the diameter of the maximum cross-section of the fully concentrated phase reaction zone to the total height of the rapid reaction bed is about 0.005: 1 to about 1: 1, preferably about 0.01: 1 to about 0.8 1: 1, more preferably from about 0.05: 1 to about 0.5: 1; the ratio of the height of the fully concentrated phase reaction zone to the total height of the rapid reaction bed is about 0.1: 1 to about 0.9: 1, preferably about 0.3: 1 to About 0.85: 1, more preferably about 0.5: 1 to about 08: 1.
  • the all-concentrated reaction zone is an inverted hollow truncated cone
  • the longitudinal section is an isosceles trapezoid
  • the diameter of the cross section at the bottom end is about 0.2-10 meters, preferably about 0.5.
  • the ratio of the ratio of the diameter of the top cross section to the diameter of the bottom cross section is greater than 1 to about 50, preferably about 1.2 to about 10, more preferably about 1.5 to about 5; maximum cross section
  • the ratio of the diameter to the total height of the rapid reaction bed is about 0.005: 1 to about 1: 1, preferably about 0.01: 1 to about 0.8: 1, more preferably about 0.05: 1 to about 0.5: 1; fully concentrated phase
  • the ratio of the height of the reaction zone to the total height of the rapid reaction bed is about 0.1: 1 to about 0.9: 1, preferably about 0.3: 1 to about 0.85: 1, and more preferably about 0.5: 1 to about 0.8: 1.
  • the all-concentrated phase reaction zone is a type of a hollow cylinder composed of an inverted truncated cone and a cylinder, and preferably the truncated cone is located below the cylinder, where
  • the longitudinal section of the truncated cone is an isosceles trapezoid, and the diameter of the cross section at the bottom end is about 0.2-10 meters, preferably about 0.5-8 meters, more preferably about 1-5 meters;
  • the ratio of the cross-sectional diameter is greater than 1 to about 50, preferably about 1.2 to about 10, and more preferably about 1.5 to about 5;
  • the diameter of the cylinder is approximately the same as the diameter of the top cross section of the frusto-conical body, and
  • the ratio of the height of the hollow cylinder to the height of the frusto-conical body is about 0.4: 1 to 2.5: 1, preferably about 0.8: 1 to about 1.5: 1;
  • the ratio of the diameter to the total height of the rapid reaction bed is about 0.005: 1 to about
  • the fully dense phase reaction zone is a hollow cylindrical type with a diameter of about 0.2-10 meters, preferably about 1-5 meters;
  • the ratio of the total height of the rapid reaction bed is about 0.005: 1 to about 1: 1, preferably about 0.01: 1 to about 0.8: 1, more preferably about 0.05: 1 to about 0.5: 1;
  • the ratio of the height to the total height of the rapid reaction bed is about 0.1: 1 to about 0.9: 1, preferably about 0.3: 1 to about 0.85: 1, and more preferably about 0.5: 1 to about 0.8: 1.
  • the height of the full dense phase reaction zone is about 2-50 meters, preferably about 5-40 meters, and more preferably about 8-20 meters.
  • the ratio of the height of the transition section to the total height of the rapid reaction bed is from about 01: 1 to about 0.9: 1, preferably from about 0.2: 1 to about 0.5: 1.
  • the transition section is a hollow truncated cone type
  • the longitudinal section is an isosceles trapezoid
  • the inclination angle ⁇ of the sides of the isosceles trapezoid is about 25-85 °, preferably about 30-75 °.
  • the reactor used in the present application may further include one or more additional fluidized bed reaction zones upstream of the dilute phase transport bed and / or downstream of the rapid fluidized bed, such as Dilute phase transport bed, dense phase fluidized bed, conventional rapid fluidized bed, etc.
  • the reactor used herein does not include an additional reaction zone upstream of the dilute phase transport bed and downstream of the rapid fluidized bed.
  • the reactor used in the present application may be arranged coaxially with the settler, or may be arranged in parallel with the settler.
  • the present application provides a system suitable for the catalytic cracking of a hydrocarbon oil, especially a heavy feedstock oil, the system comprising a catalytic cracking reactor, an oil separation device, an optional reaction product separation device, and Regenerator,
  • the catalytic cracking reactor includes a dilute phase transport bed and a rapid fluidized bed in series. According to the flow direction of the reaction material, the dilute phase transport bed is in fluid communication with the fast fluidized bed and the dilute phase transport bed is located in the rapid flow. Upstream of the chemical bed
  • the dilute phase transport bed is provided with a catalyst inlet at the bottom and a first reaction raw material inlet at the lower portion, and the fast fluidized bed is provided with an oil agent outlet at the top and an optional second reaction material inlet at the bottom.
  • the separation equipment is provided with an oil inlet, a catalyst outlet, and a reaction product outlet, and the optional reaction product separation equipment is provided with a reaction product inlet, a dry gas outlet, a liquefied gas outlet, a cracked gasoline outlet, a cracked diesel outlet, and a cracked heavy oil outlet.
  • the regenerator is provided with a catalyst inlet and a catalyst outlet,
  • the catalyst inlet of the dilute phase transport bed is in fluid communication with the catalyst outlet of the regenerator
  • the oil agent outlet of the rapid fluidized bed is in fluid communication with the oil agent inlet of the oil agent separation device
  • the oil agent separation device An outlet of the reaction product is in fluid communication with a reaction product inlet of the optional reaction product separation device
  • a catalyst outlet of the oil separation device is in fluid communication with a catalyst inlet of the regenerator.
  • the rapid fluidized bed is arranged coaxially with the dilute phase transport bed, and the rapid fluidized bed is located above the dilute phase transport bed.
  • a catalyst distribution plate is provided in the rapid fluidized bed, which may be disposed at the bottom of the rapid fluidized bed, for example, at the connection between the dilute phase transport bed and the rapid fluidized bed. Office.
  • the catalyst distribution plate may be various types of distribution plates common in the industry, for example, one or more of a flat plate shape, an arch shape, a dish shape, a ring shape, and an umbrella shape.
  • a catalyst distribution plate helps to make the catalyst uniformly contact the feedstock oil in the axial direction of the full-thickness reaction zone to conduct the catalytic cracking reaction, thereby reducing the specific coke and thermal reactions caused by the catalyst concentration being too high or too low Defocused generation.
  • one or more second reaction raw material inlets are provided at the bottom of the rapid fluidized bed, and a gas distributor is preferably provided at the position of the inlets.
  • the rapid fluidized bed has the structure as described above, that is, from bottom to top, it includes a fully concentrated phase reaction zone and a transition section, and the fully concentrated phase reaction zone and the transition section Specific settings are not described in detail here.
  • the catalytic cracking reactor may further include one or more additional fluidized bed reaction zones upstream of the dilute phase transport bed and / or downstream of the rapid fluidized bed, such as Dilute phase transport bed, dense phase fluidized bed, conventional rapid fluidized bed, etc.
  • the catalytic cracking reactor does not include additional reaction zones upstream of the dilute phase transport bed and downstream of the rapid fluidized bed.
  • the oil agent separation equipment and the reaction product separation equipment may both adopt equipment well known to those skilled in the art.
  • the oil separation device may include a cyclone, a settler, a stripper, and the like
  • the reaction product separation device may be a fractionation column and the like.
  • FIG. 1 shows a preferred embodiment of the present application, in which the pre-lifted medium enters the bottom of the dilute phase transport bed I from the bottom of the pre-lifted section 2 through the pre-lifted medium pipeline 1, and the pre-lifted medium may be dry gas or water Steam or their mixture.
  • the regenerated catalyst from the regeneration inclined pipe 11 enters the lower part of the pre-lifting section 2 and enters the dilute phase transport bed I under the lifting effect of the pre-lifting medium and moves upward.
  • the first reaction feedstock including C4 hydrocarbon fraction, C5-C6 light gasoline fraction and / or hydrocarbon oil feedstock, such as inferior heavy oil is injected into the lower portion of the dilute phase transport bed I through the first feed line 14 and Some streams are mixed and contacted and subjected to a first catalytic cracking reaction to obtain a reaction effluent comprising a first reaction product and a semi-standby catalyst.
  • the reaction effluent moves upwards, enters the bottom of the rapid fluidized bed II, and comes into contact with a supplementary catalyst introduced through the supplementary line 15, which may be a regenerated catalyst or a catalyst to be grown, and performs a second catalytic cracking reaction.
  • the supplement line 15 is connected at a height of about 0% to about 90% of the rapid fluidized bed II, preferably at a height of about 20% to about 80%, and more preferably at a height of about 30% to about 75%.
  • a second reaction feedstock including a C4 hydrocarbon fraction, a C5-C6 light gasoline fraction, and / or a hydrocarbon oil feedstock, such as an inferior heavy oil is passed to the bottom of the rapid fluidized bed II via the second feed line 16.
  • the reaction effluent containing the second reaction product and the deactivated catalyst to be produced by the reaction enters the cyclone separator 6 in the settler 4 through the outlet section 3 to achieve separation of the catalyst to be produced and the second reaction product.
  • the separated second reaction product enters the gas collection chamber 7, and the reaction products in the gas collection chamber 7 enter the subsequent product separation system (not shown in the figure) through the oil and gas pipeline 8.
  • the fine catalyst powder is returned to the settler 4 by the legs of the cyclone separator 6, and the catalyst to be produced in the settler 4 flows to the stripping section 5.
  • the reaction product steamed out of the catalyst to be produced passes through the cyclone separator 6 and enters the gas collection chamber 7.
  • the stripped catalyst to be regenerated enters the regenerator 10 through the inclined tube 9 to be regenerated, and the air is distributed to the regenerator 10 through the air distributor 13 to burn off the catalyst on the regenerated catalyst in the dense phase bed at the bottom of the regenerator Coke regenerates the deactivated standby catalyst to obtain a regenerated catalyst.
  • the regenerated catalyst returns to the pre-lifting section 2 through the regeneration inclined pipe 11, and the flue gas enters the subsequent energy recovery system (not shown in the figure) through the flue gas line 12.
  • FIG. 2 shows another preferred embodiment of the present application, which is basically the same as the embodiment shown in FIG. 1 except that the rapid fluidized bed II includes a fully concentrated phase reaction zone 17 and a transition from bottom to top in order.
  • Paragraph 18 The full dense phase reaction zone 17 is an inverted hollow truncated cone, and the longitudinal section is an isosceles trapezoid.
  • the transition section 18 is of a hollow truncated cone type, the longitudinal section is an isosceles trapezoid, and the inclination angle ⁇ of the sides of the isosceles trapezoid is about 25-85 °, preferably about 30-75 °.
  • the present application provides the following technical solutions:
  • a method for catalytic cracking using a dilute phase transport bed and a rapid fluidized bed comprising:
  • inferior heavy oil is a heavy petroleum hydrocarbon and / or other mineral oil
  • the heavy petroleum hydrocarbon is selected from the group consisting of vacuum residues, inferior atmospheric residues, inferior hydrogenation residues, coking gas oil, deasphalted oil, vacuum wax oil, high acid crude oil and high metal crude oil.
  • the other mineral oil is one or more selected from the group consisting of coal liquefied oil, oil sands oil, and shale oil.
  • the catalytic cracking catalyst comprises 1-50% by weight of zeolite and 5-99% by weight of inorganic oxidation And 0-70% by weight of clay;
  • the zeolite includes a mesoporous zeolite and an optional macroporous zeolite, the mesoporous zeolite is a ZSM series zeolite and / or a ZRP zeolite, and the macroporous zeolite is selected from the group consisting of rare earth Y, rare earth hydrogen Y, super stable Y and high One or more of silicon Y.
  • the conditions for the second catalytic cracking reaction include: a reaction temperature of 510-650 ° C, a reaction time of 1-20 seconds, a weight ratio of the agent to the oil (3-50): 1, and a weight ratio of the water to the oil (0.03-0.8) : 1, catalyst density is 120-290 kg / m 3 , gas linear velocity is 0.8-2.5 m / s, reaction pressure is 130-450 kPa, catalyst mass flow rate G s is 15-150 kg / (m 2 ⁇ s ).
  • the conditions for the second catalytic cracking reaction include: a reaction temperature of 550-620 ° C, a reaction time of 3-15 seconds, a weight ratio of the agent to the oil (10-30): 1, and a weight ratio of the water to the oil (0.05-0.5) : 1, catalyst density is 150-250 kg / m 3 , gas linear velocity is 1-1.8 m / s, and catalyst mass flow rate G s is 20-130 kg / (m 2 ⁇ s).
  • a catalytic cracking system comprising a dilute phase transport bed, a rapid fluidized bed, an oil separation device, a reaction product separation device and a regenerator;
  • the dilute phase transport bed is in fluid communication with the rapid fluidized bed and the dilute phase transport bed is located upstream of the rapid fluidized bed;
  • the dilute phase transport bed is provided with a catalyst inlet at the bottom and an inferior heavy oil inlet at the lower portion
  • the fast fluidized bed is provided with an oil agent outlet at the top
  • the oil agent separation device is provided with an oil agent inlet, a catalyst outlet, and a reaction product An outlet
  • the reaction product separation equipment is provided with a reaction product inlet, a dry gas outlet, a liquefied gas outlet, a cracked gasoline outlet, a cracked diesel outlet, and a cracked heavy oil outlet
  • the regenerator is provided with a catalyst inlet and a catalyst outlet
  • the catalyst inlet of the dilute phase transport bed is in fluid communication with the catalyst outlet of the regenerator
  • the oil agent outlet of the rapid fluidized bed is in fluid communication with the oil agent inlet of the oil agent separation device
  • the oil agent separation device An outlet of the reaction product is in fluid communication with a reaction product inlet of the reaction product separation equipment, and a catalyst outlet of the oil separation equipment is in fluid communication with a catalyst inlet of the regenerator.
  • Axial solid fraction ⁇ of the catalyst pressure difference between two axial points in the reaction zone measured by a differential pressure meter ⁇ distance between the two axial points ⁇ catalyst particle density;
  • the unit of pressure difference is kg / m 2
  • the unit of distance between two points in the axial direction is meter
  • the unit of catalyst particle density is kg / m 3 .
  • Catalyst particle density skeleton density / (catalyst pore volume ⁇ skeleton density +1), where the unit of skeleton density is kg / m3 and the unit of catalyst pore volume is m3 .
  • the skeleton density and catalyst pore volume are determined by the pycnometer method and Titration.
  • Reaction time volume of reaction zone / logarithmic average volume flow of oil and gas
  • the unit of the volume of the reaction zone is m 3
  • the unit of the logarithmic average volume flow of oil and gas is m 3 / s
  • Oil and gas logarithmic average volume flow (V out -V in ) / ln (V out / V in ), V out and V in are the oil and gas volume flow at the exit and entrance of the reaction zone, respectively;
  • m is the feed amount of raw oil and atomized steam per unit of time, in kilograms per second;
  • ⁇ 3 is the density of oil and gas at the exit of the reaction zone, in kg / m 3 ;
  • ⁇ 4 is the oil and gas at the entrance of the reaction zone Density in kg / m3 .
  • Catalyst density in the reaction zone pressure difference between two axial points in the reaction zone (or its upper, middle, and lower parts) measured by a differential pressure meter ⁇ the axis Distance between two points;
  • the unit of the pressure difference is kg / m 2.
  • the axial direction of the reaction zone is divided into three parts: upper, middle, and lower.
  • the unit of the distance between the two points in the axial direction is meters.
  • the linear velocity of the gas takes the logarithmic average of the linear velocity of the gas at the bottom of the reaction zone and the linear velocity of the gas at the top of the reaction zone.
  • the catalyst mass flow rate G of the average number of G s top and bottom end of the reaction zone takes s G s of the reaction zone;
  • the unit of catalyst circulation is kg / s;
  • Reactor catalyst circulation amount coke formation rate ⁇ (carbon content of catalyst to be regenerated-carbon content of regenerated catalyst);
  • the unit of coke generation speed is kg / s
  • the content of the catalyst carbon to be grown and the content of the regenerated catalyst carbon are both weight content
  • Coke generation speed flue gas quantity ⁇ (CO 2 % + CO%) ⁇ Vm ⁇ M;
  • Vm is the molar volume of the gas, and the value is 22.4 ⁇ 10 -3 m 3 / mol, and M is the molar mass of the carbon element, and the value is 12 ⁇ 10 -3 kg / mol;
  • Amount of flue gas (amount of regeneration air ⁇ 79% by volume) / (1-CO 2 % -CO% -O 2 %);
  • the unit of the regeneration air volume is m 3 / s
  • the unit of the flue gas volume is m 3 / s
  • CO 2 %, CO%, and O 2 % are the volume percentages of CO 2 , CO, and O 2 in the flue gas, respectively.
  • the feedstock oils used in the following examples and comparative examples are hydrogenation residues, and their properties are shown in Table 1.
  • the catalyst used was a commercial catalytic cracking catalyst purchased from the Catalyst Branch of China Petroleum and Chemical Corporation, under the brand name DMMC-2.
  • the test was carried out according to the flow chart shown in Figure 1.
  • the feedstock oil was hydrogenation residue.
  • the test was carried out on a medium-sized device using DMMC-2 catalyst.
  • the reactor was a combined reaction including a dilute phase transport bed and a rapid fluidized bed in series. Device.
  • the preheated feed oil enters the dilute phase transport bed and contacts the catalytic cracking catalyst to perform the first cracking reaction.
  • the reaction effluent enters the rapid fluidized bed from bottom to top and is mixed with the regenerated catalyst to continue the second catalytic cracking reaction.
  • the carbon content of the regenerated catalyst was 0.05% by weight, and the replenishment position of the catalyst was at 1/3 of the total height of the rapid fluidized bed.
  • the replenished catalyst accounted for 5% by weight of the reactor catalyst circulation.
  • the catalyst in the rapid fluidized bed is controlled to have a full-thickness distribution.
  • the axial solid fraction ⁇ of the catalyst in the rapid fluidized bed is from bottom to top. Within the range of 0.1-0.2.
  • the reaction product is quickly separated from the catalyst to be produced, and the reaction product is cut according to the distillation range in the product separation system.
  • the waiting catalyst enters the stripping section under the action of gravity, and the reaction products adsorbed on the waiting catalyst are lifted out by water vapor.
  • the stripped catalyst directly enters the regenerator without heat exchange, and is contacted with air for scorch regeneration and regeneration.
  • the catalyst is returned to the reactor for recycling.
  • Table 2 The operating conditions and product distribution used are listed in Table 2.
  • the ethylene yield of this example can reach 5.2% by weight
  • the propylene yield can reach 18.2% by weight
  • the light aromatics yield is 11.5% by weight
  • the dry gas and coke yields are 10.8% by weight and 8.5% by weight.
  • the test was carried out according to the flow shown in Figure 2.
  • the feedstock was hydrogenated residue and the DMMC-2 catalyst was used to conduct the test on a medium-sized device.
  • the reactor was divided into a combined type including a series of dilute phase transport beds and a rapid fluidized bed. reactor.
  • the preheated feedstock enters the bottom of the full-concentration reaction zone to contact the catalytic cracking catalyst and perform the first cracking reaction.
  • the reaction effluent enters the full-concentration reaction zone of the rapid fluidized bed from bottom to top and is mixed with the regenerated catalyst to continue.
  • the second catalytic cracking reaction The carbon content of the regenerated catalyst was 0.05% by weight, and the replenishment position of the catalyst was at 1/3 of the total height of the rapid fluidized bed.
  • the replenished catalyst accounted for 5% by weight of the reactor catalyst circulation.
  • the catalyst in the full-thickness reaction zone is controlled to have a full-thickness distribution, and the axial solid fraction of the catalyst in the full-thickness reaction zone ⁇ From bottom to top, it is in the range of 0.1-0.2.
  • the reaction product is quickly separated from the catalyst to be produced, and the reaction product is cut according to the distillation range in the product separation system.
  • the waiting catalyst enters the stripping section under the action of gravity, and the reaction products adsorbed on the waiting catalyst are lifted out by water vapor.
  • the stripped catalyst directly enters the regenerator without heat exchange, and is contacted with air for scorch regeneration and regeneration.
  • the catalyst is returned to the reactor for recycling.
  • the mixed C4 fraction obtained after cutting the reaction product is returned to the bottom of the dilute phase transport bed for further reaction.
  • the operating conditions and product distribution used are listed in Table 2.
  • the ethylene yield of this example can reach 5.9% by weight
  • the propylene yield can reach 21.1% by weight
  • the light aromatics yield is 11.8% by weight
  • the dry gas and coke yields are 10.7% by weight and 8.4% by weight.
  • the feedstock oil is hydrogenation residue oil, and DMMC-2 catalyst is used to test on a medium-sized device.
  • the reactor type is a combined reactor in which a riser and a fluidized bed are connected in series.
  • the preheated feed oil enters the lower part of the riser to be in contact with the catalyst for catalytic cracking reaction.
  • the reaction oil, gas and water vapor and the catalyst to be produced enter the dense-phase fluidized bed from the riser outlet to continue the reaction.
  • the stream enters a closed cyclone separator to quickly separate the reaction product and the catalyst to be produced, and the reaction product is cut according to the distillation range in the product separation system.
  • the waiting catalyst enters the stripping section under the action of gravity, and the reaction products adsorbed on the waiting catalyst are lifted out by water vapor.
  • the stripped catalyst directly enters the regenerator without heat exchange, and is contacted with air for scorch regeneration and regeneration. After the catalyst is returned to the riser for recycling.
  • the operating conditions and product distribution used are listed in Table 2.
  • the comparative example has an ethylene yield of 3.7% by weight, a propylene yield of 12.8% by weight, a light aromatics yield of 5.5% by weight, a dry gas and coke yield of 12.9% by weight and 13.3% by weight.
  • Comparative Example 2 is basically the same as Example 1, except that the catalyst distribution plate is not provided at the bottom of the rapid fluidized bed, and the axial solid fraction ⁇ of the catalyst in the rapid fluidized bed shows an increase of 0.1 ⁇ 0.2 ⁇ 0.3 from top to bottom.
  • the operating conditions used were the same as in Example 1.
  • the product distribution is shown in Table 2.
  • the method of the present application has higher yields of ethylene, propylene and light aromatics, and has lower yields of dry gas and coke.

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Abstract

一种烃油催化裂解方法和系统，该方法包括使烃油原料与催化裂解催化剂在包括串联的稀相输送床和快速流化床的反应器中接触反应的步骤，其中，在所述快速流化床中，所述催化剂的轴向固体分率ε控制在约0.1至约0.2的范围内。本申请的方法和系统用于烃油原料、特别是重质原料油的催化裂解时，干气和焦炭产率低，产品分布好。

Description

一种烃油催化裂解方法和系统
相关申请的交叉引用
本申请要求申请人于2018年7月16日向中国专利局提交的申请号为201810779819.0、名称为“一种采用稀相输送床与快速流化床进行催化裂解的方法和系统”的专利申请的优先权，上述专利申请的内容经此引用全文并入本文。
技术领域
本申请涉及催化裂化的技术领域，具体涉及一种烃油催化裂解方法和系统。
背景技术
以乙烯、丙烯为代表的低碳烯烃是化学工业的最基本原料，国内外多以天然气或轻质石油馏分为原料，采用乙烯联合装置中蒸汽裂解工艺生产低碳烯烃。苯、甲苯、二甲苯(BTX)是重要的基础化工原料，其中对二甲苯(PX)占BTX消费总量的45％左右。随着中国聚酯等工业的发展，预计BTX的需求将继续高速增长。约90％的乙烯、约70％的丙烯、90％的丁二烯、30％的芳烃均来自蒸汽裂解副产。虽然蒸汽裂解技术经过几十年的发展，技术不断完善，但仍具有能耗高、生产成本高、CO 2排放量大和产品结构不易调节等技术局限，石油化工如果采用传统的蒸汽裂解制乙烯、丙烯路线，将面临轻质原料油短缺、生产能力不足以及成本过高等几大制约因素，另外，随着蒸汽裂解原料的轻质化，丙烯和轻芳烃产率下降更是加剧的供需矛盾。催化裂解技术可以作为生产低碳烯烃和轻芳烃的生产工艺的有益补充，对炼油和化工一体化的企业来说，采用催化技术路线生产化工原料具有明显的社会与经济效益。
中国专利申请公开CN1234426A公开了一种从重质油同时制取低碳烯烃和高芳烃汽油的方法，包括使重质石油烃和水蒸汽在由提升管和密相流化床组成的复合反应器中进行催化裂解反应，达到提高低碳烯烃特别是丙烯的产率，同时使汽油中的芳烃含量增加到80重％左右。
中国专利申请公开CN1393510A公开了一种重质石油烃催化转化增产乙烯和丙烯的方法，包括使烃油原料，也是在提升管或流化床反 应器中与含有五元环高硅沸石的催化剂接触、反应，该方法不仅可以提高乙烯和丙烯产率，还可以在一定程度上缓解催化剂的水热失活。
中国专利申请公开CN1721510A公开了一种采用双反应区催化裂解生产低碳烯烃和芳烃的方法，其在两个反应区采用不同的重时空速达到从重质原料最大限度地生产丙烯、乙烯等低碳烯烃，其中丙烯的产率超过20重％，同时联产甲苯与二甲苯等芳烃的目的。
美国专利申请公开US2002003103A和US2002189973A公开了采用双提升管来增产丙烯的FCC装置，其中裂化反应生成的汽油(60～300°F/15-150℃)进料到第二根提升管进一步反应，催化剂是USY分子筛和ZSM-5分子筛催化剂的混合物。
美国专利申请公开US2002195373A和国际专利申请公开WO2017223310A公开的方法是采用下行式反应器，在高温(1020～1200°F/550-650℃)、短接触时间(＜0.5秒)和大剂油比(15～25)条件下操作。主催化剂(Y型八面沸石)具有低氢转移活性，其配方结合操作条件来制定，以使轻烯烃产率最大化。高效分离器在0.1秒之内将产物和催化剂分离，使得二次反应和焦炭的形成最小化。另外，使用LCO来骤冷分离出来的气体产物至约930°F/500℃，并防止进一步裂化。
美国专利US6538169A和美国专利申请公开US2003121825A公开的方法也是采用两个反应区和一个共用的再生器的构成反应-再生系统。在第一反应区，使用高温和高剂油比将重质原料裂化为轻烯烃或能转化为轻烯烃的中间产物。第二反应区由第二根提升管构成，在那儿操作条件更苛刻，从汽油产品中生产更多的轻组分。使用择形分子筛如ZSM-5辅助汽油向轻烯烃的转化，适用的原料包括VGO、HVGO和加氢的瓦斯油。
中国专利申请公开CN1403540A公开了一种制取乙烯和丙烯的催化转化方法，其中采用提升管和密相流化床串联反应器，将轻质原料注入提升管，在较高苛刻度下反应，反应产物与积炭催化剂进入流化床，在相对缓和条件下继续反应，该方法具有较高的乙烯+丙烯+丁烯收率。
中国专利申请公开CN102051213A公开了一种催化裂解方法，包括将重质原料在包括至少两个反应区的第一提升管反应器中与催化剂 接触进行裂化反应，将轻质原料和裂解重油在第二提升管反应器和流化床反应器中与催化剂接触进行裂化反应的步骤。该方法用于重油催化裂解，重油转化率和丙烯产率较高，干气和焦炭产率低。
炼油化工行业结构性矛盾日益严重，一方面传统石化产品产能过剩，成品油供需矛盾突出，另一方面资源类产品和高端石油化工产品短缺突出，炼油向化工转型已是大势所趋。作为炼油与化工桥梁的催化裂解装置面临前所未有压力与挑战。当前，催化裂解装置掺炼常压渣油比例越来越大，甚至提出掺炼减压渣油的要求，现有催化裂解技术通常以减压蜡油或石蜡基常压渣油为原料，最先进的催化裂解技术是采用双提升管或者提升管串联密相床层的反应器，在较高苛刻度的反应条件下，达到多产低碳烯烃和/或轻芳烃的目标，这类反应器在加工掺渣重油时，不可避免地出现干气和焦炭产率高的问题。采用下行式反应器时可以达到焦炭产率的降低，但是反应转化率相对较低，需要专用催化剂。随着原料的重质化，催化裂解装置掺炼渣油的要求越来越多，为了高效利用劣质重油资源，满足日益增长的化工原料如低碳烯烃和芳烃的需求，有必要开发一种将劣质重油原料转化为高附加值产品的催化裂解方法。
发明内容
本申请的目的是提供一种新型的烃油催化裂解方法和系统，其特别适用于由烃油原料、特别是重质原料油经催化裂解来生产低碳烯烃，例如乙烯和丙烯。采用本申请的方法和系统进行催化裂解的干气和焦炭产率低，产品分布好。
为了实现上述目的，在一个方面，本申请提供了一种烃油催化裂解方法，包括使烃油原料、特别是重质原料油与催化裂解催化剂在包括串联的稀相输送床和快速流化床的反应器中接触反应的步骤，其中，在所述快速流化床中，所述催化剂的轴向固体分率ε控制在约0.1至约0.2的范围内。
另一方面，本申请提供了一种适用于烃油、特别是重质原料油的催化裂解的系统，该系统包括催化裂解反应器、油剂分离设备、任选的反应产物分离设备、和再生器，
所述催化裂解反应器包括串联的稀相输送床和快速流化床，按照反应物料的流向，所述稀相输送床与快速流化床流体流通且所述稀相 输送床位于所述快速流化床的上游；
所述稀相输送床设置有底部的催化剂入口、下部的第一反应原料入口，所述快速流化床设置有顶部的油剂出口和任选的底部的第二反应原料入口，所述油剂分离设备设置有油剂入口、催化剂出口和反应产物出口，所述任选的反应产物分离设备设置有反应产物入口、干气出口、液化气出口、裂解汽油出口、裂解柴油出口和裂解重油出口，所述再生器设置有催化剂入口和催化剂出口，
所述稀相输送床的催化剂入口与所述再生器的催化剂出口流体连通，所述快速流化床的油剂出口与所述油剂分离设备的油剂入口流体连通，所述油剂分离设备的反应产物出口与所述任选的反应产物分离设备的反应产物入口流体连通，所述油剂分离设备的催化剂出口与所述再生器的催化剂入口流体连通。
本申请通过将快速流化床的催化剂轴向固体分率ε控制在约0.1至约0.2的范围内可以有效提高快速流化床内的催化剂密度，从而大幅度提高反应瞬间的催化剂和烃油原料之比，并控制相对较长的反应时间，使催化剂能够与烃油原料、特别是劣质重油进行充分反应。由此不仅能提高反应转化率，还能提高低碳烯烃和轻芳烃的产率，同时，还可以有效减少干气和焦炭的生成，使产品分布与产品质量得到改善。
进一步地，本申请通过在快速流化床中引入补充催化剂可以在较大范围内调节剂油比，为裂解反应提供更多的活性中心。同时，引入补充催化剂增强了反应温度调节的灵活性，可以有效调节快速流化床内温度与催化剂活性的梯度。
采用本申请的方法和系统可以使石化企业从廉价的劣质重油最大限度地生产高附加值的化工原料，有助于推进炼油企业的炼化一体化进程，既解决了石化原料短缺的问题，又提高了石化行业的经济效益和社会效益。
本申请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申请的进一步理解，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申请，但并不构成对本发 明的限制。在附图中：
图1为本申请的一种优选实施方式的示意图；以及
图2为本申请的另一优选实施方式的示意图。
附图标记说明
I稀相输送床        II快速流化床
1预提升介质管线    2预提升段          3出口段
4沉降器            5汽提段            6旋风分离器
7集气室            8油气管线          9待生斜管
10再生器           11再生斜管         12烟气管线
13空气分配器       14第一进料管线     15补剂管线
16第二进料管线     17全浓相反应区     18过渡段
具体实施方式
以下将通过具体的实施方式对本申请作出进一步的详细描述，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申请，但不以任何方式限制本发明。
在本文中所披露的任何具体数值(包括数值范围的端点)都不限于该数值的精确值，而应当理解为还涵盖了接近该精确值的值，例如在该精确值±5％范围内的所有可能的数值。并且，对于所披露的数值范围而言，在该范围的端点值之间、端点值与范围内的具体点值之间，以及各具体点值之间可以任意组合而得到一个或多个新的数值范围，这些新的数值范围也应被视为在本文中具体公开。
在本申请中，所述稀相输送床和快速流化床构成所述反应器的两个串联的反应区，因此也可以分别称为稀相输送床反应区和快速流化床反应区。
在本申请中，术语“稀相输送床”具有本领域技术人员所熟知的含义，具体是指其中的催化剂颗粒在流体中形成悬浮状态的稀相，并被流体从流化床中一起夹带出去的流化床形式。
在本申请中，术语“快速流化床”具有本领域技术人员所熟知的含义，具体是指其中的催化剂颗粒处于快速流态化状态的流化床形式，其中快速流态化是一种无气泡的气固接触流态化，重要特征是固体颗粒趋向于成团运动。当催化剂处于快速流态化状态时，流化床内 催化剂的轴向固体分率ε通常在约0.05至约0.4的范围内。但是，在常规快速流化床中，催化剂通常呈上稀下浓式分布，例如上部的催化剂轴向固体分率ε可在约0.05至约0.1的范围内，而下部的催化剂轴向固体分率ε可在约0.3至约0.4的范围内。
根据本申请，在所述快速流化床中，当催化剂的轴向固体分率ε自下至上始终控制在约0.1至约0.2的范围内(即在该反应区沿轴向均分的上、中、下三部分中测得的催化剂轴向固体分率ε均大于等于约0.1且小于等于约0.2)时，整个所述快速流化床内的催化剂呈拟均一的全浓相分布。相应地，其内的催化剂呈此类全浓相分布的快速流化反应区可以称为“全浓相反应区”。
在本申请中，术语“水油重量比”是指注入反应器内的总蒸汽重量与原料重量之比。
在本申请中，所谓“上游”和“下游”均是基于反应物料的流动方向而言的。例如，当反应物流自下而上流动时，“上游”表示位于下方的位置，而“下游”表示位于上方的位置。
除非另有说明，本文所用的术语具有与本领域技术人员通常所理解的相同的含义，如果术语在本文中有定义，且其定义与本领域的通常理解不同，则以本文的定义为准。
本申请中，除了明确说明的内容之外，未提到的任何事宜或事项均直接适用本领域已知的那些而无需进行任何改变。而且，本文描述的任何实施方式均可以与本文描述的一种或多种其他实施方式自由结合，由此形成的技术方案或技术思想均视为本发明原始公开或原始记载的一部分，而不应被视为是本文未曾披露或预期过的新内容，除非本领域技术人员认为该结合明显不合理。
在本文中提及的所有专利和非专利文献，包括但不限于教科书和期刊文章等，均通过引用方式全文并入本文。
在第一方面，本申请提供了一种烃油催化裂解方法，包括使烃油原料、特别是重质原料油与催化裂解催化剂在包括串联的稀相输送床和快速流化床的反应器中接触反应的步骤，其中，在所述快速流化床中，所述催化剂的轴向固体分率ε控制在约0.1至约0.2的范围内。
在优选的实施方式中，根据本申请的方法用于由重质原料油生产低碳烯烃，例如乙烯和丙烯，并且进一步包括如下步骤：
i)使包括轻质原料和/或重质原料油的第一反应原料与催化裂解催化剂在稀相输送床中接触进行第一催化裂解反应；以及
ii)使步骤i)的反应流出物和任选的包括轻质原料和/或重质原料油的第二反应原料在快速流化床中进行第二催化裂解反应，
所述轻质原料选自C4烃馏分、C5-C6轻汽油馏分和它们的任意组合，且所述第一和第二反应原料中至少一者包括所述重质原料油；
其中，在所述快速流化床中，所述催化剂的轴向固体分率ε控制在约0.1至约0.2的范围内。
根据本申请，通过将所述快速流化床中的催化剂轴向固体分率ε自下至上始终控制在约0.1至约0.2范围内，防止了催化剂在该快速流化床内呈上稀下浓式分布，使所述快速流化床上下的实际剂油比保持一致，进而能够减少干气和焦炭的产率，并提高目标产物的产率。
在进一步优选的实施方式中，本申请方法进一步包括如下步骤：
i)使所述第一反应原料与催化裂解催化剂在稀相输送床中接触进行第一催化裂解反应；
ii)使步骤i)的反应流出物和任选的第二反应原料在快速流化床中、在有效生成低碳烯烃的条件下进行第二催化裂解反应，其中在所述快速流化床中所述催化剂的轴向固体分率ε控制在约0.1至约0.2的范围内；
iii)将来自所述反应器的反应流出物分离，得到富含低碳烯烃的反应产物和待生催化剂；
iv)使所述待生催化剂再生，并将所得再生催化剂的至少一部分返回步骤i)作为所述催化裂解催化剂；以及
v)任选地，将所述反应产物分离得到干气、液化气、裂解汽油、裂解柴油和裂解重油。
根据本申请，所谓“有效生成低碳烯烃”是指至少部分反应原料在所述快速流化床中发生有效裂化、例如深度裂化，产生低碳烯烃产物如乙烯和丙烯，从而使所得的产物混合物富含低碳烯烃。
根据本申请，所谓“富含低碳烯烃”是指所述反应产物或产物混合物中的低碳烯烃(如乙烯和丙烯)的总含量高于所述反应产物或产物混合物的约10％重量，优选高于约15％重量，更优选高于约20％重量。
在某些具体实施方式中，在所述步骤i)之前和/或所述步骤ii)之后，本申请方法可以进一步包括一个或多个额外的反应步骤，例如在额外的流化床反应区如稀相输送床、密相流化床、常规快速流化床等中进行的催化裂化和/或催化异构化等反应步骤。
在另一些具体实施方式中，本申请方法在所述步骤i)之前和所述步骤ii)之后不包括额外的反应步骤。
在某些特别优选的实施方式中，本申请方法进一步包括如下步骤：
i)将包括经预热的劣质重油的第一反应原料从稀相输送床的下部引入稀相输送床中与催化裂解催化剂接触，并在由下至上通过所述稀相输送床的过程中进行第一催化裂解反应，得到包含第一反应产物和半待生催化剂的反应流出物；
ii)将步骤i)的反应流出物引入快速流化床的底部，并与任选的自快速流化床底部引入的包括轻质原料的第二反应原料一起，在由下至上通过所述快速流化床的过程中进行第二催化裂解反应，得到包含第二反应产物和待生催化剂的反应流出物，其中，在所述快速流化床中催化剂的轴向固体分率ε控制为满足：0.1≤ε≤0.2；
iii)将步骤ii)的反应流出物中的第二反应产物与待生催化剂分离；
iv)将所述待生催化剂送入再生器进行烧焦再生，并将所得再生催化剂的至少一部分返回步骤i)作为所述催化裂解催化剂；以及
v)任选地，将所得的第二反应产物分离得到干气、液化气、裂解汽油、裂解柴油和裂解重油。
在另一些特别优选的实施方式中，本申请方法进一步包括如下步骤：
i)将包括轻质原料的第一反应原料从稀相输送床的下部引入稀相输送床中与催化裂解催化剂接触，并在由下至上通过所述稀相输送床的过程中进行第一催化裂解反应，得到包含第一反应产物和半待生催化剂的反应流出物；
ii)将步骤i)的反应流出物引入快速流化床的底部，并与自快速流化床底部引入的包括经预热的劣质重油的第二反应原料一起，在由下至上通过所述快速流化床的过程中进行第二催化裂解反应，得到包含 第二反应产物和待生催化剂的反应流出物，其中，在所述快速流化床中催化剂的轴向固体分率ε控制为满足：0.1≤ε≤0.2；
iii)将步骤ii)的反应流出物中的第二反应产物与待生催化剂分离；
iv)将所述待生催化剂送入再生器进行烧焦再生，并将所得再生催化剂的至少一部分返回步骤i)作为所述催化裂解催化剂；以及
v)任选地，将所得的第二反应产物分离得到干气、液化气、裂解汽油、裂解柴油和裂解重油。
在优选的实施方式中，本申请的方法进一步包括：在所述快速流化床中引入一股或多股补充催化剂，并使其与所述快速流化床内的物料接触进行催化裂解反应。
根据本申请，所述一股或多股补充催化剂的炭含量可以各自独立地为约0-1.0重量％，例如所述一股或多股补充催化剂可以各自独立地选自再生催化剂、待生催化剂和半再生催化剂，即再生、待生和半再生的催化裂解催化剂。
根据本申请，所述一股或多股补充催化剂的总量可以占反应器催化剂循环量的约0-50重量％，优选约5-30重量％。
根据本申请，所述一股或多股补充催化剂的引入位置距所述快速流化床底部的距离各自独立地为所述快速流化床总高度的约0％至约90％。优选地，所述一股或多股补充催化剂的引入位置各自独立地位于所述快速流化床的约20％到约80％高度处，更优选约30％到约75％高度处。例如，所述引入位置可以在所述快速流化床的底部，或者在所述快速流化床总高度的约1/3处。所述补充催化剂的温度可以根据所需的反应温度进行调整，例如可以引入冷和/或热的再生催化剂，也可以引入冷和/或热的待生催化剂。
根据本申请，在所述快速流化床中引入补充催化剂可以在较大范围内调节剂油比，为裂解反应提供更多的活性中心。同时，引入补充催化剂增强了反应温度调节的灵活性，可以有效调节快速流化床内的温度和催化剂活性的梯度。另外，在快速流化床中引入补充催化剂可以尽可能地维持该流化床内催化剂密度的均匀性，有效调节催化剂密度的分布，保证裂解反应充分地、有效地进行，从而提高目标产物的选择性。
根据本申请，可以通过调节快速流化床中的气体线速，和/或在快速流化床中设置催化剂分布板来进一步调节快速流化床内的催化剂分布，由此使所述催化剂呈拟均一的全浓相分布。
根据本申请，所述烃油原料，如重质原料油、特别是劣质重油，可以在一个或多个位置引入所述包括稀相输送床和快速流化床的反应器中。例如，所述烃油原料可以在一个进料位置全部引入所述稀相输送床中，或者在一个进料位置全部引入所述快速流化床中。任选地，所述烃油原料可以从两个或两个以上的进料位置按照相同或不同的比例引入所述稀相输送床和/或快速流化床中，例如一部分烃油原料从一个进料位置引入所述稀相输送床中，而另一部分烃油原料从另一进料位置引入所述快速流化床中，或者所述烃油原料从两个或两个以上的进料位置引入所述稀相输送床中，或者所述烃油原料从两个或两个以上的进料位置引入所述快速流化床中。
根据本申请，劣质重油是指比常规重油更加不适宜催化裂解加工的重油。例如，所述劣质重油的性质可以满足以下指标中的至少一种，例如一种、两种、三种或四种：20℃的密度为约900-1000千克/米 3，优选为约910-940千克/米 3；残炭为约2-10重量％，优选为约3-8重量％；镍和钒的总含量为约2-30ppm，优选为约5-20ppm；特性因数K值小于约12.1，优选为小于约12.0。劣质重油中的残炭采用ASTMD-189康氏残炭实验方法进行测定。
作为示例，所述劣质重油可以为重质石油烃和/或其它矿物油；所述重质石油烃可以选自减压渣油(VR)、劣质的常压渣油(AR)、劣质的加氢渣油、焦化瓦斯油、脱沥青油、减压蜡油、高酸值原油、高金属原油和它们的任意组合；所述其它矿物油可以选自煤液化油、油砂油、页岩油和它们的任意组合。
本申请对所用的催化裂解催化剂没有严格的限制，例如可以是本领域技术人员熟知的适用于由烃油原料、如重质原料油生产低碳烯烃的各种催化裂解催化剂。在优选的实施方式中，以干基计并以催化裂解催化剂的干基重量为基准，所述催化裂解催化剂包括约1-50重量％，优选约5-45重量％，更优选约10-40重量％的沸石、约5-99重量％，优选约10-80重量％，更优选约20-70重量％的无机氧化物，和约0-70重量％，优选约5-60重量％，更优选约10-50重量％的粘土； 所述沸石作为活性组分，可以包括中孔沸石和任选的大孔沸石。优选地，以干基计，所述中孔沸石占沸石总重量的约0-50重量％，优选占约0-20重量％。
在本申请中，所述中孔沸石和大孔沸石沿用本领域的常规定义，即中孔沸石的平均孔径为约0.5-0.6nm，大孔沸石的平均孔径为约0.7-1.0nm。
作为示例，所述大孔沸石可以选自由稀土Y(REY)型沸石、稀土氢Y(REHY)型沸石、由不同方法得到的超稳Y型沸石和高硅Y型沸石中的一种或多种。所述中孔沸石可以选自具有MFI结构的沸石，例如ZSM系列沸石和/或ZRP沸石。任选地，还可对上述中孔沸石用磷等非金属元素和/或铁、钴、镍等过渡金属元素进行改性。有关ZRP沸石的更为详尽的描述可参见美国专利US5,232,675A。ZSM系列沸石优选选自ZSM-5、ZSM-11、ZSM-12、ZSM-23、ZSM-35、ZSM-38、ZSM-48和其它类似结构的沸石之中的一种或更多种的混合物。有关ZSM-5的更为详尽的描述可参见美国专利US3,702,886A。
根据本申请，所述无机氧化物作为粘结剂，优选为二氧化硅(SiO 2)和/或三氧化二铝(Al 2O 3)。所述粘土作为基质(即载体)，优选为高岭土和/或多水高岭土。
本申请对所用的催化裂解反应条件没有严格的限制，例如可以是本领域技术人员熟知的适用于由烃油原料、如重质原料油生产低碳烯烃的催化裂解反应条件。在优选的实施方式中，所述第一催化裂解反应的条件可以包括：反应温度为约500-600℃，反应时间为约0.05-5秒，剂油重量比为约1∶1至约50∶1，水油重量比为约0.03∶1至约0.5∶1，催化剂密度为约20-100千克/米 3，气体线速为约4-18米/秒，反应压力为约130-450千帕，催化剂质量流率G s为约180-500千克/(米 2·秒)。进一步优选地，所述第一催化裂解反应的条件包括：反应温度为约520-580℃，反应时间为约1-3秒，剂油重量比为约5∶1至约25∶1，水油重量比为约0.05∶1至约0.3∶1。
在优选的实施方式中，所述第二催化裂解反应的条件可以包括：反应温度为约510-650℃，反应时间为约1-20秒，剂油重量比为约3∶1至约50∶1，水油重量比为约0.03∶1至约0.8∶1，催化剂密度为约120-290千克/米 3，气体线速为约0.8-2.5米/秒，反应压力为约130-450 千帕，催化剂质量流率G s为约15-150千克/(米 2·秒)。进一步优选地，所述第二催化裂解反应的条件包括：反应温度为约550-620℃，反应时间为约3-15秒，剂油重量比为约10∶1至约30∶1，水油重量比为约0.05∶1至约0.5∶1，催化剂密度为约150-250千克/米 3，气体线速为约1-1.8米/秒，反应压力为约130-450千帕，催化剂质量流率G s为约20-130千克/(米 2·秒)。
根据本申请，反应产物与待生催化剂的分离可以采用本领域技术人员所熟知的方式进行，例如可以在沉降器中采用旋风分离器进行。所述反应产物进一步分离得到干气、液化气、裂解汽油、裂解柴油和裂解重油的方式也是本领域技术人员所熟知的。在优选的实施方式中，所述干气和液化气可以进一步采用本领域常规的分离手段分离得到乙烯、丙烯等目的产物等。
在优选的实施方式中，根据本申请的方法进一步包括：将C4烃馏分和/或C5-C6轻汽油馏分作为所述第一和/或第二反应原料中的轻质原料在一个或多个位置引入所述快速流化床和/或稀相输送床中进行催化裂解反应。例如，在某些优选实施方式中，所述第一反应原料和第二反应原料中至少一者包括选自C4烃馏分、C5-C6轻汽油馏分和它们的任意组合的轻质原料。在某些进一步优选的实施方式中，所述第一反应原料包括轻质原料和重质原料油，并且所述轻质原料的至少一部分在所述重质原料油引入所述稀相输送床的位置的上游引入所述稀相输送床中。在另一些进一步优选的实施方式中，所述第一反应原料包括重质原料油，例如劣质重油，而所述第二反应原料包括所述轻质原料。在另一些进一步优选的实施方式中，所述第一反应原料包括所述轻质原料，而所述第二反应原料包括重质原料油，例如劣质重油。
根据本申请，所述“C4烃馏分”是指以C4馏分为主要成分的常温、常压下以气体形式存在的低分子碳氢化合物，包括分子中碳原子数为4的烷烃、烯烃及炔烃。它既可以包括本申请方法所产的富含C4烃馏分的气态烃产品(例如液化气)，也可以包括其它装置所产的富含C4馏分的气态烃，优选本申请方法所产的C4烃馏分。所述C4烃馏分优选为富含烯烃的C4烃馏分，其中C4烯烃的含量可以为大于约50重量％，优选大于约60重量％，更优选在约70重量％以上。
根据本申请，所述“C5-C6轻汽油馏分”是指汽油中碳数为C5至 C6的组分，其可以包括本申请方法所产的裂解汽油，也可以包括其它装置所产的汽油馏分，例如可以为选自催化裂解汽油、催化裂化汽油、直馏汽油、焦化汽油、热裂解汽油、热裂化汽油和加氢汽油中的至少一种的C5-C6馏分。
根据本申请，待生催化剂可以通过本领域技术人员所熟知的方式再生，例如可以在再生器中烧焦再生，具体地可以将诸如空气等的含氧气体引入再生器中与待生催化剂接触。烧焦再生所得的烟气可以在再生器中与催化剂分离后，送入后续能量回收系统。
在本申请的某些优选实施方式中，可以将经再生器烧焦再生后的再生催化剂通过催化剂冷却器降温至约600-680℃后再返回所述反应器。热的再生催化剂经冷却后返回反应器，有助于降低油剂接触温度，改善原料油和催化剂的接触状态，进而改善对干气和生焦的选择性。
在某些优选的实施方式中，所述快速流化床从下到上依次包括全浓相反应区和过渡段，所述全浓相反应区为横截面呈大致圆形的、底端和顶端开口的、等直径或者变直径的空心柱体的型式，所述稀相输送床与所述全浓相反应区的底端相连通，所述全浓相反应区的顶端经由所述过渡段与所述反应器的出口段相连通，所述全浓相反应区的底部任选设有一个或多个供所述第二反应原料进料的入口，
其中，所述全浓相反应区的底端的横截面直径大于或等于所述稀相输送床的直径，且顶端的横截面直径大于所述出口段的直径，并且所述全浓相反应区的底部或侧壁设有一个或多个补充催化剂入口，所述一个或多个补充催化剂入口的位置各自独立地位于所述快速流化床总高度的约0％至约90％的高度处，优选约20％到约80％高度处，更优选约30％到约75％高度处。
根据本申请，所述全浓相反应区可以为各种横截面呈大致圆形的、底端和顶端开口的、等直径或者变直径的空心柱体的型式，例如等直径的空心圆柱体，或者由下至上直径连续地或不连续地增大的空心柱体的型式。
根据本申请，所谓“直径连续地增大”是指直径以线性或者非线性的方式连续不断地增大。作为“由下至上直径连续增大的空心柱体”的例子，可以举出倒置的空心截头圆锥体。
根据本申请，所谓“直径不连续地增大”是指直径以不连续，例如阶梯式，的方式增大。作为“由下至上直径不连续地增大的空心柱体”的例子，可以举出由两段或更多段直径递增的圆柱体构成的空心柱体。
作为示例，所述全浓相反应区可以为空心圆柱体、倒置的空心截头圆锥体、由两段或更多段直径递增的圆柱体构成的空心柱体、由两段或更多段直径递增的倒置的截头圆锥体构成的空心柱体、或者由一段或多段圆柱体与一段或多段倒置的截头圆锥体构成的空心柱体的型式。
在某些优选的实施方式中，所述全浓相反应区的底部设有催化剂分布板。
在某些优选的实施方式中，所述全浓相反应区的底部设有一个或多个第二反应原料入口，且优选地，所述第二反应原料入口的位置设有气体分布器。
在优选的实施方式中，所述全浓相反应区的最大横截面的直径与所述快速反应床的总高度之比为约0.005∶1至约1∶1，优选约0.01∶1至约0.8∶1，更优选约0.05∶1至约0.5∶1；全浓相反应区的高度与所述快速反应床的总高度之比为约0.1∶1至约0.9∶1，优选约0.3∶1至约0.85∶1，更优选约0.5∶1至约08∶1。
在某些优选实施方式中，所述全浓相反应区为倒置的空心截头圆锥体的型式，纵切面为等腰梯形，其底端横截面的直径为约0.2-10米，优选约0.5-8米，更优选约1-5米；顶端横截面直径与底端横截面直径比的比值为大于1至约50，优选约1.2至约10，更优选约1.5至约5；最大横截面的直径与所述快速反应床的总高度之比为约0.005∶1至约1∶1，优选约0.01∶1至约0.8∶1，更优选约0.05∶1至约0.5∶1；全浓相反应区的高度与所述快速反应床的总高度之比为约0.1∶1至约0.9∶1，优选约0.3∶1至约0.85∶1，更优选约0.5∶1至约0.8∶1。
在另一些优选的实施方式中，所述全浓相反应区为由一段倒置的截头圆锥体与一段圆柱体构成的空心柱体的型式，优选截头圆锥体位于圆柱体的下方，其中所述截头圆锥体的纵切面为等腰梯形，其底端横截面的直径为约0.2-10米，优选约0.5-8米，更优选约1-5米；顶端 横截面直径与底端横截面直径的比值为大于1至约50，优选约1.2至约10，更优选约1.5至约5；所述圆柱体的直径与所述截头圆锥体的顶端横截面的直径大致相同，并且所述空心圆柱体的高度与所述截头圆锥体的高度之比为约0.4∶1至2.5∶1，优选约0.8∶1至约1.5∶1；所述全浓相反应区的最大横截面的直径与所述快速反应床的总高度之比为约0.005∶1至约1∶1，优选约0.01∶1至约0.8∶1，更优选约0.05∶1至约0.5∶1；全浓相反应区的高度与所述快速反应床的总高度之比为约0.1∶1至约0.9∶1，优选约0.3∶1至约0.85∶1，更优选约0.5∶1至约0.8∶1。
在又一些优选的实施方式中，所述全浓相反应区为空心的圆柱体的型式，其直径为约0.2-10米，优选约1-5米；所述全浓相反应区的直径与所述快速反应床的总高度之比为约0.005∶1至约1∶1，优选约0.01∶1至约0.8∶1，更优选约0.05∶1至约0.5∶1；全浓相反应区的高度与所述快速反应床的总高度之比为约0.1∶1至约0.9∶1，优选约0.3∶1至约0.85∶1，更优选约0.5∶1至约0.8∶1。
在优选的实施方式中，所述全浓相反应区的高度为约2-50米，优选约5-40米，更优选约8-20米。
在优选的实施方式中，所述过渡段的高度与所述快速反应床的总高度之比为约01∶1至约0.9∶1，优选约0.2∶1至约0.5∶1。进一步优选地，该过渡段为空心的截头圆锥体的型式，纵切面为等腰梯形，等腰梯形侧边的内倾角α为约25-85°，优选约30-75°。
在某些具体实施方式中，本申请采用的反应器可以在所述稀相输送床的上游和/或所述快速流化床的下游进一步包括一个或多个额外的流化床反应区，例如稀相输送床、密相流化床、常规快速流化床等。
在另一些具体实施方式中，本申请采用的反应器在所述稀相输送床的上游和所述快速流化床的下游不包括额外的反应区。
在某些具体实施方式中，本申请采用的反应器可以与沉降器同轴布置，也可以与沉降器高低并列布置。
在第二方面，本申请提供了一种适用于烃油、特别是重质原料油的催化裂解的系统，该系统包括催化裂解反应器、油剂分离设备、任选的反应产物分离设备、和再生器，
所述催化裂解反应器包括串联的稀相输送床和快速流化床，按照 反应物料的流向，所述稀相输送床与快速流化床流体流通且所述稀相输送床位于所述快速流化床的上游；
所述稀相输送床设置有底部的催化剂入口、下部的第一反应原料入口，所述快速流化床设置有顶部的油剂出口和任选的底部的第二反应原料入口，所述油剂分离设备设置有油剂入口、催化剂出口和反应产物出口，所述任选的反应产物分离设备设置有反应产物入口、干气出口、液化气出口、裂解汽油出口、裂解柴油出口和裂解重油出口，所述再生器设置有催化剂入口和催化剂出口，
所述稀相输送床的催化剂入口与所述再生器的催化剂出口流体连通，所述快速流化床的油剂出口与所述油剂分离设备的油剂入口流体连通，所述油剂分离设备的反应产物出口与所述任选的反应产物分离设备的反应产物入口流体连通，所述油剂分离设备的催化剂出口与所述再生器的催化剂入口流体连通。
在优选的实施方式中，所述快速流化床与稀相输送床上下同轴设置，且快速流化床位于稀相输送床的上方。
在某些优选的实施方式中，所述快速流化床中设有催化剂分布板，其可以设置在所述快速流化床的底部，例如在所述稀相输送床与快速流化床的连接处。
根据本申请，所述催化剂分布板可以为工业上常见的各种型式的分布板，例如为平板形、拱形、碟形、环形和伞形中的一种或多种。采用催化剂分布板有助于使催化剂在全浓相反应区的轴向上浓度均一地与原料油接触进行催化裂解反应，从而减少催化剂浓度过高或过低带来的剂油比焦和热反应焦的生成。
在某些优选的实施方式中，所述快速流化床底部设有一个或多个第二反应原料入口，优选地在所述入口的位置设有气体分布器。
在某些优选的实施方式中，所述快速流化床具有如上文所述的结构，即从下到上依次包括全浓相反应区和过渡段，所述全浓相反应区和过渡段的具体设置在此不再详细描述。
在某些具体实施方式中，所述催化裂解反应器可以在所述稀相输送床的上游和/或所述快速流化床的下游进一步包括一个或多个额外的流化床反应区，例如稀相输送床、密相流化床、常规快速流化床等。
在另一些具体实施方式中，所述催化裂解反应器在所述稀相输送 床的上游和所述快速流化床的下游不包括额外的反应区。
根据本申请，所述油剂分离设备和反应产物分离设备均可采用本领域技术人员所熟知的设备。例如，所述油剂分离设备可以包括旋风分离器、沉降器和汽提器等，而所述反应产物分离设备可以是分馏塔等。
下面将结合附图所示的优选实施方式来进一步说明本申请，但是并不因此而限制本申请。
图1示出了本申请的一种优选实施方式，其中预提升介质经预提升介质管线1由预提升段2底部进入稀相输送床I的底部，所述预提升介质可以为干气、水蒸气或它们的混合物。来自再生斜管11的再生催化剂进入预提升段2的下部，在预提升介质的提升作用下进入稀相输送床I并向上运动。包括C4烃馏分、C5-C6轻汽油馏分和/或烃油原料、如劣质重油的第一反应原料经第一进料管线14注入稀相输送床I的下部，与稀相输送床I内已有的物流混合接触并进行第一催化裂解反应，得到包含第一反应产物和半待生催化剂的反应流出物。该反应流出物向上运动，进入快速流化床II的底部，与经由补剂管线15引入的补充催化剂，其可为再生催化剂或待生催化剂，接触并进行第二催化裂解反应。补剂管线15连接在快速流化床II的约0％至约90％高度处，优选约20％到约80％高度处，更优选约30％到约75％高度处。任选地，经由第二进料管线16向快速流化床II的底部通入包括C4烃馏分、C5-C6轻汽油馏分和/或烃油原料、如劣质重油的第二反应原料。反应产生的包含第二反应产物和失活的待生催化剂的反应流出物经出口段3进入沉降器4中的旋风分离器6，实现待生催化剂与第二反应产物的分离。分离出的第二反应产物进入集气室7，集气室7中的反应产物经过油气管线8进入后续的产物分离系统(图中未示出)。催化剂细粉由旋风分离器6的料腿返回沉降器4，沉降器4中的待生催化剂流向汽提段5。从待生催化剂中汽提出的反应产物经旋风分离器6后进入集气室7。汽提后的待生催化剂经待生斜管9进入再生器10，空气经空气分配器13分配后进入再生器10，烧去位于再生器10底部的密相床层中的待生催化剂上的焦炭，使失活的待生催化剂再生，得到再生催化剂。再生催化剂经再生斜管11返回预提升段2，烟气经烟气管线12进入后续的能量回收系统(图中未示出)。
图2示出了本申请的另一优选实施方式，其与图1所示的实施方式基本相同，区别仅在于所述快速流化床II从下到上依次包括全浓相反应区17和过渡段18。全浓相反应区17为倒置的空心截头圆锥体的型式，纵切面为等腰梯形。过渡段18为空心截头圆锥体型式，纵切面为等腰梯形，等腰梯形侧边的内倾角α为约25-85°，优选约30-75°。
在某些优选的实施方式中，本申请提供了以下的技术方案：
1、一种采用稀相输送床与快速流化床进行催化裂解的方法，该方法包括：
i)将预热的劣质重油从稀相输送床的下部引入稀相输送床中与催化裂解催化剂接触并由下至上进行第一催化裂解反应，得到第一反应产物和半待生催化剂；
ii)将所得第一反应产物和半待生催化剂引入快速流化床的底部并由下至上进行第二催化裂解反应，得到第二反应产物和待生催化剂；其中，所述快速流化床中催化剂呈全浓相分布，所述快速流化床中轴向固体分率ε分布满足：0.1≤ε≤0.2；
iii)将待生催化剂送入再生器进行烧焦再生，至少将部分所得再生催化剂作为所述催化裂解催化剂返回稀相输送床的底部；以及
iv)将所得第二反应产物进行分离，得到干气、液化气、裂解汽油、裂解柴油和裂解重油。
2、根据项目1所述的方法，其中，所述劣质重油的性质满足以下指标中的一种、两种、三种或四种：20℃密度为900-1000千克/米 3，残炭为2-10重量％，镍和钒总含量为2-30ppm，特性因数K值小于12.1。
3、根据项目1所述的方法，其中，所述劣质重油的性质满足以下指标中的一种、两种、三种或四种：20℃密度为910-940千克/米 3，残炭为3-8重量％，镍和钒总含量为5-20ppm，特性因数K值小于12.0。
4、根据项目1所述的方法，其中，所述劣质重油为重质石油烃和/或其它矿物油；
所述重质石油烃为选自减压渣油、劣质的常压渣油、劣质的加氢渣油、焦化瓦斯油、脱沥青油、减压蜡油、高酸值原油和高金属原油中的一种或多种，所述其它矿物油为选自煤液化油、油砂油和页岩油中的一种或多种。
5、根据项目1所述的方法，其中，以干基计并以催化裂解催化剂的干基重量为基准，所述催化裂解催化剂包括1-50重量％的沸石、5-99重量％的无机氧化物和0-70重量％的粘土；
所述沸石包括中孔沸石和任选的大孔沸石，所述中孔沸石为ZSM系列沸石和/或ZRP沸石，所述大孔沸石为选自稀土Y、稀土氢Y、超稳Y和高硅Y中的一种或多种。
6、根据项目5所述的方法，其中，以干基计，所述中孔沸石占沸石总重量的0-50重量％。
7、根据项目5所述的方法，其中，以干基计，所述中孔沸石占沸石总重量的0-20重量％。
8、根据项目1所述的方法，其中，所述第一催化裂解反应的条件包括：反应温度为500-600℃，反应时间为0.05-5秒，剂油重量比为(1-50)∶1，水油重量比为(0.03-0.5)∶1，催化剂密度为20-100千克/米 3，气体线速为4-18米/秒，反应压力为130-450千帕，催化剂质量流速G s为180-500千克/(米 2·秒)；
所述第二催化裂解反应的条件包括：反应温度为510-650℃，反应时间为1-20秒，剂油重量比为(3-50)∶1，水油重量比为(0.03-0.8)∶1，催化剂密度为120-290千克/米 3，气体线速为0.8-2.5米/秒，反应压力为130-450千帕，催化剂质量流速G s为15-150千克/(米 2·秒)。
9、根据项目1所述的方法，其中，所述第一催化裂解反应的条件包括：反应温度为520-580℃，反应时间为1-3秒，剂油重量比为(5-25)∶1，水油重量比为(0.05-0.3)∶1；
所述第二催化裂解反应的条件包括：反应温度为550-620℃，反应时间为3-15秒，剂油重量比为(10-30)∶1，水油重量比为(0.05-0.5)∶1，催化剂密度为150-250千克/米 3，气体线速为1-1.8米/秒，催化剂质量流速G s为20-130千克/(米 2·秒)。
10、根据项目1所述的方法，所述方法还包括：将C4烃馏分和/或C5-C6轻汽油馏分引入所述快速流化床和/或稀相输送床进行催化裂解反应。
11、根据项目10所述的方法，其中，在劣质重油引入稀相输送床的进料位置之前引入所述C4烃馏分和/或C5-C6轻汽油馏分。
12、根据项目1所述的方法，所述方法还包括：将催化剂补充入快速流化床中与所述第一反应产物和半待生催化剂一起进行所述第二催化裂解反应；其中，所补充催化剂的炭含量为0-1.0重量％。
13、根据项目12所述的方法，其中，所补充的催化剂占稀相输送床和快速流化床催化剂总循环量的0-50重量％。
14、根据项目12所述的方法，其中，所补充的催化剂占稀相输送床和快速流化床催化剂总循环量的5-30重量％。
15、根据项目12所述的方法，其中，催化剂在快速流化床补充的位置为快速流化床的底部。
16、一种催化裂解系统，该系统包括稀相输送床、快速流化床、油剂分离设备、反应产物分离设备和再生器；
按照反应物料的流向，所述稀相输送床与快速流化床流体流通且所述稀相输送床位于所述快速流化床上游；
所述稀相输送床设置有底部的催化剂入口、下部的劣质重油入口，所述快速流化床设置有顶部的油剂出口，所述油剂分离设备设置有油剂入口、催化剂出口和反应产物出口，所述反应产物分离设备设置有反应产物入口、干气出口、液化气出口、裂解汽油出口、裂解柴油出口和裂解重油出口，所述再生器设置有催化剂入口和催化剂出口；
所述稀相输送床的催化剂入口与所述再生器的催化剂出口流体连通，所述快速流化床的油剂出口与所述油剂分离设备的油剂入口流体连通，所述油剂分离设备的反应产物出口与所述反应产物分离设备的反应产物入口流体连通，所述油剂分离设备的催化剂出口与所述再生器的催化剂入口流体连通。
17、根据项目16所述的系统，其中，所述快速流化床与稀相输送床上下同轴设置且快速流化床位于稀相输送床上方。
本申请中各参数的定义和计算方式如下：
(1)、催化剂的轴向固体分率ε＝由压差计测得的反应区内轴向两点间的压差÷所述轴向两点间的距离÷催化剂颗粒密度；
其中压差的单位为千克/米 2，轴向两点间距离的单位为米，催化剂颗粒密度的单位为千克/米 3。
催化剂颗粒密度＝骨架密度/(催化剂孔体积×骨架密度+1)，其中 骨架密度的单位为千克/米 3，催化剂孔体积的单位为米 3，骨架密度和催化剂孔体积分别由比重瓶法和水滴定法测定。
(2)、反应时间＝反应区的体积/油气对数平均体积流量；
其中，反应区体积的单位为米 3，油气对数平均体积流量的单位为米 3/秒；
油气对数平均体积流量＝(V out-V in)/ln(V out/V in)，V out和V in分别为反应区出口处和入口处的油气体积流量；
反应区出口油气体积流量V out＝m/ρ 3；
反应区入口油气体积流量V in＝m/ρ 4；
其中，m为单位时间原料油和雾化蒸汽的进料量，单位为千克/秒；ρ 3为反应区出口处的油气密度，单位为千克/米 3；ρ 4为反应区入口处的油气密度，单位为千克/米 3。
(3)、反应区(或其上、中、下部)的催化剂密度＝由压差计测得的反应区内(或其上、中、下部)轴向两点间的压差÷所述轴向两点间的距离；
其中压差的单位为千克/米 2，反应区轴向平均分为上、中、下三部分，轴向两点间距离的单位为米。
(4)、气体线速＝油气对数平均体积流量÷反应区横截面积；
当反应区为非圆柱体型式时，气体线速取反应区底端的气体线速与反应区顶端的气体线速的对数平均值。
(5)、催化剂质量流率G s＝反应器催化剂循环量÷反应区横截面积；
当反应区为非圆柱体型式时，催化剂质量流率G s取反应区底端的G s与反应区顶端的G s的对数平均值；
其中，催化剂循环量的单位为千克/秒；
反应器催化剂循环量＝焦炭生成速度÷(待生催化剂炭含量-再生催化剂炭含量)；
其中，焦炭生成速度单位为千克/秒，待生催化剂炭含量和再生催化剂炭含量均为重量含量；
焦炭生成速度＝烟气量×(CO 2％+CO％)÷Vm×M；
其中，Vm为气体摩尔体积，取值为22.4×10 -3米 3/摩尔，M为碳元素的摩尔质量，取值为12×10 -3千克/摩尔；
烟气量＝(再生空气量×79体积％)/(1-CO 2％-CO％-O 2％)；
其中，再生空气量的单位为米 3/秒，烟气量的单位为米 3/秒，CO 2％、CO％、O 2％分别为烟气中CO 2、CO和O 2的体积百分比。
实施例
下面的实施例将对本申请方法予以进一步的说明，但并不因此而限制本申请。
以下实施例与对比例中所用的原料油均为加氢渣油，其性质如表1所示。所用的催化剂是购自中国石油化工股份有限公司催化剂分公司的商业催化裂解催化剂，商品牌号为DMMC-2。
表1所用原料油的性质
密度(20℃)/克·厘米 -3 0.9237
折光指数/70℃ 1.4914
碱性氮/微克·克 -1 506
残炭/重量％ 3.11
特性因数K值 11.8
馏程/℃  
5体积％ 357
10体积％ 387
30体积％ 443
50体积％ 490
70体积％ 550
金属含量/微克·克 -1  
Fe 34.4
Ni 4.4
Ca 7.8
V 4.3
Na 2.0
实施例1
按照图1所示的流程进行试验，原料油为加氢渣油，采用DMMC-2催化剂，在中型装置上进行试验，反应器为包括串联的稀相输送床和快速流化床的组合式反应器。预热的原料油进入稀相输送床 与催化裂解催化剂接触并进行第一裂解反应，反应流出物由下至上进入快速流化床与补充的再生催化剂混合后继续进行第二催化裂解反应。所补充的再生催化剂的炭含量为0.05重量％，催化剂补充位置位于快速流化床总高度的1/3高度处，所补充的催化剂占反应器催化剂循环量的5重量％。通过调整气体线速和在快速流化床底部设置伞形催化剂分布板，控制快速流化床中催化剂呈全浓相分布，快速流化床中催化剂的轴向固体分率ε由下至上均在0.1-0.2范围内。反应产物和待生催化剂快速分离，反应产物在产物分离系统中按馏程进行切割。待生催化剂在重力作用下进入汽提段，由水蒸气汽提出待生催化剂上吸附的反应产物，汽提后的催化剂不经换热直接进入到再生器，与空气接触进行烧焦再生，再生催化剂返回反应器中循环使用。所用的操作条件和产品分布列于表2。
从表2可以看出，本实施例的乙烯产率可达5.2重量％，丙烯产率可达18.2重量％，轻芳烃产率为11.5重量％，干气和焦炭产率分别为10.8重量％和8.5重量％。
实施例2
按照图2所示的流程进行试验，原料油为加氢渣油，采用DMMC-2催化剂，在中型装置上进行试验，反应器分为包括串联的稀相输送床和快速流化床的组合式反应器。预热的原料油进入全浓相反应区底部与催化裂解催化剂接触并进行第一裂解反应，反应流出物由下至上进入快速流化床的全浓相反应区与补充的再生催化剂混合后继续进行第二催化裂解反应。所补充的再生催化剂的炭含量为0.05重量％，催化剂补充位置位于快速流化床总高度的1/3高度处，所补充的催化剂占反应器催化剂循环量的5重量％。通过调整气体线速和在全浓相反应区底部设置伞形催化剂分布板，控制全浓相反应区中催化剂呈全浓相分布，所述全浓相反应区中催化剂的轴向固体分率ε由下至上均在0.1-0.2范围内。反应产物和待生催化剂快速分离，反应产物在产物分离系统中按馏程进行切割。待生催化剂在重力作用下进入汽提段，由水蒸气汽提出待生催化剂上吸附的反应产物，汽提后的催化剂不经换热直接进入到再生器，与空气接触进行烧焦再生，再生催化剂返回反应器中循环使用。反应产物切割后得到的混合C4馏分返回稀相输送床底部进行进一步反应。所用的操作条件和产品分布列于表2。
从表2可以看出，本实施例的乙烯产率可达5.9重量％，丙烯产率可达21.1重量％，轻芳烃产率为11.8重量％，干气和焦炭产率分别为10.7重量％和8.4重量％。
对比例1
原料油为加氢渣油，采用DMMC-2催化剂，在中型装置上进行试验，反应器型式为提升管与流化床串联的组合式反应器。预热的原料油进入提升管下部与催化剂接触进行催化裂解反应，反应油气和水蒸气以及待生催化剂从提升管出口进入密相流化床继续反应。反应后物流进入密闭式旋风分离器，使反应产物和待生催化剂快速分离，反应产物在产物分离系统中按馏程进行切割。待生催化剂在重力作用下进入汽提段，由水蒸气汽提出待生催化剂上吸附的反应产物，汽提后的催化剂不经换热直接进入到再生器，与空气接触进行烧焦再生，再生后的催化剂返回提升管中循环使用。所用的操作条件和产品分布列于表2。
从表2的结果可以看出，该对比例的乙烯产率为3.7重量％，丙烯产率为12.8重量％，轻芳烃产率为5.5重量％，干气和焦炭产率分别为12.9重量％和13.3重量％。
对比例2
对比例2与实施例1基本相同，区别在于快速流化床底部未设置催化剂分布板，快速流化床中催化剂的轴向固体分率ε由上至下表现为0.1→0.2→0.3递增。所用的操作条件同实施例1，产品分布列于表2。
表2实施例1-2和对比例1-2的反应结果对比
Figure PCTCN2019095950-appb-000001
由以上实施例和对比例的结果可以看出，本申请方法具有较高的乙烯、丙烯和轻芳烃产率，同时具有较低的干气和焦炭产率。
以上详细描述了本申请的优选实施方式，但是，本申请并不限于上述实施方式中的具体细节，在本申请的技术构思范围内，可以对本申请的技术方案进行多种简单变型，这些简单变型均属于本申请的保护范围。
另外需要说明的是，在上述具体实施方式中所描述的各个具体技术特征，在不矛盾的情况下，可以通过任何合适的方式进行组合，为了避免不必要的重复，本申请对各种可能的组合方式不再另行说明。
此外，本申请的各种不同的实施方式之间也可以进行任意组合，只要其不违背本申请的思想，其同样应当视为本申请所发明的内容。

Claims (16)

  1. 一种烃油催化裂解方法，包括使烃油原料、特别是重质原料油与催化裂解催化剂在包括串联的稀相输送床和快速流化床的反应器中接触反应的步骤，其中，在所述快速流化床中，所述催化剂的轴向固体分率ε控制在约0.1至约0.2的范围内。
  2. 根据权利要求1所述的方法，所述方法用于由重质原料油生产低碳烯烃，并且进一步包括如下步骤：
    i)使包括轻质原料和/或重质原料油的第一反应原料与催化裂解催化剂在稀相输送床中接触进行第一催化裂解反应；以及
    ii)使步骤i)的反应流出物和任选的包括轻质原料和/或重质原料油的第二反应原料在快速流化床中进行第二催化裂解反应，
    所述轻质原料选自C4烃馏分、C5-C6轻汽油馏分和它们的任意组合，且所述第一和第二反应原料中至少一者包括所述重质原料油；
    其中，在所述快速流化床中所述催化剂的轴向固体分率ε控制在约0.1至约0.2的范围内。
  3. 根据权利要求2所述的方法，进一步包括如下步骤：
    i)使所述第一反应原料与催化裂解催化剂在稀相输送床中接触进行第一催化裂解反应；
    ii)使步骤i)的反应流出物和任选的第二反应原料在快速流化床中、在有效生成低碳烯烃的条件下进行第二催化裂解反应，其中在所述快速流化床中所述催化剂的轴向固体分率ε控制在约0.1至约0.2的范围内；
    iii)将来自所述反应器的反应流出物分离，得到富含低碳烯烃的反应产物和待生催化剂；
    iv)使所述待生催化剂再生，并将所得再生催化剂的至少一部分返回步骤i)作为所述催化裂解催化剂；以及
    v)任选地，将所述反应产物分离得到干气、液化气、裂解汽油、裂解柴油和裂解重油。
  4. 根据在先权利要求中任一项所述的方法，进一步包括：在所述快速流化床中引入一股或多股补充催化剂，并使其与所述快速流化床内的物料接触进行催化裂解反应，
    其中，所述一股或多股补充催化剂的炭含量各自独立地为约0-1.0重量％，并且各自独立地选自再生、半再生或待生的催化裂解催化剂，且所述一股或多股补充催化剂的总量占反应器催化剂循环量的约0-50重量％，优选约5-30重量％；
    所述一股或多股补充催化剂的引入位置各自独立地位于所述快速流化床总高度的约0％至约90％高度处，优选约20％到约80％高度处，更优选约30％到约75％高度处。
  5. 根据在先权利要求中任一项所述的方法，其中，以干基计并以催化裂解催化剂的干基重量为基准，所述催化裂解催化剂包括约1-50重量％，优选约5-45重量％，更优选约10-40重量％的沸石、约5-99重量％，优选约10-80重量％，更优选约20-70重量％的无机氧化物，和约0-70重量％，优选约5-60重量％，更优选约10-50重量％的粘土；
    所述沸石包括中孔沸石和任选的大孔沸石，所述中孔沸石选自ZSM系列沸石、ZRP沸石，和它们的任意组合，所述大孔沸石选自稀土Y型沸石、稀土氢Y型沸石、超稳Y型沸石、高硅Y型沸石，和它们的任意组合；
    优选地，以干基计，所述中孔沸石占沸石总重量的约0-50重量％，更优选约0-20重量％。
  6. 根据权利要求2-5中任一项所述的方法，其中，所述第一催化裂解反应的条件包括：反应温度为约500-600℃，反应时间为约0.05-5秒，剂油重量比为约1∶1至约50∶1，水油重量比为约0.03∶1至约0.5∶1，催化剂密度为约20-100千克/米 3，气体线速为约4-18米/秒，反应压力为约130-450千帕，催化剂质量流率G s为约180-500千克/(米 2·秒)，
    优选地，所述第一催化裂解反应的条件包括：反应温度为约520-580℃，反应时间为约1-3秒，剂油重量比为约5∶1至约25∶1，水油重量比为约0.05∶1至约0.3∶1；以及
    所述第二催化裂解反应的条件包括：反应温度为约510-650℃，反应时间为约1-20秒，剂油重量比为约3∶1至约50∶1，水油重量比为约0.03∶1至约0.8∶1，催化剂密度为约120-290千克/米 3，气体线速为约0.8-2.5米/秒，反应压力为约130-450千帕，催化剂质量流率G s为约15-150千克/(米 2·秒)，
    优选地，所述第二催化裂解反应的条件包括：反应温度为约550-620℃，反应时间为约3-15秒，剂油重量比为约10∶1至约30∶1，水油重量比为约0.05∶1至约0.5∶1，催化剂密度为约150-250千克/米 3，气体线速为约1-1.8米/秒，反应压力为约130-450千帕，催化剂质量流率G s为约20-130千克/(米 2·秒)。
  7. 根据在先权利要求中任一项所述的方法，其中，所述重质原料油为劣质重油，其性质满足以下指标中的至少一种：20℃密度为约900-1000千克/米 3，残炭为约2-10重量％，镍和钒的总含量为约2-30ppm，特性因数K值小于约12.1；
    优选满足以下指标中的至少一种：20℃密度为约910-940千克/米 3，残炭为约3-8重量％，镍和钒的总含量为约5-20ppm，特性因数K值小于约12.0。
  8. 根据在先权利要求中任一项所述的方法，其中，所述重质原料油为选自重质石油烃、其它矿物油和它们的任意组合的劣质重油；
    其中，所述重质石油烃选自减压渣油、劣质的常压渣油、劣质的加氢渣油、焦化瓦斯油、脱沥青油、减压蜡油、高酸值原油、高金属原油，和它们的任意组合；并且
    所述其它矿物油选自煤液化油、油砂油、页岩油，和它们的任意组合。
  9. 根据权利要求2-8中任一项所述的方法，其中所述第一反应原料和第二反应原料中至少一者包括所述选自C4烃馏分、C5-C6轻汽油馏分和它们的任意组合的轻质原料；
    优选地，所述第一反应原料包括轻质原料和重质原料油，并且所述轻质原料的至少一部分在所述重质原料油引入所述稀相输送床的位置的上游引入所述稀相输送床中；
    优选地，所述第一反应原料包括重质原料油，而所述第二反应原料包括所述轻质原料；或者
    优选地，所述第一反应原料包括所述轻质原料，而所述第二反应原料包括重质原料油。
  10. 根据在先权利要求中任一项所述的方法，其中所述快速流化床从下到上依次包括全浓相反应区和过渡段，所述全浓相反应区为横截面呈大致圆形的、底端和顶端开口的、等直径或者变直径的空心柱 体的型式，所述稀相输送床与所述全浓相反应区的底端相连通，所述全浓相反应区的顶端经由所述过渡段与所述反应器的出口段相连通，所述全浓相反应区的底部任选设有一个或多个供所述第二反应原料进料的入口，
    其中，所述全浓相反应区的底端的横截面直径大于或等于所述稀相输送床的直径，且顶端的横截面直径大于所述出口段的直径，并且所述全浓相反应区的底部或侧壁设有一个或多个补充催化剂入口，所述一个或多个补充催化剂入口的位置各自独立地位于所述快速流化床总高度的约0％至约90％的高度处，优选约20％到约80％高度处，更优选约30％到约75％高度处。
  11. 根据权利要求10所述的方法，其中所述全浓相反应区为等直径的空心圆柱体型式，或者为由下至上直径连续地或不连续地增大的空心柱体型式，例如倒置的空心截头圆锥体、由两段或更多段直径递增的圆柱体构成的空心柱体、由两段或更多段直径递增的倒置的截头圆锥体构成的空心柱体、或者由一段或多段圆柱体与一段或多段倒置的截头圆锥体构成的空心柱体的型式。
  12. 一种适用于烃油、特别是重质原料油的催化裂解的系统，该系统包括催化裂解反应器、油剂分离设备、任选的反应产物分离设备、和再生器，
    所述催化裂解反应器包括串联的稀相输送床和快速流化床，按照反应物料的流向，所述稀相输送床与快速流化床流体流通且所述稀相输送床位于所述快速流化床的上游；
    所述稀相输送床设置有底部的催化剂入口、下部的第一反应原料入口，所述快速流化床设置有顶部的油剂出口和任选的底部的第二反应原料入口，所述油剂分离设备设置有油剂入口、催化剂出口和反应产物出口，所述任选的反应产物分离设备设置有反应产物入口、干气出口、液化气出口、裂解汽油出口、裂解柴油出口和裂解重油出口，所述再生器设置有催化剂入口和催化剂出口，
    所述稀相输送床的催化剂入口与所述再生器的催化剂出口流体连通，所述快速流化床的油剂出口与所述油剂分离设备的油剂入口流体连通，所述油剂分离设备的反应产物出口与所述任选的反应产物分离设备的反应产物入口流体连通，所述油剂分离设备的催化剂出口与所 述再生器的催化剂入口流体连通。
  13. 根据权利要求12所述的系统，其中，所述快速流化床与所述稀相输送床上下同轴设置且快速流化床位于稀相输送床的上方，
    优选地，所述快速流化床的底部设有催化剂分布板，和/或所述第二反应原料入口的位置设有气体分布器。
  14. 根据权利要求12或13所述的系统，其中所述快速流化床从下到上依次包括全浓相反应区和过渡段，所述全浓相反应区为横截面呈大致圆形的、底端和顶端开口的、等直径或者变直径的空心柱体的型式，所述稀相输送床与所述全浓相反应区的底端相连通，所述全浓相反应区的顶端经由所述过渡段与所述反应器的出口段相连通，所述全浓相反应区的底部任选设有一个或多个第二反应原料入口，
    其中，所述全浓相反应区的底端的横截面直径大于或等于所述稀相输送床的直径，且顶端的横截面直径大于所述出口段的直径，并且所述全浓相反应区的底部或侧壁设有一个或多个补充催化剂入口，所述一个或多个补充催化剂入口的位置各自独立地位于所述快速流化床总高度的约0％至约90％的高度处，优选约20％到约80％高度处，更优选约30％到约75％高度处。
  15. 根据权利要求14所述的系统，其中所述全浓相反应区为等直径的空心圆柱体型式，或者为由下至上直径连续地或不连续地增大的空心柱体型式，例如倒置的空心截头圆锥体、由两段或更多段直径递增的圆柱体构成的空心柱体、由两段或更多段直径递增的倒置的截头圆锥体构成的空心柱体、或者由一段或多段圆柱体与一段或多段倒置的截头圆锥体构成的空心柱体的型式。
  16. 根据权利要求15所述的系统，其中所述全浓相反应区为倒置的空心截头圆锥体型式，纵切面为等腰梯形，其底端横截面的直径为约0.2-10米，优选约0.5-8米，更优选约1-5米；顶端横截面直径与底端横截面直径的比值为大于1至约50，优选约1.2至约10，更优选约1.5至约5；最大横截面的直径与快速流化床总高度之比为约0.005∶1至约1∶1，优选约0.01∶1至约0.8∶1，更优选约0.05∶1至约0.5∶1；全浓相反应区的高度与快速流化床总高度之比为约0.1∶1至约0.9∶1，优选约0.3∶1至约0.85∶1，更优选约0.5∶1至约0.8∶1，和/或
    所述过渡段为空心截头圆锥体型式，纵切面为等腰梯形，等腰梯 形侧边的内倾角α为约25-85°，优选约30-75°，且所述过渡段的高度与快速流化床总高度之比为约0.1∶1至约0.9∶1，优选约0.2∶1至约0.5∶1。
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