WO2019080791A1 - 一种多产异丁烷和/或轻质芳烃的催化裂化方法 - Google Patents

一种多产异丁烷和/或轻质芳烃的催化裂化方法

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Publication number
WO2019080791A1
WO2019080791A1 PCT/CN2018/111179 CN2018111179W WO2019080791A1 WO 2019080791 A1 WO2019080791 A1 WO 2019080791A1 CN 2018111179 W CN2018111179 W CN 2018111179W WO 2019080791 A1 WO2019080791 A1 WO 2019080791A1
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Prior art keywords
oil
catalytic cracking
reaction
catalyst
weight
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PCT/CN2018/111179
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English (en)
French (fr)
Inventor
许友好
王新
张毓莹
刘涛
白旭辉
戴立顺
张执刚
梁家林
姜楠
Original Assignee
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 KR1020207014490A priority Critical patent/KR102574048B1/ko
Priority to RU2020115286A priority patent/RU2776179C2/ru
Priority to SG11202003403WA priority patent/SG11202003403WA/en
Priority to US16/756,424 priority patent/US11427773B2/en
Publication of WO2019080791A1 publication Critical patent/WO2019080791A1/zh

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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
    • C10G69/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process
    • C10G69/02—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only
    • C10G69/08—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only including at least one step of reforming naphtha
    • 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
    • C10G69/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process
    • C10G69/02—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only
    • C10G69/04—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only including at least one step of catalytic cracking in the absence of hydrogen
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C4/00—Preparation of hydrocarbons from hydrocarbons containing a larger number of carbon atoms
    • C07C4/02—Preparation of hydrocarbons from hydrocarbons containing a larger number of carbon atoms by cracking a single hydrocarbon or a mixture of individually defined hydrocarbons or a normally gaseous hydrocarbon fraction
    • C07C4/06—Catalytic processes
    • 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
    • C10G11/02—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils characterised by the catalyst used
    • C10G11/04—Oxides
    • C10G11/05—Crystalline alumino-silicates, e.g. molecular sieves
    • 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
    • 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
    • C10G21/00—Refining of hydrocarbon oils, in the absence of hydrogen, by extraction with selective solvents
    • C10G21/06—Refining of hydrocarbon oils, in the absence of hydrogen, by extraction with selective solvents characterised by the solvent used
    • C10G21/12—Organic compounds only
    • 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
    • C10G69/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C9/00—Aliphatic saturated hydrocarbons
    • C07C9/02—Aliphatic saturated hydrocarbons with one to four carbon atoms
    • C07C9/10—Aliphatic saturated hydrocarbons with one to four carbon atoms with four carbon atoms
    • C07C9/12—Iso-butane
    • 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
    • C10G2300/1096—Aromatics or polyaromatics
    • 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
    • C10G2300/4018—Spatial velocity, e.g. LHSV, WHSV
    • 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
    • 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
    • C10G2400/30—Aromatics

Definitions

  • the present application relates to the field of catalytic cracking technology, and in particular to a catalytic cracking process for producing isobutane and/or light aromatic hydrocarbons.
  • Benzene, toluene and xylene (BTX) in aromatic hydrocarbons are important chemical raw materials.
  • Benzene can be used to synthesize products such as styrene, phenol and aniline.
  • Toluene can be used as an excellent solvent for organic synthesis and is an ideal raw material for the synthesis of cresol.
  • O-xylene, m-xylene and p-xylene in xylene are the basic raw materials for organic synthesis.
  • Isobutane is used as an important chemical raw material for alkylation with C 3- C 5 olefins to produce alkylated oils; co-oxidation with propylene to produce propylene oxide (PO), and parallel production of tert-butanol (TBA) either to produce methyl tert-butyl ether (MTBE); or to prepare isobutylene by isobutane dehydrogenation.
  • PO propylene oxide
  • TAA tert-butanol
  • MTBE methyl tert-butyl ether
  • the conventional catalytic cracking process is mainly used for the production of gasoline, and has achieved a breakthrough in gasoline yield of up to 50% by weight, which satisfies the requirement of lead-free gasoline and improves the octane number of gasoline.
  • the gasoline octane number is increased by changing the process conditions or by using a novel zeolite catalyst, the octane number of the gasoline is increased by increasing the olefin content in the gasoline composition.
  • the content of olefins in the catalytic cracking gasoline composition can reach 35-65% by weight, which is far from the requirements of the Chinese national standard for olefin content.
  • the liquefied gas composition has a higher olefin content of about 70% by weight, wherein the butene is several times that of isobutane, which is difficult to use as an alkylation raw material.
  • a process for the preparation of BTX from catalytically cracked naphtha is disclosed in U.S. Patent No. 5,685,972.
  • the raw materials used in the process are catalytically cracked naphtha and cokered naphtha, and the catalyst used is preferably ZSM-5 or a catalyst having a hydrogenation functional component.
  • Chinese Patent Application Publication No. CN104560166A discloses a catalytic conversion process in which a catalytic cracking light cycle oil is cut to obtain a light fraction and a heavy fraction, and a heavy fraction is hydrotreated to obtain a hydrogenated heavy fraction, and the light fraction and the hydrogenated heavy fraction are separately separated into
  • the secondary riser reactor of the catalytic cracking unit the heavy petroleum hydrocarbons enter the main riser reactor of the catalytic cracking unit.
  • the process maximizes the critical conditions required to meet the catalytic cracking reactions of different fractions of light cycle oils to maximize the production of high octane catalytic cracked gasoline.
  • Cipheral Patent Application Publication No. CN1232069A discloses a catalytic conversion method for preparing isobutane and isoparaffin-rich gasoline, which has improved heavy oil processing capacity, decreased dry gas and slurry yield, and greatly reduced gasoline olefin and sulfur content, and device Energy consumption is further reduced.
  • the present application provides a catalytic cracking method comprising the following steps:
  • the separation in step c) further results in a light cycle oil fraction, a heavy cycle oil fraction, and optionally a slurry, and the method further comprises the steps of:
  • step c) hydrotreating at least a portion of the light cycle oil fraction, the heavy cycle oil fraction and the optional slurry obtained in step c) to obtain a hydrogenated tail oil;
  • step a) further comprises pretreating the starting stock oil having a bicycloacyclic or higher naphthenic content of no greater than about 25% by weight to provide catalytic cracking of said bicyclo or higher naphthenic content greater than about 25% by weight.
  • Raw oil having a bicycloacyclic or higher naphthenic content of no greater than about 25% by weight
  • the pretreatment comprises aromatic extraction and/or hydrotreating.
  • the first catalytic cracking reaction is carried out under conditions of a reaction temperature of about 520-620 ° C, a reaction time of about 0.5-3.0 seconds, and a weight ratio of the agent oil of from about 3:1 to about 15:
  • the second catalytic cracking reaction is carried out under the following conditions: a reaction temperature of about 480 to 600 ° C, a reaction time of about 2 to 30 seconds, and a weight ratio of the agent oil of from about 3:1 to about 18:1.
  • the hydrotreating is carried out under the following conditions: a hydrogen partial pressure of about 6.0-30.0 MPa, a reaction temperature of about 300-450 ° C, and a liquid phase volumetric space velocity of about 0.1-10.0 h -1 . hydrogen oil volume ratio of about 300-3000Nm 3 / m 3.
  • feedstock oil rich in polycyclic naphthenes or feedstock oil which can be obtained by pretreatment to obtain high polycyclic naphthenic content, so as to achieve efficient use of such feedstock oil;
  • the gasoline yield is greater than 40% by weight, and the content of light aromatic hydrocarbons in the product gasoline is increased.
  • FIG. 1 is a schematic flow diagram of a preferred embodiment of the method disclosed herein.
  • FIG. 2 is a schematic flow diagram of another preferred embodiment of the method disclosed herein.
  • FIG. 3 is a schematic flow diagram of still another preferred embodiment of the method disclosed herein.
  • any specific numerical values (including the endpoints of the numerical ranges) disclosed herein are not limited to the precise value of the numerical value, but should be understood to cover the value close to the precise value. Moreover, for the disclosed numerical range, one or more new ones can be obtained between the endpoint values of the range, the endpoint values and the specific point values in the range, and the specific point values. Numerical ranges, these new numerical ranges are also considered to be specifically disclosed herein.
  • any matters or matters not mentioned are directly applicable to those known in the art without any change other than those explicitly stated.
  • any of the embodiments described herein can be freely combined with one or more other embodiments described herein, and the resulting technical solution or technical idea is considered to be part of the original disclosure or original description of the present invention, and should not be It is considered to be new content that has not been disclosed or anticipated herein, unless it is apparent to those skilled in the art that the combination is clearly unreasonable.
  • bicyclocycloalkane and “polycyclic cycloalkane” are used interchangeably and refer to a cycloalkane having more than two carbocycles.
  • bicyclic aromatic hydrocarbon and “polycyclic aromatic hydrocarbon” are used interchangeably and refer to an aromatic compound having two or more aromatic rings.
  • water to oil ratio refers to the mass ratio of atomized steam to feedstock oil.
  • the present application provides a catalytic cracking method comprising the following steps:
  • the catalytic cracking feedstock oil has a bicyclo or higher cycloalkane content of greater than about 40% by weight, and the higher the content, the better.
  • the separation in step c) further results in a light cycle oil fraction, a heavy cycle oil fraction, and optionally a slurry, and the method further comprises the steps of:
  • step c) hydrotreating at least a portion of the light cycle oil fraction, the heavy cycle oil fraction and the optional slurry obtained in step c) to obtain a hydrogenated tail oil;
  • the hydrotreating in step e) is used to produce a hydrogenated tailstock (or referred to as a hydrogenated distillate) enriched in cyclohexane or higher.
  • the hydrotreating can be carried out under the following conditions: a hydrogen partial pressure of about 6.0 to 30.0 MPa, preferably about 8 to 20 MPa, a reaction temperature of about 300 to 450 ° C, preferably about 330 to 430 ° C, and a liquid. phase space velocity of about 0.1-10.0h -1, preferably about 0.2-5h -1, hydrogen to oil volume ratio of about 300-3000Nm 3 / m 3, preferably from about 500-2500Nm 3 / m 3.
  • the hydrotreating can be carried out under the following conditions: a hydrogen partial pressure of about 8-20 MPa, a reaction temperature of about 330-430 ° C, a liquid phase volume space velocity of about 0.2-5 h -1 , and a hydrogen oil volume ratio of about 500-2500Nm 3 / m 3.
  • the hydrotreating may be carried out in the presence of a hydrotreating catalyst, which may comprise a hydrogenation active component and a support
  • the hydrogenation active component may be selected from Group VIB non-noble metals, Group VIII non-noble metals, and combinations thereof, may be selected from the group consisting of alumina, silica, and amorphous silica-alumina, and combinations thereof.
  • the Group VIII non-noble metal is present in an amount of from about 1 to 99% by weight, preferably from about 1 to 60% by weight, based on the weight of the hydrotreating catalyst; and/or
  • the Group VIB non-noble metal is present in an amount of from about 1 to 99% by weight, preferably from about 1 to 70% by weight.
  • the Group VIII non-noble metal is cobalt and/or nickel
  • the Group VIB non-noble metal is molybdenum and/or tungsten.
  • the hydrotreating catalyst has excellent aromatic hydrogenation saturation ability, and can effectively hydrogenate the light cycle oil fraction, the heavy cycle oil fraction and the aromatic hydrocarbon in the slurry to a naphthenic hydrocarbon, thereby providing the catalytic cracking reaction of the present invention. raw material.
  • the slurry produced by the catalytic cracking reaction may optionally enter the slurry filtration system and then enter the hydrogenation unit for hydrogenation and at least partially recycle to the catalytic cracking reactor. Or the resulting slurry can be directly pulled out of the device.
  • step a) further comprises pretreating the starting stock oil having a bicycloacyclic or higher naphthenic content of no greater than about 25% by weight to obtain a naphthene or higher naphthenic content of greater than about 25% by weight, preferably More than about 40% by weight of the catalytic cracking feedstock oil.
  • the starting feedstock having a bicyclocyclo-cycloalkane content of no greater than about 25% by weight may be a feedstock oil having a total content of bicyclo- or higher cycloalkanes and bicyclic or higher aromatics greater than about 25 wt%.
  • the initial feedstock oil may be hydrotreated to hydrogenate the bicyclic or higher aromatic hydrocarbon to a bicyclo or higher cycloalkane to obtain a catalytically cracked feedstock oil having a bicyclo or higher cycloparaffin content of greater than about 25% by weight.
  • the hydrotreating can be carried out under the following conditions: a hydrogen partial pressure of about 6.0 to 30.0 MPa, preferably about 8 to 20 MPa, a reaction temperature of about 300 to 450 ° C, preferably about 330 to 430 ° C, and a liquid. phase space velocity of about 0.1-10.0h -1, preferably about 0.2-5h -1, hydrogen to oil volume ratio of about 300-3000Nm 3 / m 3, preferably from about 500-2500Nm 3 / m 3.
  • the hydrotreating may be carried out in the presence of a hydrotreating catalyst, which may comprise a hydrogenation active component and a support
  • the hydrogenation active component may be selected from Group VIB non-noble metals, Group VIII non-noble metals, and combinations thereof, may be selected from the group consisting of alumina, silica, and amorphous silica-alumina, and combinations thereof.
  • the Group VIII non-noble metal is present in an amount of from about 1 to 99% by weight, preferably from about 1 to 60% by weight, based on the weight of the hydrotreating catalyst; and/or
  • the Group VIB non-noble metal is present in an amount of from about 1 to 99% by weight, preferably from about 1 to 70% by weight.
  • the Group VIII non-noble metal is cobalt and/or nickel, and the Group VIB non-noble metal is molybdenum and/or tungsten.
  • the aromatic feedstock may be subjected to an aromatic hydrocarbon extraction of a biogenic hydrocarbon having a naphthene content of not more than about 25% by weight, and then optionally subjected to the above hydrotreatment.
  • the initial feedstock oil is separated into a polycyclic naphtheter-rich pumping oil and a polycyclic aromatic hydrocarbon-rich extracted oil by aromatic extraction; the extracted oil is optionally further subjected to hydrorefining to make the polycyclic ring therein
  • the aromatic hydrocarbon is saturated to form a polycyclic cycloalkane.
  • the aromatics extraction can be carried out in a manner well known to those skilled in the art.
  • the aromatic hydrocarbon extraction can be carried out under the following conditions: a temperature of about 50-70 ° C, a solvent to raw material weight ratio of about 0.5-2, and a solvent selected from the group consisting of furfural, dimethyl sulfoxide, dimethylformamide, Monoethanolamine, ethylene glycol, and 1,2-propanediol, and combinations thereof.
  • the catalytic cracking feedstock oil suitable for use in the process disclosed herein can be any bicyclic or higher naphthenic content greater than about 25% by weight, preferably greater than about 40% by weight of the feedstock oil; the initial feedstock oil suitable for use in the process disclosed herein can be any After the pretreatment, a feedstock oil having a cyclocycloalkane content of greater than about 25% by weight, preferably greater than about 40% by weight, can be achieved.
  • the catalytic cracking feedstock oil or initial feedstock oil may be selected from the group consisting of deep hydrogenated light cycle oil, heavy distillate (CGO) of delayed coker unit, catalytic cracking light cycle oil (LCO), catalytic cracking heavy cycle oil (HCO).
  • catalytic cracking heavy distillate FGO
  • oil slurry hydrocracked diesel
  • residual hydrocracking diesel wax hydrocracking diesel
  • biodiesel shale oil diesel fraction
  • coal liquefied diesel fraction atmospheric pressure tower Oil, atmospheric distillation tower distillate, straight-run vacuum oil, hydrogenated wax oil, coking wax oil, deasphalted oil (DAO), oil extraction, pumping oil, atmospheric residue, vacuum residue, The hydrogenation tail oil obtained by hydrogenating the above raw material oil, and a combination thereof.
  • DAO deasphalted oil
  • the first and second catalytic cracking reactions can be carried out under different reaction conditions in a manner well known to those skilled in the art, wherein the first catalytic cracking reaction mainly comprises a cracking reaction.
  • the second catalytic cracking reaction mainly includes a selective hydrogen transfer reaction, an isomerization reaction, and an aromatization reaction.
  • the first catalytic cracking reaction can be carried out under the following conditions: a reaction temperature of about 520-620 ° C, preferably about 530-600 ° C; a reaction time of about 0.5-3.0 seconds, preferably about 0.8-2.0 seconds;
  • the oil weight ratio is from about 3:1 to about 15:1, preferably from about 4:1 to about 12:1; the water to oil ratio is from about 0.03:1 to about 0.3:1, preferably from about 0.05:1 to about 0.3:1.
  • the pressure is 130-450 kPa; and/or the second catalytic cracking reaction can be carried out under the following conditions: a reaction temperature of about 480-600 ° C, preferably about 500-550 ° C, or about 420-530 ° C, preferably about 460-510 ° C; reaction time is about 2-30 seconds, preferably about 3-15 seconds; the weight ratio of the agent oil is from about 3:1 to about 18:1, preferably from about 4:1 to about 15:1; water to oil ratio It is from about 0.03:1 to about 0.3:1, preferably from about 0.05:1 to about 0.3:1; and the pressure is from 130 to 450 kPa.
  • the catalytic cracking reactor suitable for use in the process disclosed herein can be any catalytic cracking reactor well known to those skilled in the art, as long as the first and second catalytic cracking reactions can be carried out under different reaction conditions.
  • the catalytic cracking reactor may be a composite of an equal diameter riser reactor, an equal line riser reactor, a variable diameter riser reactor, a fluidized bed reactor, or an equal diameter riser and a fluidized bed.
  • the reactor preferably a reduced diameter riser reactor.
  • the catalytic cracking reactor is a variable diameter riser reactor, and the variable diameter riser reactor is provided with a coaxial and fluidly connected pre-lifting section in a vertical direction from bottom to top, a first reaction zone, a second reaction zone and an outlet zone, the end of the outlet zone is provided with a horizontal pipe for connecting the settler, the inner diameter of the first reaction zone is smaller than the second reaction zone, and the inner diameter of the second reaction zone More than the outlet zone, the catalytic cracking catalyst is sent to the pre-lifting section, the catalytic cracking feedstock oil is sent to the lower portion of the first reaction zone, and the first catalytic cracking reaction is carried out in the first reaction zone. The second catalytic cracking reaction is carried out in the second reaction zone.
  • the total height of the reduced riser reactor (including the pre-elevation section, the first reaction zone, the second reaction zone, and the outlet zone) is between about 10 and 60 meters.
  • the pre-lift section has the same diameter as in a conventional equal-diameter riser reactor, typically from about 0.2 to 5 meters, and its height is from about 5 to 20% of the total height of the reactor.
  • the structure of the first reaction zone is similar to that of a conventional equal-diameter riser reactor, and its diameter may be the same as that of the pre-lift section, or may be slightly larger than the pre-lift section, and the ratio of the diameter of the first reaction zone to the diameter of the pre-lift section is From about 1:1 to about 2:1, the height is from about 10% to about 30% of the total height of the reactor.
  • the cracking reaction mainly occurs at a higher reaction temperature and a ratio of the agent to the oil, and a shorter reaction time (generally about 0.5 to 3.0 seconds).
  • the second reaction zone has a larger diameter than the first reaction zone, and the ratio of the diameter to the diameter of the first reaction zone is from about 1.5:1 to about 5.0:1, and the height thereof is about 10-60% of the total height of the reactor. It is to reduce the flow rate and reaction temperature of oil and gas, and to cause selective hydrogen transfer reaction, isomerization reaction and aromatization reaction.
  • the reaction time of the oil and gas in the reaction zone may be long, for example, about 2 to 30 seconds.
  • the ratio of the diameter of the outlet zone to the diameter of the first reaction zone is from about 0.8:1 to about 1.5:1, the height of which is from about 0% to about 50% of the total height of the reactor.
  • the binding sites of the first and second reaction zones are in the shape of a truncated cone, and the apex angle ⁇ of the isosceles trapezoid in the longitudinal section is about 30-80°, and the joint portion of the second reaction zone and the exit zone is a truncated cone shape, and the longitudinal section of the isosceles trapezoidal top
  • the angle ⁇ is about 45-85°.
  • the method for lowering the reaction temperature of the second reaction zone may include injecting a chilling medium from a joint portion of the zone and the first reaction zone, and/or suppressing the second time by providing a heat extractor in the zone to remove a part of heat The purpose of cracking reaction, increasing isomerization reaction and hydrogen transfer reaction.
  • the chilling medium can be selected from the group consisting of a cold horn, a cooled regenerated catalyst, and a cooled semi-regenerated catalyst, and combinations thereof.
  • the cold shock agent may be selected from the group consisting of a liquefied gas, a crude gasoline, a stabilized gasoline, a light cycle oil fraction, a heavy light cycle oil fraction, water, and combinations thereof; the cooled regenerated catalyst and the cooled semi-regenerated catalyst may be treated
  • the biocatalyst is obtained by two stages of regeneration and a section of regeneration and then cooling.
  • the carbon content of the regenerated catalyst is generally about 0.1% by weight or less, preferably about 0.05% by weight or less, and the carbon content of the semi-regenerated catalyst is generally from about 0.1 to 0.9% by weight, preferably from about 0.15 to 0.7% by weight. If a heat extractor is provided, its height is about 50-90% of the height of the second reaction zone.
  • the bonding region of the first reaction zone and the second reaction zone is provided with at least one chill medium inlet for inputting a chill medium; and/or the second reaction zone is provided with The heat extractor has a height of about 50-90% of the height of the second reaction zone.
  • the chill medium is selected from the group consisting of a cold shock agent, a cooled regenerated catalyst, a cooled semi-regenerated catalyst, a fresh catalyst, and a combination thereof;
  • the cold shock agent is selected from the group consisting of liquefied gas, coarse Gasoline, stabilized gasoline, light cycle oil, heavy cycle oil, water, and combinations thereof.
  • Catalytic cracking catalysts suitable for use in the processes disclosed herein can be those well known to those skilled in the art.
  • the catalytic cracking catalyst may comprise a cracking active component and a carrier; the cracking active component may comprise from about 0 to about 100% by weight, preferably on a dry basis, and based on the weight of the cracking active component.
  • the FAU-type zeolite is preferably selected from the group consisting of Y-type zeolite, HY-type zeolite, ultra-stable Y-type zeolite, and combinations thereof, the five-membered ring
  • the structural zeolite is preferably selected from the group consisting of ZSM-5 series zeolites, high silica zeolites, ferrierites, and combinations thereof, which may optionally contain rare earths and/or phosphorus.
  • the catalyst to be produced after the catalytic cracking catalyst is used can be sent to a regenerator for charring regeneration, and the regenerated catalyst obtained by the regeneration is recycled to the catalytic cracking catalyst as a catalytic cracking catalyst.
  • a portion of the catalyst to be produced in the settler may optionally be sent to the second reaction zone as a chill medium. To produce isoparaffins and light aromatics.
  • the light aromatic hydrocarbon (BTX) is recovered from the gasoline fraction by extraction refining in step d).
  • the extraction purification can be carried out in a manner well known to those skilled in the art, for example, the extraction purification can be carried out under the following conditions: the extraction solvent is selected from the group consisting of sulfolane, dimethyl sulfoxide, N-formylmorpholine, and tetra Glycol, triethylene glycol, N-methylpyridinone, and combinations thereof, at a temperature of from about 50 to about 110 ° C, and a weight ratio of extraction solvent to gasoline fraction of from about 2 to about 6.
  • the catalytic cracking process disclosed herein includes the following steps:
  • step c) optionally, the gasoline fraction obtained in step c) is subjected to extraction and purification to obtain a light aromatic hydrocarbon (for example, benzene, toluene, xylene), and/or the liquefied gas fraction obtained in the step c) is separated.
  • a light aromatic hydrocarbon for example, benzene, toluene, xylene
  • step e) optionally, the light cycle oil fraction obtained in step c), the heavy cycle oil fraction and optionally the oil slurry (preferably filtered) are fed to a hydrogenation unit in contact with a hydrotreating catalyst and hydrotreated; as well as
  • the resulting hydrogenated tailings are fed to the catalytic cracking reactor as the catalytic cracking feedstock oil.
  • the preheated catalytic cracking feedstock oil is sent to the variable diameter riser reactor, and enters the catalyst from the lower portion of the first reaction zone of the reactor, mainly causing cracking.
  • Reaction; the mixture of oil and gas formed after the reaction is carried up to the lower part of the second reaction zone of the reactor, optionally in contact with the cooled catalyst, mainly performing a hydrogen transfer reaction and an isomerization reaction, and the effluent enters the settler after the reaction
  • the reaction oil and gas (ie, the catalytic cracking reaction product) and the catalyst to be produced are separated.
  • the reaction oil is separated to obtain a reaction product including a dry gas, a liquefied gas fraction, a gasoline fraction, a light cycle oil fraction, a heavy cycle oil fraction, and a slurry.
  • a reaction product including a dry gas, a liquefied gas fraction, a gasoline fraction, a light cycle oil fraction, a heavy cycle oil fraction, and a slurry.
  • the light cycle oil fraction, the heavy cycle oil distillate and/or the slurry, and optionally the distillate from the other processing unit are sent to a hydrogenation unit for hydrotreating under the action of a hydrotreating catalyst,
  • the resulting hydrogenated tail oil is returned to the riser reactor.
  • the spent catalyst is stripped and regenerated and returned to the riser reactor.
  • Fig. 1 schematically shows an exemplary flow for producing isobutane and aromatic-rich gasoline using a variable diameter riser reactor, wherein the shape, size and the like of the equipment and piping are not limited by the drawings, but Determine according to the specific circumstances.
  • the pre-lifting steam enters from the pre-lift section 2 of the variable-diameter riser reactor via line 1, and the hot regenerated catalyst enters the pre-lift section 2 via the regeneration inclined tube 16 and is lifted by the pre-lifting steam.
  • the catalytic cracking feedstock oil from line 50 e.g., distillate oil from other processing units
  • the catalytic cracking feedstock oil from line 50 is combined with the hydrogenated tail oil from line 41 and passed through line 4 from the pre-elevation section 2 with atomized steam from line 3.
  • After mixing with the hot regenerated catalyst it enters the first reaction zone 5 where the first catalytic cracking reaction is carried out.
  • the reactant stream is mixed with a cold shock agent from line 6 and/or a cooled catalyst (not shown) into second reaction zone 7 for a second catalytic cracking reaction.
  • the reacted stream enters the outlet zone 8, which increases the line speed of the stream, allowing the reactant stream to quickly enter the settler 9 and cyclone 10 of the gas-solids separation system for separation.
  • the separated catalyst to be produced enters the stripper 12, is stripped, enters the riser 14 to be produced, and is lifted by the air from the line 17 into the regenerator 15 where it is charred and regenerated.
  • the flue gas is separated by the cyclone separator 13 and discharged to the regenerator via line 18, and the hot regenerated catalyst is returned to the bottom of the riser via the regeneration inclined tube 16 for recycling.
  • a portion of the catalyst to be produced in the settler can be delivered to the second reaction zone 7 as a chilling medium.
  • the separated reaction oil is sent to the fractionation system 19 via the line 11, and the dry gas obtained by fractional distillation is taken out through the line 20; the liquefied gas fraction is sent to a subsequent treatment device (not shown) via the line 21 to separate the isobutane;
  • the gasoline fraction preferably enters the extraction refining unit 45 via line 22, and the resulting benzene, toluene, xylene, and extract are sent from line 46, line 47, line 48, and line 49, respectively; and the light cycle oil fraction is withdrawn via line 23.
  • Hydrogen enters the hydrogen recycle compressor 32 via line 33 for pressurization and is then recycled to the hydrotreater 31 via line 30 and line 29 and line 28.
  • the hydrogenated tail oil enters the pre-elevation section 2 via line 41, line 4, and atomized steam from line 3.
  • the initial feedstock oil is treated by a hydrotreating unit to provide a hydrogenated tailstock which is fed to the variable riser reactor as a catalytic cracking feedstock oil.
  • the preheated hydrogenated tail oil enters the catalyst from the lower part of the first reaction zone of the reactor, and mainly undergoes a cracking reaction, and the mixture of the oil and gas and the catalyst formed after the reaction is ascended to the lower part of the second reaction zone of the reactor, optionally In contact with the catalyst after cooling, the hydrogen transfer reaction and the isomerization reaction are mainly carried out, and after the reaction, the effluent enters the settler to separate the reaction oil and the catalyst to be produced.
  • the reaction oil is separated to obtain a reaction product including a dry gas, a liquefied gas fraction, a gasoline fraction, a light cycle oil fraction, a heavy cycle oil fraction, and a slurry.
  • a reaction product including a dry gas, a liquefied gas fraction, a gasoline fraction, a light cycle oil fraction, a heavy cycle oil fraction, and a slurry.
  • the light cycle oil fraction, the heavy cycle oil fraction and/or the slurry, and optionally the distillate from the other treatment unit are sent to a hydrogenation unit for hydrotreating under the action of a hydrotreating catalyst.
  • the hydrogenated tail oil is returned to the riser reactor.
  • the spent catalyst is stripped and regenerated and returned to the riser reactor.
  • Fig. 2 schematically shows another exemplary flow for producing isobutane and aromatic-rich gasoline using a variable diameter riser reactor, wherein the shape, size and the like of the equipment and piping are not limited by the drawings, and It is determined on a case-by-case basis.
  • the pre-lifting steam enters from the pre-lift section 2 of the variable-diameter riser reactor via line 1, and the hot regenerated catalyst enters the pre-lift section 2 via the regeneration inclined tube 16 and is lifted by the pre-lifting steam.
  • the initial feedstock oil from line 27 e.g., distillate oil from other processing units
  • the hydrogen from line 28 and optionally the feed from line 26 is passed to a hydrotreating unit 31 for hydrotreating, and the hydrogenated product is subjected to a high pressure.
  • the separator 34 is separated from the low pressure separator 35, and the liquid product is separated from the line 36 through the hydrogenation unit fractionation column 37 to obtain gas, hydrogenated gasoline, hydrogenated light cycle oil and hydrogenated tail oil, respectively, via line 38, line 39, and line. 40 and line 41 are sent out.
  • Hydrogen enters the hydrogen recycle compressor 32 via line 33 for pressurization and is then recycled to the hydrotreater 31 via line 30 and line 29 and line 28.
  • the hydrogenation tail oil is used as the catalytic cracking feedstock oil from the pre-elevation section 2 through the pipeline 41, the pipeline 4 and the atomized steam from the pipeline 3, and is mixed with the hot regenerated catalyst and then enters the first reaction zone 5, where the first is carried out. Catalytic cracking reaction.
  • the reactant stream is mixed with a cold shock agent from line 6 and/or a cooled catalyst (not shown) into second reaction zone 7 for a second catalytic cracking reaction.
  • the reacted stream enters the outlet zone 8, which increases the line speed of the stream, allowing the reactant stream to quickly enter the settler 9 and cyclone 10 of the gas-solids separation system for separation.
  • the separated catalyst to be produced enters the stripper 12, is stripped, enters the riser 14 to be produced, and is lifted by the air from the line 17 into the regenerator 15 where it is charred and regenerated.
  • the flue gas is separated by the cyclone separator 13 and discharged to the regenerator via line 18, and the hot regenerated catalyst is returned to the bottom of the riser via the regeneration inclined tube 16 for recycling.
  • a portion of the catalyst to be produced in the settler can be delivered to the second reaction zone 7 as a chilling medium.
  • the separated reaction oil is sent to the fractionation system 19 via the line 11, and the dry gas obtained by fractional distillation is taken out through the line 20; the liquefied gas fraction is sent to a subsequent treatment device (not shown) via the line 21 to separate the isobutane;
  • the gasoline fraction preferably enters the extraction refining unit 45 via line 22, and the resulting benzene, toluene, xylene, and extract are sent from line 46, line 47, line 48, and line 49, respectively; and the light cycle oil fraction is withdrawn via line 23.
  • the initial feedstock oil is treated with an aromatics extraction unit to provide a polycyclic aromatic hydrocarbon-enriched extractive oil and a polycyclic naphthenic-rich convulsion oil.
  • the extracted oil is treated by a hydrotreating unit to obtain a hydrogenated tail oil, and the hydrogenated tail oil and the pumping oil are combined and sent to the variable diameter riser reactor as a catalytic cracking feedstock oil.
  • the preheated catalytic cracking feedstock oil enters the catalyst from the lower part of the first reaction zone of the reactor, and mainly undergoes a cracking reaction, and the mixture of the oil and gas and the catalyst formed after the reaction is ascended to the lower part of the second reaction zone of the reactor, optionally In contact with the catalyst after cooling, the hydrogen transfer reaction and the isomerization reaction are mainly carried out, and after the reaction, the effluent enters the settler to separate the reaction oil and the catalyst to be produced.
  • the reaction oil is separated to obtain a reaction product including a dry gas, a liquefied gas fraction, a gasoline fraction, a light cycle oil fraction, a heavy cycle oil fraction, and a slurry.
  • the light cycle oil fraction, the heavy cycle oil fraction and/or the slurry, and optionally the distillate from other processing units are sent to a hydrogenation unit for hydrotreating under the action of a hydrotreating catalyst.
  • the hydrogenated tail oil is returned to the riser reactor.
  • the spent catalyst is stripped and regenerated and returned to the riser reactor.
  • Fig. 3 schematically shows another exemplary flow for producing isobutane and aromatic-rich gasoline using a variable diameter riser reactor, wherein the shape, size, and the like of the equipment and piping are not limited by the drawings, and It is determined on a case-by-case basis.
  • the pre-lifting steam enters from the pre-lift section 2 of the variable-diameter riser reactor via line 1, and the hot regenerated catalyst enters the pre-lift section 2 via the regeneration inclined tube 16 and is lifted by the pre-lifting steam.
  • the initial feedstock oil from line 43 e.g., distillate from other processing units
  • the withdrawn oil is passed through line 27 with hydrogen from line 28 and optionally from line 26 to a hydrotreating unit 31 for hydrotreating, and the hydrogenated product is separated by high pressure separator 34 and low pressure separator 35.
  • the hydrogenation unit fractionation column 37 separates the gas, hydrogenated gasoline, hydrogenated light cycle oil and hydrogenated tail oil, which are sent through line 38, line 39, line 40 and line 41, respectively.
  • Hydrogen enters the hydrogen recycle compressor 32 via line 33 for pressurization and is then recycled to the hydrotreater 31 via line 30 and line 29 and line 28.
  • the hydrogenated tail oil is combined with the pumping oil from line 44 as a catalytic cracking feedstock oil, which is fed from the pre-elevation section 2 together with the atomized steam from line 3 via line 41, line 4, and mixed with the hot regenerated catalyst.
  • a reaction zone 5 a first catalytic cracking reaction is carried out therein.
  • the reactant stream is mixed with a cold shock agent from line 6 and/or a cooled catalyst (not shown) into second reaction zone 7 for a second catalytic cracking reaction.
  • the reacted stream enters the outlet zone 8, which increases the line speed of the stream, allowing the reactant stream to quickly enter the settler 9 and cyclone 10 of the gas-solids separation system for separation.
  • the separated catalyst to be produced enters the stripper 12, is stripped, enters the riser 14 to be produced, and is lifted by the air from the line 17 into the regenerator 15 where it is charred and regenerated.
  • the flue gas is separated by the cyclone separator 13 and discharged to the regenerator via line 18, and the hot regenerated catalyst is returned to the bottom of the riser via the regeneration inclined tube 16 for recycling.
  • a portion of the catalyst to be produced in the settler can be delivered to the second reaction zone 7 as a chilling medium.
  • the separated reaction oil is sent to the fractionation system 19 via the line 11, and the dry gas obtained by fractional distillation is taken out through the line 20; the liquefied gas fraction is sent to a subsequent treatment device (not shown) via the line 21 to separate the isobutane;
  • the gasoline fraction preferably enters the extraction refining unit 45 via line 22, and the resulting benzene, toluene, xylene, and extract are sent from line 46, line 47, line 48, and line 49, respectively; and the light cycle oil fraction is withdrawn via line 23.
  • the present application provides the following technical solutions:
  • a catalytic cracking process for producing isobutane and light aromatics comprising:
  • catalytic cracking feedstock oil feedstock oil into the catalytic cracking reactor, contacting the catalytic cracking catalyst, and sequentially performing the first catalytic cracking reaction and the second catalytic cracking reaction under different reaction conditions to obtain a catalytic cracking reaction product and a catalyst to be produced.
  • the catalytic cracking feedstock oil has a bicyclo or higher cycloalkane content of greater than about 25% by weight;
  • the light cycle oil fraction, the heavy cycle oil fraction and the slurry obtained in the step (2) are optionally fed to a hydrogenation unit to be contacted with a hydrotreating catalyst and selectively hydrotreated to obtain a hydrogenated tail oil. Feeding the catalytic cracking feedstock oil into the catalytic cracking reactor;
  • the gasoline obtained in the step (2) is subjected to extraction and purification to obtain a light aromatic hydrocarbon.
  • catalytic cracking feedstock oil is selected from the group consisting of deep hydrogenated light cycle oil, heavy distillate of delayed coking unit, catalytic cracking light cycle oil, catalytic cracking heavy cycle oil, catalytic cracking heavy fraction Oil, slurry, hydrocracked diesel, residue hydrocracking diesel, wax hydrocracking diesel, biodiesel, shale oil diesel fraction, coal liquefied diesel fraction, atmospheric top oil, atmospheric pressure tower extracted fraction
  • a secondary processing feedstock, and/or at least one of the primary processing distillate and/or secondary processing feedstock is hydrogenated tailings obtained by hydrogenation.
  • A5. The method according to item A1, wherein the conditions of the first catalytic cracking reaction comprise: a reaction temperature of about 520-620 ° C, a reaction time of about 0.5-3.0 seconds, and a weight ratio of the agent oil of about 3:1 to About 15:1;
  • the conditions of the second catalytic cracking reaction include a reaction temperature of about 420 to 530 ° C, a reaction time of about 2 to 30 seconds, and a weight ratio of the agent to the oil of from about 3:1 to about 18:1.
  • A6 The method according to item A1, wherein the conditions of the first catalytic cracking reaction comprise: a reaction temperature of about 530-600 ° C, a reaction time of about 0.8-2.0 seconds, and a weight ratio of the agent oil of about 4:1. Up to about 12:1;
  • the conditions of the second catalytic cracking reaction include a reaction temperature of about 460 to 510 ° C, a reaction time of about 3 to 15 seconds, and a weight ratio of the agent oil of from about 4:1 to about 15:1.
  • catalytic cracking reactor is an equal diameter riser, an equal line speed riser, a variable diameter riser, a fluidized bed, or an equal diameter riser and a fluidized bed.
  • Composite reactor is an equal diameter riser, an equal line speed riser, a variable diameter riser, a fluidized bed, or an equal diameter riser and a fluidized bed.
  • variable diameter riser is provided with a coaxial and fluidly connected pre-lift section, a first reaction zone, a second reaction zone and an exit zone in a vertical direction from bottom to top.
  • the end of the outlet zone is provided with a horizontal pipe for connecting the settler, the inner diameter of the first reaction zone is smaller than the second reaction zone, the inner diameter of the second reaction zone is larger than the outlet zone, and the total of the variable diameter riser a height of about 10-60 meters,
  • the catalytic cracking catalyst is sent to the pre-lifting section, the catalytic cracking feedstock oil is sent to a lower portion of the first reaction zone, and the first catalytic cracking reaction is in a first reaction In the zone, the second catalytic cracking reaction is carried out in the second reaction zone.
  • a heat extractor is disposed in the second reaction zone, and the height of the heat extractor is about 50-90% of the height of the second reaction zone.
  • chilling medium is selected from the group consisting of a cold blasting agent, a cooled regenerated catalyst, a cooled semi-regenerated catalyst, and a fresh catalyst, and combinations thereof, the chilling agent being selected from the group consisting of liquefaction Gas, crude gasoline, stabilized gasoline, light cycle oil fraction, heavy cycle oil fraction and water, and combinations thereof.
  • the catalytic cracking catalyst comprises a cracking active component and a carrier; and the cracking active component comprises from about 0 to 100% by weight, preferably from about 10 to 90% by weight on a dry basis.
  • the FAU type zeolite is selected from the group consisting of Y type zeolite, HY type zeolite and ultra stable Y type zeolite, and their
  • the five-membered ring structure zeolite is selected from the group consisting of ZSM-5 series zeolites, high silica zeolites and ferrierites, and combinations thereof, the five-membered ring structure zeolites with or without rare earths and with or without phosphorus.
  • the hydrotreating catalyst comprises a hydrogenation active component and a carrier selected from the group VIB non-noble metals, Group VIII non-noble metals, And combinations thereof, the support is selected from the group consisting of alumina, silica, and amorphous silica-alumina, and combinations thereof.
  • extracting the solvent is selected from the group consisting of sulfolane, dimethyl sulfoxide, N-formylmorpholine, tetraethylene glycol, and three Glycol and N-methylpyridinone, and combinations thereof, have a temperature of from about 50 to about 110 ° C and a weight ratio of extraction solvent to gasoline of from about 2 to about 6.
  • aromatic hydrocarbon extraction conditions include: a temperature of about 50-70 ° C, a solvent to raw material weight ratio of about 0.5-2, and a solvent selected from the group consisting of furfural, dimethyl sulfoxide, and Methylformamide, monoethanolamine, ethylene glycol, and 1,2-propanediol, and combinations thereof.
  • A16 The method according to item A1 or A4, wherein the conditions of the selective hydrotreating comprise: a hydrogen partial pressure of about 6.0-30.0 MPa, a reaction temperature of about 300-450 ° C, and a liquid phase volume space velocity of about 0.1-10.0 h -1 , the hydrogen oil volume ratio is about 300-3000 Nm 3 /m 3 ; preferably, the hydrogen partial pressure is about 8-20 MPa, the reaction temperature is about 330-430 ° C, and the liquid hour volume space velocity is about 0.2. -5h -1, hydrogen to oil volume ratio of about 500-2500Nm 3 / m 3.
  • a catalytic cracking method comprising the steps of:
  • step c) further results in a light cycle oil fraction, a heavy cycle oil fraction, and optionally a slurry, and the method further comprises the steps of:
  • step c) hydrotreating at least a portion of the light cycle oil fraction, the heavy cycle oil fraction and the optional slurry obtained in step c) to obtain a hydrogenated tail oil;
  • step a) further comprises pretreating the starting stock oil having a naphthene-containing naphthenic content of no greater than about 25% by weight to obtain a naphthene-containing naphthenic content greater than about 25
  • the catalytically cracked feedstock oil is % by weight, preferably greater than about 40% by weight.
  • the hydrogen treatment catalyst comprises a hydrogenation active component and a support, preferably selected from the group consisting of Group VIB non-noble metals, Group VIII non-noble metals, and combinations thereof, preferably selected from the group consisting of alumina, dioxide. Silicon, amorphous silicon aluminum, and combinations thereof.
  • catalytic cracking feedstock oil or the initial feedstock oil is selected from the group consisting of deep hydrogenated light cycle oil, heavy distillate of a delayed coker unit, catalytic cracking light cycle oil, Catalytic cracking heavy cycle oil, catalytic cracking heavy distillate, oil slurry, hydrocracked diesel, residual hydrocracking diesel, wax hydrocracking diesel, biodiesel, shale oil diesel fraction, coal liquefied diesel fraction, atmospheric pressure Tower top oil, distillate oil extracted from atmospheric tower, straight-run depressurized wax oil, hydrogenated wax oil, coking wax oil, deasphalted oil, extracted oil, pumping oil, atmospheric residue, vacuum residue, above A hydrogenated tailstock obtained by hydrogenation of a feedstock oil, and combinations thereof.
  • the first catalytic cracking reaction is carried out under the following conditions: a reaction temperature of about 520-620 ° C; a reaction time of about 0.5-3.0 seconds; a weight ratio of the agent oil of about 3:1 to about 15:1;
  • the second catalytic cracking reaction is carried out under the following conditions: a reaction temperature of about 480 to 600 ° C; a reaction time of about 2 to 30 seconds; and a ratio of the agent to oil of from about 3:1 to about 18:1.
  • the first catalytic cracking reaction is carried out under the following conditions: a reaction temperature of about 530-600 ° C; a reaction time of about 0.8-2.0 seconds; a weight ratio of the agent oil of about 4:1 to about 12:1;
  • the second catalytic cracking reaction is carried out under the following conditions: a reaction temperature of about 500 to 550 ° C; a reaction time of about 3 to 15 seconds; and a weight ratio of the agent to oil of from about 4:1 to about 15:1.
  • the catalytic cracking reactor is an equal diameter riser reactor, an equal line riser reactor, a variable diameter riser reactor, a fluidized bed reactor, Or a composite reactor consisting of a diameter riser and a fluidized bed.
  • the catalytic cracking reactor is a variable diameter riser reactor, and the variable diameter riser reactor is coaxially arranged in a vertical direction from bottom to top.
  • a pre-elevation section a first reaction zone, a second reaction zone, and an outlet zone in fluid communication, wherein an inner diameter of the first reaction zone is smaller than a second reaction zone, an inner diameter of the second reaction zone is larger than an outlet zone, and the catalytic cracking Catalyst is fed into the pre-elevation section, the catalytic cracking feedstock oil is fed to a lower portion of the first reaction zone, the first catalytic cracking reaction is carried out in a first reaction zone, and the second catalytic cracking reaction is The second reaction zone is carried out.
  • a heat extractor is disposed in the second reaction zone, and the height of the heat extractor is about 50-90% of the height of the second reaction zone.
  • chilling medium is selected from the group consisting of a cold shock, a cooled regenerated catalyst, a cooled semi-regenerated catalyst, a fresh catalyst, and combinations thereof;
  • cold shock agent is selected from the group consisting of liquefied gas , crude gasoline, stabilized gasoline, light cycle oil, heavy cycle oil, water, and combinations thereof.
  • the catalytic cracking catalyst comprises a cracking active component and a carrier; the cracking active group is based on the dry basis and based on the weight of the cracking active component
  • the fraction comprises about 0-100% by weight of a FAU-type zeolite and about 0-100% by weight of a five-membered ring-structured zeolite, wherein the total amount of the FAU-type zeolite and the five-membered ring-structured zeolite is 100% by weight.
  • the catalytic cracking catalyst comprises a cracking active component and a carrier; the cracking active group is based on the dry basis and based on the weight of the cracking active component
  • the fraction comprises from about 10 to 90% by weight of the FAU type zeolite and from about 10 to 90% by weight of the five-membered ring structure zeolite, wherein the total amount of the FAU type zeolite and the five-membered ring structure zeolite is 100% by weight.
  • the FAU type zeolite is selected from the group consisting of Y type zeolite, HY type zeolite, ultra stable Y type zeolite, and combinations thereof; and the five-membered ring structure zeolite is selected from ZSM- A series 5 zeolite, a high silica zeolite, a ferrierite, and combinations thereof, the five-membered ring structure zeolite optionally containing rare earth and/or phosphorus.
  • the extraction solvent is selected from the group consisting of sulfolane, dimethyl sulfoxide, N-formylmorpholine, tetraethylene glycol, triethylene glycol, N - Methylpyridinone, and combinations thereof, at a temperature of from about 50 to about 110 ° C, and a weight ratio of extraction solvent to gasoline fraction of from about 2 to about 6.
  • the catalytic cracking catalyst IBA-1 used in the following examples and comparative examples used high silica Y zeolite and ZRP zeolite as cracking active components, and the weight ratio of the two was 30% by weight of high silica Y zeolite and 70% by weight of ZRP zeolite.
  • the specific preparation method of the catalyst is as follows:
  • catalytic cracking catalysts of the commercial grades CGP-1, MLC-500 and DMMC-1 used in the following examples and comparative examples were all produced by Qilu Catalyst Plant of Sinopec Catalyst Branch, and their physical and chemical properties are listed in Table 3.
  • hydrotreating catalysts of commercial grades RN-32V used in the following examples and comparative examples, and the protective agents of commercial grades RG-30A, RG-30B, RG-1, are all produced by Sinopec Catalyst Branch, plus The loading ratio of the hydrotreating catalyst and the protecting agent in the hydrogen apparatus was 95:5.
  • micro-reaction activity (MAT) of catalytic cracking catalysts is based on the standard method of RIPP 92-90 (see “Petrochemical Analysis Methods (RIPP Test Methods)", edited by Yang Cuiding, Science Press, September 1990, first edition, 263 - 268 pages)
  • the measurement conditions were as follows: catalyst 5.0 g; oil feed amount 1.56 g; reaction time 70 sec; reaction temperature 460 ° C; agent/oil 3.2; space velocity 16 h -1 .
  • the specific surface area and total pore volume of the catalytic cracking catalyst were measured by AS-3, AS-6 static nitrogen adsorption meter manufactured by Quantachrome Instruments, and the specific measurement method was as follows: the sample was placed in a sample processing system, and vacuum was taken at 300 ° C.
  • the preheated catalytic cracking feedstock A is contacted with the catalyst IBA-1 and subjected to a catalytic cracking reaction on a medium-sized variable-diameter riser catalytic cracking unit.
  • the preheated feedstock oil A enters the variable diameter riser reactor as shown in FIG. 1 and is first contacted with the hot catalytic cracking catalyst IBA-1 in the first reaction zone and the second reaction zone in the presence of water vapor.
  • Catalytic cracking reaction and second catalytic cracking reaction separating catalytic cracking reaction product to obtain dry gas, isobutane-rich liquefied gas, aromatic-rich gasoline, light cycle oil and heavy cycle oil (no oil slurry); light cycle oil
  • the outer pumping device recycles the heavy cycle oil to the hydrogenation unit for hydrotreating, and the obtained hydrogenated tail oil A' is recycled back to the riser reactor for catalytic cracking reaction; the catalyst to be produced is stripped and then enters the regenerator and is burned. The focus is recycled after regeneration.
  • the properties of the feedstock oil are listed in Table 1, the operating conditions are listed in Table 4, the product distribution is listed in Table 5, and the properties of the gasoline obtained by extraction and purification are shown in Table 5.
  • Example 2 The operation was carried out in a similar manner to that described in Example 1, in which a conventional catalytic cracking catalyst CGP-1 was employed.
  • the preheated catalytic cracking feedstock oil A enters the variable diameter riser reactor as shown in FIG. 1 and is contacted with the hot catalytic cracking catalyst CGP-1 in the first reaction zone and the second reaction zone in the presence of water vapor.
  • First catalytic cracking reaction and second catalytic cracking reaction separating catalytic cracking reaction product to obtain dry gas, isobutane-rich liquefied gas, aromatic hydrocarbon-rich gasoline, light cycle oil and heavy cycle oil (no oil slurry);
  • the circulating oil external scooping device recirculates the oil to the hydrogenation unit for hydrotreating, and the obtained hydrogenated tail oil is recycled to the riser reactor for catalytic cracking reaction; the catalyst to be produced is stripped and then enters the regenerator and is burned. The focus is recycled after regeneration.
  • the properties of the feedstock oil are listed in Table 1, the operating conditions are listed in Table 4, the product distribution is listed in Table 5, and the properties of the gasoline obtained by extraction and purification are shown in Table 5.
  • Example 2 The operation was carried out in a similar manner to that described in Example 1, in which a conventional catalytic cracking catalyst MLC-500 was employed.
  • the preheated catalytic cracking feedstock oil A enters a conventional equal-diameter riser reactor, and a chill medium is injected in the middle of the equal-diameter riser reactor to form a first reaction zone at the lower portion and a second reaction at the upper portion. Area.
  • the feedstock oil A is contacted with the hot catalytic cracking catalyst MLC-500 in the first reaction zone and the second reaction zone in the presence of steam to carry out the first catalytic cracking reaction and the second catalytic cracking reaction, and the catalytic cracking reaction product is separated.
  • Example 2 The operation was carried out in a similar manner to that described in Example 1, in which a conventional catalytic cracking catalyst DMMC-1 was employed.
  • the preheated catalytic cracking feedstock oil A enters a riser reactor + fluidized bed composite reactor wherein the riser reactor acts as a first reaction zone and the fluidized bed acts as a second reaction zone.
  • the feedstock oil A is contacted with the hot catalytic cracking catalyst DMMC-1 in the first reaction zone and the second reaction zone in the presence of steam to carry out the first catalytic cracking reaction and the second catalytic cracking reaction, and the catalytic cracking reaction product is separated.
  • Example 2 The operation was carried out in a similar manner to that described in Example 1, in which a feedstock oil D having a low polycycloalkane content was used.
  • the preheated feedstock oil D enters the variable diameter riser reactor as shown in Fig. 1, and reacts with the hot catalytic cracking catalyst CGP-1 in the presence of water vapor to separate the catalytic cracking reaction product to obtain dry gas and liquefied gas.
  • Gasoline, light cycle oil, heavy cycle oil and oil slurry; light cycle oil and oil slurry are all externally pumped out, heavy cycle oil is recycled to the hydrogenation unit for hydrotreating and then returned to the riser reactor, and the catalyst is to be produced. After being stripped, it enters the regenerator, and is burned and recycled for recycling.
  • the properties of the feedstock oil are listed in Table 1, the operating conditions are listed in Table 4, the product distribution is listed in Table 5, and the properties of the gasoline obtained by extraction and purification are shown in Table 5.
  • Example 2 The operation was carried out in a similar manner to that described in Example 1, in which a conventional catalytic cracking catalyst MLC-500 was employed.
  • the preheated feedstock oil A enters a conventional equal-diameter riser reactor, reacts with a hot catalytic cracking catalyst in the presence of steam, and separates the catalytic cracking reaction product to obtain dry gas, liquefied gas, gasoline, light cycle oil, Heavy cycle oil and oil slurry; light cycle oil and oil slurry are both externally pumped out, heavy cycle oil is recycled to the hydrogenation unit for hydrotreating, and then returned to the riser reactor, and the catalyst is stripped and then enters the regenerator , burnt and recycled after recycling.
  • the properties of the feedstock oil are listed in Table 1, the operating conditions are listed in Table 4, the product distribution is listed in Table 5, and the properties of the gasoline obtained by extraction and purification are shown in Table 5.
  • Example 2 The operation was carried out in a similar manner to that described in Example 1, in which a catalytic cracking catalyst IBA-1 was employed.
  • the preheated feedstock oil A enters a conventional equal-diameter riser reactor, reacts with a hot catalytic cracking catalyst in the presence of steam, and separates the catalytic cracking reaction product to obtain dry gas, liquefied gas, gasoline, light cycle oil, Heavy cycle oil and oil slurry; light cycle oil and oil slurry are both externally pumped out, heavy cycle oil is recycled to the hydrogenation unit for hydrotreating, and then returned to the riser reactor, and the catalyst is stripped and then enters the regenerator , burnt and recycled after recycling.
  • the properties of the feedstock oil are listed in Table 1, the operating conditions are listed in Table 4, the product distribution is listed in Table 5, and the properties of the gasoline obtained by extraction and purification are shown in Table 5.
  • the initial feedstock oil B is first subjected to hydrogenation to obtain a catalytic cracking feedstock oil C.
  • the preheated catalytic cracking feedstock oil C enters the variable diameter riser reactor as shown in FIG.
  • the feedstock oil B is not hydrogenated, and the preheated feedstock oil B is directly fed into the variable diameter riser reactor as shown in Fig. 1, in the presence of water vapor, with the hot catalyst.
  • IBA-1 contact reaction separation of catalytic cracking reaction products to obtain dry gas, liquefied gas, gasoline, light cycle oil, heavy cycle oil and slurry; light cycle oil and slurry external extraction device, heavy cycle oil circulation to hydrogenation unit
  • the reactor is returned to the riser reactor; the catalyst to be produced is stripped and then enters the regenerator, and is burned and recycled for recycling.
  • the properties of the feedstock oil are listed in Table 1, the operating conditions are listed in Table 6, the product distribution is listed in Table 7, and the properties of the gasoline obtained by extraction and purification are shown in Table 7.
  • the raw material oil vacuum distillate E is treated by an aromatic hydrocarbon extraction device to obtain a vacuum distillate oil extraction oil G and a vacuum distillate oil extraction oil F.
  • the vacuum distillate extracts the oil G and the hydrogen mixture, and then enters the hydrogenation unit for hydrotreating, and obtains the hydrogenated tail oil G′ from the hydrogenation unit, and preheats the hydrogenated tail oil and the vacuum distillate oil to extract the oil F. After mixing, it is used as catalytic cracking feedstock oil to enter the variable-diameter riser reactor shown in Fig. 3.
  • the catalytic cracking reaction product In contact with hot catalyst IBA-1 in the presence of water vapor, the catalytic cracking reaction product is separated to obtain dry gas and rich in Butane liquefied gas, aromatic hydrocarbon-rich gasoline, light cycle oil and heavy cycle oil (no oil slurry); light cycle oil extraction device, heavy cycle oil and vacuum distillate oil extraction oil G mixed and sent to hydrogenation
  • the device is subjected to hydrotreating, and the obtained hydrogenated tail oil is recycled to the riser reactor for catalytic cracking reaction; the catalyst to be produced is stripped and then enters the regenerator, and is recycled after scorch regeneration.
  • the properties of the feedstock oil are listed in Table 2, the operating parameters are listed in Table 8, the product distribution is listed in Table 9, and the properties of the gasoline obtained by extraction and purification are shown in Table 9.
  • the feedstock oil E is directly fed into the reduced-diameter riser reactor as shown in Figure 1 without extraction and hydrotreating, in the presence of water vapor.
  • the hot catalyst IBA-1 separating the catalytic cracking reaction product to obtain dry gas, liquefied gas, gasoline, light cycle oil, heavy cycle oil and oil slurry; among them, light circulation oil, oil slurry external extraction device, heavy circulation
  • the oil is recycled to the hydrogenation unit for hydrotreating and then returned to the riser reactor; the catalyst to be produced is stripped and then enters the regenerator, and is recycled after scorch regeneration.
  • the properties of the raw materials are listed in Table 2, the operating parameters are listed in Table 8, the product distribution is listed in Table 9, and the properties of the products obtained by extracting and purifying gasoline are shown in Table 9.
  • the raw material oil vacuum residue H is treated by an aromatic hydrocarbon extraction device to obtain a vacuum residue oil J and a vacuum residue oil I, and the aromatics extraction device conditions: temperature 60 °C, the solvent to raw material weight ratio is 1.5, and the solvent is furfural.
  • the vacuum residue extracts the oil J and the hydrogen mixture, and then enters the hydrogenation unit for hydrotreating, and obtains the hydrogenated tail oil J' from the hydrogenation unit, and preheats the hydrogenated tail oil and the vacuum residue oil to extract the oil I After mixing, it is used as catalytic cracking feedstock oil to enter the variable-diameter riser reactor shown in Fig. 3.
  • the catalytic cracking reaction product In contact with hot catalyst IBA-1 in the presence of water vapor, the catalytic cracking reaction product is separated to obtain dry gas and rich in Butane liquefied gas, aromatic-rich gasoline, light cycle oil and heavy cycle oil (no oil slurry), light cycle oil extraction device, heavy cycle oil and vacuum residue oil J mixed and sent to hydrogenation
  • the device is subjected to hydrotreating, and the obtained hydrogenated tail oil is recycled to the riser reactor for catalytic cracking reaction; the catalyst to be produced is stripped and then enters the regenerator, and is recycled after scorch regeneration.
  • the properties of the feedstock oil are listed in Table 2, the operating parameters are listed in Table 8, the product distribution is listed in Table 9, and the properties of the gasoline obtained by extraction and purification are shown in Table 9.
  • the feedstock oil vacuum residue H is not extracted and hydrotreated, and the preheated feedstock H is directly fed into the variable diameter riser reactor as shown in Figure 1.
  • the hot catalyst IBA-1 reacts with the hot catalyst IBA-1 to separate the catalytic cracking reaction product to obtain dry gas, liquefied gas, gasoline, light cycle oil, heavy cycle oil and oil slurry; among them, light cycle oil and oil slurry are discharged.
  • the device recycles the heavy cycle oil to the hydrogenation unit and returns to the riser reactor after being hydrotreated; the catalyst to be produced is stripped and then enters the regenerator, and is recycled after being burned.
  • the properties of the feedstock oil are listed in Table 2, the operating parameters are listed in Table 8, the product distribution is listed in Table 9, and the properties of the gasoline obtained by extraction and purification are shown in Table 9.
  • the process of the present invention is capable of producing isobutane and/or light aromatics, particularly in the case of a variable diameter riser reactor.

Abstract

一种多产异丁烷和/或轻质芳烃的催化裂化方法，包括如下步骤：a)提供催化裂化原料油，催化裂化原料油中双环以上环烷烃的含量大于25重量%；b)使催化裂化原料油在不同的反应条件下依次进行第一催化裂化反应和第二催化裂化反应，得到催化裂化反应产物；c)分离所得的催化裂化反应产物，得到包含异丁烷的液化气馏分和包含轻质芳烃的汽油馏分；以及d)从液化气馏分中回收异丁烷，和/或从汽油馏分中回收轻质芳烃。

Description

一种多产异丁烷和/或轻质芳烃的催化裂化方法
相关申请的交叉引用
本申请要求申请人于2017年10月26日向中国专利局提交的申请号为201711022235.0、名称为“一种多产异丁烷和轻质芳烃的催化裂化方法”的专利申请的优先权，该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技术领域，具体涉及一种多产异丁烷和/或轻质芳烃的催化裂化方法。
背景技术
芳烃中的苯、甲苯、二甲苯(简称BTX)是重要的化工原料，苯可以用来合成苯乙烯、苯酚和苯胺等产品，甲苯可以作为有机合成的优良溶剂，而且是合成甲酚的理想原料；二甲苯中的邻二甲苯、间二甲苯和对二甲苯是有机合成的基础原料。异丁烷作为重要的化工原料，可用于与C 3-C 5烯烃发生烷基化反应制取烷基化油；与丙烯发生共氧化反应生产环氧丙烷(PO)，并联产叔丁醇(TBA)或者生产甲基叔丁基醚(MTBE)；或者通过异丁烷脱氢反应制备异丁烯。随着化工行业的发展，市场对于BTX和异丁烷的需求日益增加，如何扩展其来源是目前化工技术发展的目标之一。
常规的催化裂化工艺主要用于生产汽油，并已取得汽油产率高达50重量％以上的突破，满足了汽油无铅化的要求，提高了汽油的辛烷值。但是，无论是通过改变工艺条件，还是通过使用新型的沸石催化剂来提高汽油辛烷值，都是以提高汽油组成中的烯烃含量来实现增加汽油的辛烷值。目前，催化裂化汽油组成中的烯烃含量可达35-65重量％，这与中国国家标准对烯烃含量的要求相差甚远。液化气组成中烯烃含量更高，大约在70重量％，其中丁烯是异丁烷的数倍，难以作为烷基化原料。
美国专利US 5,154,818A公开了一种采用多种原料催化裂化生产较多高辛烷值汽油的方法，其中轻质烃类原料与待生催化剂在常规提 升管反应器的第一反应区接触发生芳构化反应和低聚反应，生成的油气和重质烃类原料进入第二反应区，与再生催化剂接触发生裂化反应，生成的油气和待生催化剂在沉降器中分离，油气去分离系统分离，汽提后的待生催化剂一部分返回第一反应区，另一部分则进入再生器进行烧焦再生，热的再生催化剂返回第二反应区循环使用。
美国专利US 5,685,972A公开了一种从催化裂化石脑油制取BTX的方法。该方法使用的原料为催化裂化石脑油和焦化石脑油，使用的催化剂优选ZSM-5或者带有加氢功能组分的催化剂。
中国专利申请公开CN104560166A公开了一种催化转化方法，其中催化裂化轻循环油经切割得到轻馏分和重馏分、重馏分经加氢处理得到加氢重馏分，轻馏分和加氢重馏分分开进入到催化裂化装置的副提升管反应器、重质石油烃进入催化裂化装置的主提升管反应器。该方法能够最大限度地优化满足轻循环油的不同馏分的催化裂化反应所需的苛刻条件，从而最大限度地生产高辛烷值的催化裂化汽油。
中国专利申请公开CN1232069A公开了一种制取异丁烷和富含异构烷烃汽油的催化转化方法，其重油加工能力提高、干气和油浆产率下降、汽油烯烃和硫含量大幅降低、装置能耗进一步减小。
尽管如此，本领域仍然需要一种能够进一步提高异丁烷和/或轻质芳烃产率的催化裂化方法。
发明内容
本申请的目的是提供一种新型的催化裂化方法，其可以多产异丁烷和/或轻质芳烃。
为了实现上述目的，本申请提供了一种催化裂化方法，包括如下步骤：
a)提供催化裂化原料油，其中以所述催化裂化原料油的重量计，所述催化裂化原料油中双环以上环烷烃的含量大于约25重量％；
b)使所述催化裂化原料油在催化裂化反应器中与催化裂化催化剂接触，并在不同的反应条件下依次进行第一催化裂化反应和第二催化裂化反应，得到催化裂化反应产物；
c)分离所得的催化裂化反应产物，得到包含异丁烷的液化气馏分和包含轻质芳烃的汽油馏分；以及
d)任选地，从所述液化气馏分中回收异丁烷，和/或从所述汽油馏分中回收轻质芳烃。
在优选实施方式中，步骤c)中的所述分离还得到轻循环油馏分、重循环油馏分，和任选的油浆，并且所述方法进一步包括如下步骤：
e)对步骤c)中得到的轻循环油馏分、重循环油馏分和任选的油浆的至少一部分进行加氢处理得到加氢尾油；以及
f)将所得的加氢尾油的至少一部分再循环到所述催化裂化反应器中。
在某些优选实施方式中，步骤a)进一步包括对双环以上环烷烃含量不大于约25重量％的初始原料油进行预处理，以得到所述双环以上环烷烃含量大于约25重量％的催化裂化原料油。
进一步优选地，所述预处理包括芳烃抽提和/或加氢处理。
在优选实施方式中，所述第一催化裂化反应在如下条件下进行：反应温度为约520-620℃，反应时间为约0.5-3.0秒，剂油重量比为约3∶1至约15∶1；并且，所述第二催化裂化反应在如下条件下进行：反应温度为约480-600℃，反应时间为约2-30秒，剂油重量比为约3∶1至约18∶1。
在优选实施方式中，所述加氢处理在如下条件下进行：氢分压为约6.0-30.0MPa，反应温度为约300-450℃，液相体积空速为约0.1-10.0h -1，氢油体积比为约300-3000Nm 3/m 3。
本申请的催化裂化方法可以提供一个或多个以下的优点：
1、充分利用富含多环环烷烃的原料油或者可以通过预处理获得高多环环烷烃含量的原料油，实现对此类原料油的高效利用；
2、产物液化气中异丁烷含量得到有效提升；以及
3、汽油产率大于40重量％，产物汽油中轻质芳烃含量提高。
本发明的其他特征和优点将在随后的具体实施方式部分予以详细说明。
附图说明
附图是用来帮助对本发明的进一步理解，并且构成说明书的一部分，与下面的具体实施方式一起用于解释本发明，但并不构成对本发明的限制。在附图中：
图1是本申请所公开方法的一种优选实施方式的流程示意图。
图2是本申请所公开方法的另一优选实施方式的流程示意图。
图3是本申请所公开方法的又一优选实施方式的流程示意图。
附图标记说明
1管线              2预提升段          3管线
4管线              5第一反应区        6管线
7第二反应区        8出口区            9沉降器
10旋风分离器       11管线             12汽提器
13旋风分离器       14待生立管         15再生器
16再生斜管         17管线             18管线
19分馏系统         20管线             21管线
22管线             23管线             24管线
25管线             26管线             27管线
28管线             29管线             30管线
31加氢装置         32循环压缩机       33管线
34高压分离器       35低压分离器       36管线
37加氢装置分馏塔   38管线             39管线
40管线             41管线             42芳烃抽提装置
43管线             44管线             45抽提精制装置
46管线             47管线             48管线
49管线             50管线             51管线
52管线
具体实施方式
以下将结合附图通过具体的实施方式对本发明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发明，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。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。
如无特殊说明，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，术语“双环以上环烷烃”和“多环环烷烃”可以互换使用，指具有两个以上的碳环的环烷烃。
在本文中，术语“双环以上芳烃”和“多环芳烃”可以互换使用，指具有两个以上的芳环的芳族化合物。
在本文中，术语“水油比”指雾化蒸汽与原料油的质量比。
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。
如上所述，本申请提供了一种催化裂化方法，其包括如下步骤：
a)提供催化裂化原料油，其中以所述催化裂化原料油的重量计，所述催化裂化原料油中双环以上环烷烃的含量大于约25重量％；
b)使所述催化裂化原料油在催化裂化反应器中与催化裂化催化剂接触，并在不同的反应条件下依次进行第一催化裂化反应和第二催化裂化反应，得到催化裂化反应产物；
c)分离所得的催化裂化反应产物，得到包含异丁烷的液化气馏分和包含轻质芳烃的汽油馏分；以及
d)任选地，从所述液化气馏分中回收异丁烷，和/或从所述汽油馏分中回收轻质芳烃。
在某些优选的实施方式中，所述催化裂化原料油中双环以上环烷烃的含量大于约40重量％，且含量越高越好。
在某些优选的实施方式中，步骤c)中的所述分离还得到轻循环油馏分、重循环油馏分，和任选的油浆，并且所述方法进一步包括如下步骤：
e)对步骤c)中得到的轻循环油馏分、重循环油馏分和任选的油浆的至少一部分进行加氢处理得到加氢尾油；以及
f)将所得的加氢尾油的至少一部分再循环到所述催化裂化反应器中。
在这类优选实施方式中，步骤e)中的所述加氢处理用于产生富含双环以上环烷烃的加氢尾油(或称为加氢馏分油)。优选地，所述加氢处理可以在如下条件下进行：氢分压为约6.0-30.0MPa，优选为约8-20MPa，反应温度为约300-450℃，优选为约330-430℃，液相体积空速为约0.1-10.0h -1，优选为约0.2-5h -1，氢油体积比为约300-3000Nm 3/m 3，优选约500-2500Nm 3/m 3。例如，所述加氢处理可以在如下条件下进行：氢分压为约8-20MPa，反应温度为约330-430℃，液相体积空速为约0.2-5h -1，氢油体积比为约500-2500Nm 3/m 3。
进一步优选地，所述加氢处理可以在加氢处理催化剂存在下进行，该加氢处理催化剂可以包括加氢活性组分和载体，所述加氢活性组分可以选自第VIB族非贵金属、第VIII族非贵金属，和它们的组合，所述载体可以选自氧化铝、二氧化硅和无定型硅铝，和它们的组合。优选地，以氧化物计并以所述加氢处理催化剂的重量为基准，所述第VIII族非贵金属的含量为约1-99重量％，优选约1-60重量％；和/或，所述第VIB族非贵金属的含量为约1-99重量％，优选约1-70重量％。优选地，所述第VIII族非贵金属为钴和/或镍，所述第VIB族非贵金属为钼和/或钨。所述的加氢处理催化剂具有优良的芳烃加氢饱和能力，可以有效地将轻循环油馏分、重循环油馏分和油浆中的芳烃加氢饱和为环烷烃，为本发明的催化裂化反应提供原料。
在这类优选实施方式中，催化裂化反应所产生的油浆可以任选地先进入油浆过滤系统，然后进入加氢装置进行加氢饱和，并至少部分再循环至所述催化裂化反应器，或者所产生的油浆可以直接外甩出装置。
在某些优选的实施方式中，步骤a)进一步包括对双环以上环烷烃含量不大于约25重量％的初始原料油进行预处理，以得到所述双环以上环烷烃含量大于约25重量％、优选大于约40重量％的催化裂化原料油。
例如，所述双环以上环烷烃含量不大于约25重量％的初始原料油 可以是双环以上环烷烃与双环以上芳烃的总含量大于约25重量％的原料油。此时，可以对所述初始原料油进行加氢处理，使所述双环以上芳烃加氢饱和为双环以上环烷烃，从而得到所述双环以上环烷烃含量大于约25重量％的催化裂化原料油。
优选地，所述加氢处理可以在如下条件下进行：氢分压为约6.0-30.0MPa，优选为约8-20MPa，反应温度为约300-450℃，优选为约330-430℃，液相体积空速为约0.1-10.0h -1，优选为约0.2-5h -1，氢油体积比为约300-3000Nm 3/m 3，优选约500-2500Nm 3/m 3。
进一步优选地，所述加氢处理可以在加氢处理催化剂存在下进行，该加氢处理催化剂可以包括加氢活性组分和载体，所述加氢活性组分可以选自第VIB族非贵金属、第VIII族非贵金属，和它们的组合，所述载体可以选自氧化铝、二氧化硅和无定型硅铝，和它们的组合。优选地，以氧化物计并以所述加氢处理催化剂的重量为基准，所述第VIII族非贵金属的含量为约1-99重量％，优选约1-60重量％；和/或，所述第VIB族非贵金属的含量为约1-99重量％，优选约1-70重量％。优选地，所述第VIII族非贵金属为钴和/或镍，所述第VIB族非贵金属为钼和/或钨。
或者，可以对双环以上环烷烃含量不大于约25重量％的初始原料油进行芳烃抽提，然后任选进行上述加氢处理。例如，通过芳烃抽提将所述初始原料油分为富含多环环烷烃的抽佘油和富含多环芳烃的抽出油；所述抽出油任选地再通过加氢精制使其中的多环芳烃饱和，形成多环环烷烃。
在此类实施方式中，所述芳烃抽提可以按本领域技术人员熟知的方式进行。优选地，所述芳烃抽提可以在如下条件下进行：温度为约50-70℃，溶剂与原料重量比为约0.5-2，溶剂选自糠醛、二甲亚砜、二甲基甲酰胺、单乙醇胺、乙二醇和1，2-丙二醇，和它们的组合。
适用于本申请所公开方法的催化裂化原料油可以是任何双环以上环烷烃含量大于约25重量％、优选大于约40重量％的原料油；适用于本申请所公开方法的初始原料油可以是任何经过所述预处理后可以达到双环以上环烷烃含量大于约25重量％、优选大于约40重量％的原料油。例如，所述催化裂化原料油或初始原料油可以选自深度加氢轻循环油、延迟焦化装置的重馏分油(CGO)、催化裂化轻循环油(LCO)、 催化裂化重循环油(HCO)、催化裂化重馏分油(FGO)、油浆、加氢裂化柴油、渣油加氢裂化柴油、蜡油加氢裂化柴油、生物柴油、页岩油柴油馏分、煤液化柴油馏分、常压塔顶油、常压塔抽出的馏分油、直馏减压蜡油、加氢蜡油、焦化蜡油、脱沥青油(DAO)、抽出油、抽佘油、常压渣油、减压渣油、上述原料油经加氢得到的加氢尾油，和它们的组合。
在本申请所公开的方法中，所述第一和第二催化裂化反应可以按本领域技术人员所熟知的方式，在不同的反应条件下进行，其中所述第一催化裂化反应主要包括裂化反应，所述第二催化裂化反应主要包括选择性氢转移反应、异构化反应和芳构化反应。优选地，所述第一催化裂化反应可以在如下条件下进行：反应温度为约520-620℃，优选约530-600℃；反应时间为约0.5-3.0秒，优选约0.8-2.0秒；剂油重量比为约3∶1至约15∶1，优选约4∶1至约12∶1；水油比为约0.03∶1至约0.3∶1，优选为约0.05∶1至约0.3∶1；压力为130-450kPa；和/或，所述第二催化裂化反应可以在如下条件下进行：反应温度为约480-600℃，优选约500-550℃，或者约420-530℃，优选约460-510℃；反应时间为约2-30秒，优选约3-15秒；剂油重量比为约3∶1至约18∶1，优选约4∶1至约15∶1；水油比为约0.03∶1至约0.3∶1，优选为约0.05∶1至约0.3∶1；压力为130-450kPa。
适用于本申请所公开方法的催化裂化反应器可以是任何本领域技术人员所熟知的催化裂化反应器，只要能够在其中实现在不同的反应条件下进行所述第一和第二催化裂化反应即可。例如，所述催化裂化反应器可以为等直径提升管反应器、等线速提升管反应器、变径提升管反应器、流化床反应器、或等直径提升管与流化床构成的复合反应器，优选为变径提升管反应器。
在某些优选的实施方式中，所述催化裂化反应器为变径提升管反应器，该变径提升管反应器沿竖直方向由下至上依次设置有同轴且流体连通的预提升段、第一反应区、第二反应区和出口区，所述出口区末端设置有用于连接沉降器的水平管，所述第一反应区的内径小于第二反应区，所述第二反应区的内径大于出口区，所述催化裂化催化剂送入所述预提升段中，所述催化裂化原料油送入所述第一反应区的下部，所述第一催化裂化反应在第一反应区中进行，所述第二催化裂化 反应在第二反应区中进行。
在某些特别优选的实施方式中，所述变径提升管反应器的总高度(包括预提升段、第一反应区、第二反应区和出口区)为约10-60米。所述预提升段的直径与常规的等直径提升管反应器中相同，一般为约0.2-5米，其高度占反应器总高度的约5-20％。第一反应区的结构类似于常规的等直径提升管反应器，其直径可与预提升段相同，也可较预提升段稍大，第一反应区的直径与预提升段的直径之比为约1∶1至约2∶1，其高度占反应器总高度的约10-30％。原料油和催化剂在该区混合后，在较高的反应温度和剂油比、较短的反应时间(一般为约0.5-3.0秒)下，主要发生裂化反应。第二反应区比第一反应区直径大，其直径与第一反应区的直径之比为约1.5∶1至约5.0∶1，其高度占反应器总高度的约10-60％，其作用是降低油气和催化剂的流速和反应温度，使其主要发生选择性氢转移反应、异构化反应和芳构化反应。油气在该反应区内的反应时间可以较长，例如为约2-30秒。出口区直径与第一反应区直径之比为约0.8∶1至约1.5∶1，其高度占反应器总高度的约0-50％。第一、二反应区结合部位为圆台形，其纵剖面等腰梯形的顶角α为约30-80°，第二反应区与出口区结合部位为圆台形，其纵剖面等腰梯形的顶角β为约45-85°。降低第二反应区反应温度的方法，可以包括从该区与第一反应区的结合部位注入激冷介质，和/或通过在该区设置取热器以取走部分热量，从而达到抑制二次裂化反应、增加异构化反应和氢转移反应的目的。所述激冷介质可以选自冷激剂、冷却的再生催化剂和冷却的半再生催化剂，和它们的组合。所述冷激剂可以选自液化气、粗汽油、稳定汽油、轻循环油馏分、重轻循环油馏分、水，和它们的组合；所述冷却的再生催化剂和冷却的半再生催化剂可以由待生催化剂分别经两段再生和一段再生后冷却得到。再生催化剂碳含量一般为约0.1重量％以下，优选为约0.05重量％以下，半再生催化剂碳含量一般为约0.1-0.9重量％，优选为约0.15-0.7重量％。若设置取热器，则其高度占第二反应区高度的约50-90％。
在进一步优选的实施方式中，所述第一反应区和第二反应区的结合区域设置有至少一个用于输入激冷介质的激冷介质入口；和/或所述第二反应区中设置有取热器，所述取热器的高度占第二反应区高度的约50-90％。
在更进一步优选的实施方式中，所述激冷介质选自冷激剂、冷却的再生催化剂、冷却的半再生催化剂、新鲜催化剂，和它们的组合；所述冷激剂选自液化气、粗汽油、稳定汽油、轻循环油、重循环油、水，和它们的组合。
适用于本申请所公开方法的催化裂化催化剂可以是本领域技术人员所熟知的那些。例如，所述催化裂化催化剂可以包括裂化活性组分和载体；以干基计并以所述裂化活性组分的重量为基准，所述裂化活性组分可以包括约0-100重量％、优选约10-90重量％、更优选约20-40重量％的FAU型沸石和约0-100重量％、优选约10-90重量％、更优选约60-80重量％的五元环结构沸石，其中所述FAU型沸石和五元环结构沸石的总量为100重量％；所述FAU型沸石优选选自Y型沸石、HY型沸石、超稳Y型沸石，和它们的组合，所述五元环结构沸石优选选自ZSM-5系列沸石、高硅沸石、镁碱沸石，和它们的组合，所述五元环结构沸石可以任选地含有稀土和/或磷。
在本申请所公开方法的某些具体实施方式中，所述催化裂化催化剂使用后产生的待生催化剂可以送入再生器中进行烧焦再生，再生所得的再生催化剂作为催化裂化催化剂循环回催化裂化反应器中。为了增加第二反应区中的氢转移反应和异构化反应，如图1-3所示，可以任选地将沉降器中的待生催化剂的一部分作为激冷介质输送至第二反应区，以多产异构烷烃和轻芳烃。
在本申请所公开方法的某些优选实施方式中，在步骤d)中通过抽提精制从所述汽油馏分中回收轻质芳烃(BTX)。所述抽提精制可以按本领域技术人员熟知的方式进行，例如，所述抽提精制可以在如下条件下进行：抽提溶剂选自环丁砜、二甲亚砜、N-甲酰基吗啉、四甘醇、三甘醇、N-甲基吡啶烷酮，和它们的组合，温度为约50-110℃，抽提溶剂与汽油馏分的重量比为约2-6。
在一特别优选的实施方式中，本申请所公开的催化裂化方法包括如下步骤：
a)提供催化裂化原料油，其中以所述催化裂化原料油的重量计，所述催化裂化原料油中双环以上环烷烃的含量大于约25重量％；
b)将催化裂化原料油送入催化裂化反应器中与催化裂化催化剂接触并依次在不同反应条件下进行第一催化裂化反应和第二催化裂化反 应，得到催化裂化反应产物和待生催化剂；
c)将所得催化裂化反应产物进行分离，得到干气、液化气馏分、汽油馏分、轻循环油馏分、重循环油馏分、和任选的油浆；
d)任选地，将步骤c)中所得的汽油馏分进行抽提精制，得到轻质芳烃(例如苯、甲苯、二甲苯)，和/或将步骤c)中所得的液化气馏分进行分离得到丙烯和异丁烷产品；
e)任选地，将步骤c)中所得的轻循环油馏分、重循环油馏分和任选的油浆(优选经过过滤)送入加氢装置与加氢处理催化剂接触并进行加氢处理；以及
f)任选地，将所得的加氢尾油送入所述催化裂化反应器中作为所述催化裂化原料油。
在本申请所公开方法的一类优选实施方式中，将预热后的催化裂化原料油送入变径提升管反应器中，从反应器的第一反应区下部进入与催化剂接触，主要发生裂化反应；反应后生成的油气和催化剂的混合物上行至反应器的第二反应区下部，任选地与降温后的催化剂接触，主要进行氢转移反应和异构化反应，反应后流出物进入沉降器，分离得到反应油气(即催化裂化反应产物)和待生催化剂。反应油气经分离得到包括干气、液化气馏分、汽油馏分、轻循环油馏分、重循环油馏分和油浆的反应产物。任选地，将轻循环油馏分、重循环油馏分油和/或油浆，以及任选的来自其他处理单元的馏分油送往加氢装置，在加氢处理催化剂作用下进行加氢处理，所得的加氢尾油返回提升管反应器。待生催化剂经汽提和再生后返回提升管反应器。
下面结合附图1对本发明的此类优选实施方式予以进一步的说明，但是并不因此而限制本发明。
图1示意性地显示了用变径提升管反应器制取异丁烷和富含芳烃汽油的一种示例性流程，其中的设备和管线的形状、尺寸等不受附图的限制，而是根据具体情况确定。
如图1所示，预提升蒸汽经管线1从变径提升管反应器的预提升段2进入，热的再生催化剂经再生斜管16进入预提升段2，由预提升蒸汽进行提升。来自管线50的催化裂化原料油(例如，来自其他处理单元的馏分油)与来自管线41的加氢尾油合并后，经管线4与来自管线3的雾化蒸汽一起从预提升段2进入，与热再生催化剂混合后进入 第一反应区5内，在其中进行第一催化裂化反应。反应物流与来自管线6的冷激剂和/或冷却的催化剂(图中未示出)混合进入第二反应区7，进行第二催化裂化反应。反应后的物流进入出口区8，该区段可提高物流的线速，使反应物流快速进入气固分离系统的沉降器9和旋风分离器10，进行分离。分离得到的待生催化剂进入汽提器12，经汽提后进入待生立管14，在来自管线17的空气提升下进入再生器15，在其中烧焦再生。烟气经旋风分离器13分离后经管线18排出再生器，热的再生催化剂经再生斜管16返回提升管底部循环使用。任选地，可以将沉降器中的待生催化剂的一部分作为激冷介质输送至第二反应区7。分离得到的反应油气经管线11去往分馏系统19，经分馏得到的干气经管线20取出；液化气馏分经管线21送往后续处理装置(图中未示出)，以分离异丁烷；汽油馏分优选地经管线22进入抽提精制装置45，得到的苯、甲苯、二甲苯以及抽佘物分别从管线46、管线47、管线48和管线49送出；轻循环油馏分经管线23取出后，任选地经管线52送出催化裂化装置或者经管线26送往加氢装置31；重循环油馏分经管线24取出后，任选地经管线26送往加氢装置31；油浆任选地经管线25外甩出装置或者经管线51和管线26送往加氢装置31。管线26输送的物料与来自管线28的氢气一起进入加氢装置31进行加氢处理，加氢产物经高压分离器34和低压分离器35分离，液体产品从管线36经加氢装置分馏塔37分离得到气体、加氢汽油、加氢轻循环油和加氢尾油，分别经管线38、管线39、管线40和管线41送出。氢气经管线33进入到氢气循环压缩机32进行加压，而后经管线30以及管线29和管线28循环至加氢装置31。加氢尾油经管线41、管线4与来自管线3的雾化蒸汽一起进入预提升段2。
在本申请所公开方法的另一类优选实施方式中，初始原料油经加氢装置处理得到加氢尾油，其作为催化裂化原料油送入变径提升管反应器中。预热后的加氢尾油从反应器的第一反应区下部进入与催化剂接触，主要发生裂化反应，反应后生成的油气和催化剂的混合物上行至反应器的第二反应区下部，任选地与降温后的催化剂接触，主要进行氢转移反应和异构化反应，反应后流出物进入沉降器，分离得到反应油气和待生催化剂。反应油气经分离得到包括干气、液化气馏分、汽油馏分、轻循环油馏分、重循环油馏分和油浆的反应产物。任选地， 将轻循环油馏分、重循环油馏分和/或油浆，以及任选的来自其他处理单元的馏分油送往加氢装置，在加氢处理催化剂作用下进行加氢处理，所得的加氢尾油返回提升管反应器。待生催化剂经汽提和再生后返回提升管反应器。
下面结合附图2对本发明的此类优选实施方式予以进一步的说明，但是并不因此而限制本发明。
图2示意性地显示了用变径提升管反应器制取异丁烷和富含芳烃汽油的另一种示例性流程，其中的设备和管线的形状、尺寸等不受附图的限制，而是根据具体情况确定。
如图2所示，预提升蒸汽经管线1从变径提升管反应器的预提升段2进入，热的再生催化剂经再生斜管16进入预提升段2，由预提升蒸汽进行提升。来自管线27的初始原料油(例如，来自其他处理单元的馏分油)与来自管线28的氢气和任选的来自管线26的物料一起进入到加氢装置31进行加氢处理，加氢产物经高压分离器34和低压分离器35分离，液体产品从管线36经加氢装置分馏塔37分离得到气体、加氢汽油、加氢轻循环油和加氢尾油，分别经管线38、管线39、管线40和管线41送出。氢气经管线33进入到氢气循环压缩机32进行加压，而后经管线30以及管线29和管线28循环至加氢装置31。加氢尾油作为催化裂化原料油经管线41、管线4与来自管线3的雾化蒸汽一起从预提升段2进入，与热再生催化剂混合后进入第一反应区5内，在其中进行第一催化裂化反应。反应物流与来自管线6的冷激剂和/或冷却的催化剂(图中未示出)混合进入第二反应区7，进行第二催化裂化反应。反应后的物流进入出口区8，该区段可提高物流的线速，使反应物流快速进入气固分离系统的沉降器9和旋风分离器10，进行分离。分离得到的待生催化剂进入汽提器12，经汽提后进入待生立管14，在来自管线17的空气提升下进入再生器15，在其中烧焦再生。烟气经旋风分离器13分离后经管线18排出再生器，热的再生催化剂经再生斜管16返回提升管底部循环使用。任选地，可以将沉降器中的待生催化剂的一部分作为激冷介质输送至第二反应区7。分离得到的反应油气经管线11去往分馏系统19，经分馏得到的干气经管线20取出；液化气馏分经管线21送往后续处理装置(图中未示出)，以分离异丁烷；汽油馏分优选地经管线22进入抽提精制装置45，得到的苯、甲苯、二甲苯 以及抽佘物分别从管线46、管线47、管线48和管线49送出；轻循环油馏分经管线23取出后，任选地经管线52送出催化裂化装置或者经管线26送往加氢装置31；重循环油馏分经管线24取出后，任选地经管线26送往加氢装置31；油浆任选地经管线25外甩出装置或者经管线51和管线26送往加氢装置31。管线26输送的物料与来自管线27的初始原料油和来自管线28的氢气一起进入加氢装置31进行加氢处理。
在本申请所公开方法的另一类优选实施方式中，初始原料油经芳烃抽提装置处理得到富含多环芳烃的抽出油和富含多环环烷烃的抽佘油。抽出油经加氢装置处理得到加氢尾油，所述加氢尾油和所述抽佘油合并后作为催化裂化原料油送入变径提升管反应器中。预热后的催化裂化原料油从反应器的第一反应区下部进入与催化剂接触，主要发生裂化反应，反应后生成的油气和催化剂的混合物上行至反应器的第二反应区下部，任选地与降温后的催化剂接触，主要进行氢转移反应和异构化反应，反应后流出物进入沉降器，分离得到反应油气和待生催化剂。反应油气经分离得到包括干气、液化气馏分、汽油馏分、轻循环油馏分、重循环油馏分和油浆的反应产物。任选地，将轻循环油馏分、重循环油馏分和/或油浆，以及任选的来自其他处理单元的馏分油送往加氢装置，在加氢处理催化剂作用下进行加氢处理，所得的加氢尾油返回提升管反应器。待生催化剂经汽提和再生后返回提升管反应器。
下面结合附图3对本发明的此类优选实施方式予以进一步的说明，但是并不因此而限制本发明。
图3示意性地显示了用变径提升管反应器制取异丁烷和富含芳烃汽油的另一种示例性流程，其中的设备和管线的形状、尺寸等不受附图的限制，而是根据具体情况确定。
如图3所示，预提升蒸汽经管线1从变径提升管反应器的预提升段2进入，热的再生催化剂经再生斜管16进入预提升段2，由预提升蒸汽进行提升。来自管线43的初始原料油(例如，来自其他处理单元的馏分油)进入芳烃抽提装置42，经抽提得到富含多环芳烃的抽出油和富含多环环烷烃的抽佘油。抽出油经管线27与来自管线28的氢气和任选的来自管线26的物料一起进入到加氢装置31进行加氢处理， 加氢产物经高压分离器34和低压分离器35分离，液体产品经加氢装置分馏塔37分离得到气体、加氢汽油、加氢轻循环油和加氢尾油，分别经管线38、管线39、管线40和管线41送出。氢气经管线33进入到氢气循环压缩机32进行加压，而后经管线30以及管线29和管线28循环至加氢装置31。加氢尾油与来自管线44的抽佘油合并后作为催化裂化原料油，经管线41、管线4与来自管线3的雾化蒸汽一起从预提升段2进入，与热再生催化剂混合后进入第一反应区5内，在其中进行第一催化裂化反应。反应物流与来自管线6的冷激剂和/或冷却的催化剂(图中未示出)混合进入第二反应区7，进行第二催化裂化反应。反应后的物流进入出口区8，该区段可提高物流的线速，使反应物流快速进入气固分离系统的沉降器9和旋风分离器10，进行分离。分离得到的待生催化剂进入汽提器12，经汽提后进入待生立管14，在来自管线17的空气提升下进入再生器15，在其中烧焦再生。烟气经旋风分离器13分离后经管线18排出再生器，热的再生催化剂经再生斜管16返回提升管底部循环使用。任选地，可以将沉降器中的待生催化剂的一部分作为激冷介质输送至第二反应区7。分离得到的反应油气经管线11去往分馏系统19，经分馏得到的干气经管线20取出；液化气馏分经管线21送往后续处理装置(图中未示出)，以分离异丁烷；汽油馏分优选地经管线22进入抽提精制装置45，得到的苯、甲苯、二甲苯和抽佘物分别从管线46、管线47、管线48和管线49送出；轻循环油馏分经管线23取出后，任选地经管线52送出催化裂化装置或者经管线26送往加氢装置31；重循环油馏分经管线24取出后，任选地经管线26送往加氢装置31；油浆任选地经管线25外甩出装置或者经管线51和管线26送往加氢装置31。管线26输送的物料与来自管线27的抽出油和来自管线28的氢气一起进入加氢装置31进行加氢处理。
在优选的实施方式中，本申请提供了以下的技术方案：
A1、一种多产异丁烷和轻质芳烃的催化裂化方法，该方法包括：
(1)、将催化裂化原料油送入催化裂化反应器中与催化裂化催化剂接触并依次在不同反应条件下进行第一催化裂化反应和第二催化裂化反应，得到催化裂化反应产物和待生催化剂；其中，所述催化裂化原料油中双环以上环烷烃的含量大于约25重量％；
(2)、将所得催化裂化反应产物进行分离，得到干气、液化气、 汽油、轻循环油馏分、重循环油馏分、以及得到或不得到油浆；
(3)、将步骤(2)中所得轻循环油馏分、重循环油馏分和油浆任选地送入加氢装置与加氢处理催化剂接触并进行选择性加氢处理，得到加氢尾油作为所述催化裂化原料油送入所述催化裂化反应器中；
(4)、任选将步骤(2)中所得汽油进行抽提精制，得到轻质芳烃。
A2、根据项目A1所述的方法，其中，所述催化裂化原料油中双环以上环烷烃的含量大于约40重量％。
A3、根据项目A1所述的方法，其中所述催化裂化原料油选自深度加氢轻循环油、延迟焦化装置的重馏分油、催化裂化轻循环油、催化裂化重循环油、催化裂化重馏分油、油浆、加氢裂化柴油、渣油加氢裂化柴油、蜡油加氢裂化柴油、生物柴油、页岩油柴油馏分、煤液化柴油馏分、常压塔顶油、常压塔抽出的馏分油、直馏减压蜡油、加氢蜡油、焦化蜡油、脱沥青油、抽出油、抽佘油、常压渣油和减压渣油中的至少一种一次加工馏分油和/或二次加工原料，和/或至少一种所述一次加工馏分油和/或二次加工原料进行加氢所得的加氢尾油。
A4、根据项目A3所述的方法，其中若所述二次加工原料中双环以上环烷烃的含量不大于约25重量％，将所述二次加工原料进行芳烃抽提和/或所述选择性加氢处理后再作为所述催化裂化原料油。
A5、根据项目A1所述的方法，其中所述第一催化裂化反应的条件包括：反应温度为约520-620℃，反应时间为约0.5-3.0秒，剂油重量比为约3∶1至约15∶1；
所述第二催化裂化反应的条件包括：反应温度为约420-530℃，反应时间为约2-30秒，剂油重量比为约3∶1至约18∶1。
A6、根据项目A1所述的方法，其中，所述第一催化裂化反应的条件包括：反应温度为约530-600℃，反应时间为约0.8-2.0秒，剂油重量比为约4∶1至约12∶1；
所述第二催化裂化反应的条件包括：反应温度为约460-510℃，反应时间为约3-15秒，剂油重量比为约4∶1至约15∶1。
A7、根据项目A1所述的方法，其中，所述催化裂化反应器为等直径提升管、等线速提升管、变径提升管、流化床、或等直径提升管与流化床构成的复合反应器。
A8、根据项目A7所述的方法，其中，所述变径提升管沿竖直方 向由下至上依次设置有同轴且流体连通的预提升段、第一反应区、第二反应区和出口区，所述出口区末端设置有用于连接沉降器的水平管，所述第一反应区的内径小于第二反应区，所述第二反应区的内径大于出口区，所述变径提升管的总高度为约10-60米，所述催化裂化催化剂送入所述预提升段中，所述催化裂化原料油送入所述第一反应区的下部，所述第一催化裂化反应在第一反应区中进行，所述第二催化裂化反应在第二反应区中进行。
A9、根据项目A8所述的方法，其中，所述第一反应区和第二反应区的结合区域设置有至少一个用于送入激冷介质的激冷介质入口；和/或
所述第二反应区中设置有取热器，所述取热器的高度占第二反应区高度的约50-90％。
A10、根据项目A9所述的方法，其中，所述激冷介质选自冷激剂、冷却的再生催化剂、冷却的半再生催化剂和新鲜催化剂，和它们的组合，所述冷激剂选自液化气、粗汽油、稳定汽油、轻循环油馏分、重循环油馏分和水，和它们的组合。
A11、根据项目A1所述的方法，其中，所述催化裂化催化剂包括裂化活性组分和载体；以干基计，所述裂化活性组分包括约0-100重量％、优选约10-90重量％的FAU型沸石和约0-100重量％、优选约10-90重量％的五元环结构沸石；所述FAU型沸石选自Y型沸石、HY型沸石和超稳Y型沸石，和它们的组合，所述五元环结构沸石选自ZSM-5系列沸石、高硅沸石和镁碱沸石，和它们的组合，所述五元环结构沸石含或不含稀土以及含或不含磷。
A12、根据项目A1或A4所述的方法，其中，所述选择性加氢处理的条件包括：氢分压为10.0-30.0MPa，反应温度为300-500℃，液相体积空速为约0.1-10.0h -1，氢油体积比为100-1500Nm 3/m 3。
A13、根据项目A1或A4所述的方法，其中，所述加氢处理催化剂包括加氢活性组分和载体，所述加氢活性组分选自第VIB族非贵金属、第VIII族非贵金属，和它们的组合，所述载体选自氧化铝、二氧化硅和无定型硅铝，和它们的组合。
A14、根据项目A1所述的方法，其中，步骤(4)中所述抽提精制的条件包括：抽提溶剂选自环丁砜、二甲亚砜、N-甲酰基吗啉、四 甘醇、三甘醇和N-甲基吡啶烷酮，和它们的组合，温度为约50-110℃，抽提溶剂与汽油的重量比为约2-6。
A15、根据项目A4所述的方法，其中所述芳烃抽提的条件包括：温度为约50-70℃，溶剂与原料重量比为约0.5-2，溶剂选自糠醛、二甲亚砜、二甲基甲酰胺、单乙醇胺、乙二醇和1，2-丙二醇，和它们的组合。
A16、根据项目A1或A4所述的方法，其中所述选择性加氢处理的条件包括：氢分压为约6.0-30.0MPa，反应温度为约300-450℃，液相体积空速为约0.1-10.0h -1，氢油体积比为约300-3000Nm 3/m 3；优选地，氢分压为约8-20MPa，反应温度为约330-430℃，液时体积空速为约0.2-5h -1，氢油体积比为约500-2500Nm 3/m 3。
B1、一种催化裂化方法，包括如下步骤：
a)提供催化裂化原料油，其中以所述催化裂化原料油的重量计，所述催化裂化原料油中双环以上环烷烃的含量大于约25重量％、优选大于约40重量％；
b)使所述催化裂化原料油在催化裂化反应器中与催化裂化催化剂接触，并在不同的反应条件下依次进行第一催化裂化反应和第二催化裂化反应，得到催化裂化反应产物；
c)分离所得的催化裂化反应产物，得到包含异丁烷的液化气馏分和包含轻质芳烃的汽油馏分；以及
d)任选地，从所述液化气馏分中回收异丁烷，和/或从所述汽油馏分中回收轻质芳烃。
B2、根据项目B1所述的方法，其中步骤c)中的所述分离还得到轻循环油馏分、重循环油馏分，和任选的油浆，并且所述方法进一步包括如下步骤：
e)对步骤c)中得到的轻循环油馏分、重循环油馏分和任选的油浆的至少一部分进行加氢处理得到加氢尾油；以及
f)将所得的加氢尾油的至少一部分再循环到所述催化裂化反应器中。
B3、根据项目B1或B2所述的方法，其中步骤a)进一步包括对双环以上环烷烃含量不大于约25重量％的初始原料油进行预处理，以得到所述双环以上环烷烃含量大于约25重量％、优选大于约40重量％的 催化裂化原料油。
B4、根据项目B3所述的方法，其中所述预处理包括芳烃抽提和/或加氢处理。
B5、根据项目B2或B4所述的方法，其中所述步骤e)的加氢处理和/或作为所述预处理的加氢处理在如下条件下进行：氢分压为约6.0-30.0MPa；反应温度为约300-450℃；液相体积空速为约0.1-10.0h -1；氢油体积比为约300-3000Nm 3/m 3。
B6、根据项目B2或B4所述的方法，其中所述步骤e)的加氢处理和/或作为所述预处理的加氢处理在如下条件下进行：氢分压为约8-20MPa；反应温度为约330-430℃；液相体积空速为约0.2-5h -1；氢油体积比为约500-2500Nm 3/m 3。
B7、根据项目B4-B6中任一项所述的方法，其中所述步骤e)的加氢处理和/或作为所述预处理的加氢处理在加氢处理催化剂存在下进行，所述加氢处理催化剂包括加氢活性组分和载体，所述加氢活性组分优选选自第VIB族非贵金属、第VIII族非贵金属，和它们的组合，所述载体优选选自氧化铝、二氧化硅、无定型硅铝，和它们的组合。
B8、根据项目B4-B6中任一项所述的方法，其中所述芳烃抽提在如下条件下进行：温度为约50-70℃，溶剂与原料重量比为约0.5-2，溶剂选自糠醛、二甲亚砜、二甲基甲酰胺、单乙醇胺、乙二醇、1，2-丙二醇，和它们的组合。
B9、根据前述项目中任一项所述的方法，其中所述催化裂化原料油或者所述初始原料油选自深度加氢轻循环油、延迟焦化装置的重馏分油、催化裂化轻循环油、催化裂化重循环油、催化裂化重馏分油、油浆、加氢裂化柴油、渣油加氢裂化柴油、蜡油加氢裂化柴油、生物柴油、页岩油柴油馏分、煤液化柴油馏分、常压塔顶油、常压塔抽出的馏分油、直馏减压蜡油、加氢蜡油、焦化蜡油、脱沥青油、抽出油、抽佘油、常压渣油、减压渣油、上述原料油经加氢得到的加氢尾油，和它们的组合。
B10、根据前述项目中任一项所述的方法，其中：
所述第一催化裂化反应在如下条件下进行：反应温度为约520-620℃；反应时间为约0.5-3.0秒；剂油重量比为约3∶1至约15∶1；
所述第二催化裂化反应在如下条件下进行：反应温度为约480-600 ℃；反应时间为约2-30秒；剂油重量比为约3∶1至约18∶1。
B11、根据前述项目中任一项所述的方法，其中：
所述第一催化裂化反应在如下条件下进行：反应温度为约530-600℃；反应时间为约0.8-2.0秒；剂油重量比为约4∶1至约12∶1；
所述第二催化裂化反应在如下条件下进行：反应温度为约500-550℃；反应时间为约3-15秒；剂油重量比为约4∶1至约15∶1。
B12、根据前述项目中任一项所述的方法，其中所述催化裂化反应器为等直径提升管反应器、等线速提升管反应器、变径提升管反应器、流化床反应器、或等直径提升管与流化床构成的复合反应器。
B13、根据前述项目中任一项所述的方法，其中所述催化裂化反应器为变径提升管反应器，所述变径提升管反应器沿竖直方向由下至上依次设置有同轴且流体连通的预提升段、第一反应区、第二反应区和出口区，所述第一反应区的内径小于第二反应区，所述第二反应区的内径大于出口区，所述催化裂化催化剂送入所述预提升段中，所述催化裂化原料油送入所述第一反应区的下部，所述第一催化裂化反应在第一反应区中进行，所述第二催化裂化反应在第二反应区中进行。
B14、根据项目B13所述的方法，其中所述第一反应区和第二反应区的结合区域设置有至少一个用于输入激冷介质的激冷介质入口；和/或
所述第二反应区中设置有取热器，所述取热器的高度占第二反应区高度的约50-90％。
B15、根据项目B14所述的方法，其中所述激冷介质选自冷激剂、冷却的再生催化剂、冷却的半再生催化剂、新鲜催化剂，和它们的组合；所述冷激剂选自液化气、粗汽油、稳定汽油、轻循环油、重循环油、水，和它们的组合。
B16、根据前述项目中任一项所述的方法，其中所述催化裂化催化剂包括裂化活性组分和载体；以干基计并以所述裂化活性组分的重量为基准，所述裂化活性组分包括约0-100重量％的FAU型沸石和约0-100重量％的五元环结构沸石，其中所述FAU型沸石和五元环结构沸石的总量为100重量％。
B17、根据前述项目中任一项所述的方法，其中所述催化裂化催化剂包括裂化活性组分和载体；以干基计并以所述裂化活性组分的重量 为基准，所述裂化活性组分包括约10-90重量％的FAU型沸石和约10-90重量％的五元环结构沸石，其中所述FAU型沸石和五元环结构沸石的总量为100重量％。
B18、根据项目B16或B17所述的方法，其中所述FAU型沸石选自Y型沸石、HY型沸石、超稳Y型沸石，和它们的组合；所述五元环结构沸石选自ZSM-5系列沸石、高硅沸石、镁碱沸石，和它们的组合，所述五元环结构沸石任选地含有稀土和/或磷。
B19、根据前述项目中任一项所述的方法，其中在步骤d)中通过抽提精制从所述汽油馏分中回收轻质芳烃。
B20、根据项目B19所述的方法，其中所述抽提精制在如下条件下进行：抽提溶剂选自环丁砜、二甲亚砜、N-甲酰基吗啉、四甘醇、三甘醇、N-甲基吡啶烷酮，和它们的组合，温度为约50-110℃，抽提溶剂与汽油馏分的重量比为约2-6。
实施例
下面将通过实施例来对本发明做进一步说明，但是本发明并不因此而受到任何限制。
原料和试剂
以下实施例和对比例中所涉及的各原料油的性质列于表1和表2。
表1实施例1-2和对比例1-2所涉及的各原料油的性质
Figure PCTCN2018111179-appb-000001
注：“-”表示未测定。
表2实施例3-4和对比例3-4所涉及的各原料油的性质
Figure PCTCN2018111179-appb-000002
注：“-”表示未测定。
表2(续)实施例3-4和对比例3-4所涉及的各原料油的性质
Figure PCTCN2018111179-appb-000003
注：“-”表示未测定。
以下实施例和对比例中所用的催化裂化催化剂IBA-1采用高硅Y沸石和ZRP沸石作为裂化活性组分，两者的重量比例为高硅Y沸石30重量％，ZRP沸石70重量％。所述催化剂的具体制备方法如下：
用4300克脱阳离子水将969克多水高岭土(中国高岭土公司产品，固含量73％)打浆，再加入781克拟薄水铝石(山东淄博铝石厂产品，固含量64％)和144ml盐酸(浓度30％，比重1.56)搅拌均匀，在60℃静置老化1小时，保持pH为2-4，降至常温，再加入预先准备好的含800克高硅Y沸石(干基)(中国石化催化剂分公司齐鲁催化剂厂)和2000克含化学水的ZRP沸石(中国石化催化剂分公司齐鲁催化剂厂)浆液，搅拌均匀，喷雾干燥，洗去游离Na +，得催化剂。将所得催化剂在800℃和100％水蒸汽条件下老化12小时，得到催化剂IBA-1，其理化性质列于表3。
以下实施例和对比例中所用的商业牌号为CGP-1、MLC-500和DMMC-1的催化裂化催化剂均由中国石化催化剂分公司齐鲁催化剂厂所生产，其理化性质列于表3。
表3实施例和对比例中所用催化裂化催化剂的理化性质
Figure PCTCN2018111179-appb-000004
以下实施例和对比例中使用的商业牌号为RN-32V的加氢处理催化剂，和商业牌号为RG-30A、RG-30B、RG-1的保护剂，均由中国石化催化剂分公司生产，加氢装置内加氢处理催化剂和保护剂的装填体积比为95∶5。
检测方法
催化裂化催化剂的微反活性(MAT)采用RIPP 92-90的标准方法(参见《石油化工分析方法(RIPP试验方法)》，杨翠定等编，科学出版社，1990年9月第一版，第263-268页)进行测定，具体测定条件 如下：催化剂5.0克；进油量1.56克；反应时间70秒；反应温度460℃；剂/油3.2；空速16h -1。
催化裂化催化剂的比表面积和总孔体积采用Quantachrome仪器公司生产的AS-3，AS-6静态氮吸附仪测定，具体测定方法如下所述：将样品置于样品处理系统，在300℃下抽真空至1.33×10 -2Pa，保温保压4h，净化样品；在液氮温度-196℃下，测试净化样品在不同比压P/P 0条件下对氮气的吸附量和脱附量，获得N 2吸附-脱附等温曲线；然后利用两参数BET公式计算总比表面积、微孔比表面积和中孔比表面积，取比压P/P 0＝0.98以下的吸附量为样品的总孔体积。
实施例1-A
采用图1所示的工艺流程，预热的催化裂化原料油A在中型变径提升管催化裂化装置上，与催化剂IBA-1接触并进行催化裂化反应。预热的原料油A进入如图1所示的变径提升管反应器内，在水蒸汽存在下，与热的催化裂化催化剂IBA-1在第一反应区和第二反应区接触进行第一催化裂化反应和第二催化裂化反应，分离催化裂化反应产物得到干气、富含异丁烷的液化气、富含芳烃的汽油、轻循环油和重循环油(无油浆)；轻循环油外甩出装置，重循环油循环至加氢装置进行加氢处理，所得加氢尾油A′循环回所述提升管反应器进行催化裂化反应；待生催化剂经汽提后进入再生器，烧焦再生后循环使用。原料油性质列于表1，操作条件列于表4，产品分布列于表5，汽油经抽提精制得到的产品性质见表5。
实施例1-B
按照与实施例1所述类似的方式进行操作，其中采用常规催化裂化催化剂CGP-1。预热的催化裂化原料油A进入如图1所示的变径提升管反应器内，在水蒸汽存在下，与热的催化裂化催化剂CGP-1在第一反应区和第二反应区接触进行第一催化裂化反应和第二催化裂化反应，分离催化裂化反应产物得到干气、富含异丁烷的液化气、富含芳烃的汽油、轻循环油和重循环油(无油浆)；轻循环油外甩出装置，重循环油循环至加氢装置进行加氢处理，所得加氢尾油循环回所述提升管反应器进行催化裂化反应；待生催化剂经汽提后进入再生器，烧焦 再生后循环使用。原料油性质列于表1，操作条件列于表4，产品分布列于表5，汽油经抽提精制得到的产品性质见表5。
实施例1-C
按照与实施例1所述类似的方式进行操作，其中采用常规催化裂化催化剂MLC-500。预热的催化裂化原料油A进入常规的等径提升管反应器内，在该等径提升管反应器的中部注入激冷介质使其形成位于下部的第一反应区和位于上部的第二反应区。所述原料油A在水蒸汽存在下，与热的催化裂化催化剂MLC-500在第一反应区和第二反应区接触进行第一催化裂化反应和第二催化裂化反应，分离催化裂化反应产物得到干气、富含异丁烷的液化气、富含芳烃的汽油、轻循环油和重循环油(无油浆)；轻循环油外甩出装置，重循环油循环至加氢装置进行加氢处理，所得加氢尾油循环回所述提升管反应器进行催化裂化反应；待生催化剂经汽提后进入再生器，烧焦再生后循环使用。原料油性质列于表1，操作条件列于表4，产品分布列于表5，汽油经抽提精制得到的产品性质见表5。
实施例1-D
按照与实施例1所述类似的方式进行操作，其中采用常规催化裂化催化剂DMMC-1。预热的催化裂化原料油A进入提升管反应器+流化床的复合反应器内，其中所述提升管反应器充当第一反应区，所述流化床充当第二反应区。所述原料油A在水蒸汽存在下，与热的催化裂化催化剂DMMC-1在第一反应区和第二反应区接触进行第一催化裂化反应和第二催化裂化反应，分离催化裂化反应产物得到干气、含异丁烷的液化气、富含芳烃的汽油、轻循环油和重循环油(无油浆)；轻循环油外甩出装置，重循环油循环至加氢装置进行加氢处理，所得加氢尾油循环回所述提升管反应器进行催化裂化反应；待生催化剂经汽提后进入再生器，烧焦再生后循环使用。原料油性质列于表1，操作条件列于表4，产品分布列于表5，汽油经抽提精制得到的产品性质见表5。
对比例1-A
按照与实施例1所述类似的方式进行操作，其中采用多环环烷烃含量低的原料油D。预热的原料油D进入如图1所示的变径提升管反应器内，在水蒸汽存在下，与热的催化裂化催化剂CGP-1接触反应，分离催化裂化反应产物得到干气、液化气、汽油、轻循环油、重循环油和油浆；轻循环油和油浆均外甩出装置，重循环油循环至加氢装置进行加氢处理后返回所述提升管反应器，待生催化剂经汽提后进入再生器，烧焦再生后循环使用。原料油性质列于表1，操作条件列于表4，产品分布列于表5，汽油经抽提精制得到的产品性质见表5。
对比例1-B
按照与实施例1所述类似的方式进行操作，其中采用常规催化裂化催化剂MLC-500。预热的原料油A进入常规的等径提升管反应器内，在水蒸汽存在下，与热的催化裂化催化剂接触反应，分离催化裂化反应产物得到干气、液化气、汽油、轻循环油、重循环油和油浆；轻循环油和油浆均外甩出装置，重循环油循环至加氢装置进行加氢处理后返回所述提升管反应器，待生催化剂经汽提后进入再生器，烧焦再生后循环使用。原料油性质列于表1，操作条件列于表4，产品分布列于表5，汽油经抽提精制得到的产品性质见表5。
对比例1-C
按照与实施例1所述类似的方式进行操作，其中采用催化裂化催化剂IBA-1。预热的原料油A进入常规的等径提升管反应器内，在水蒸汽存在下，与热的催化裂化催化剂接触反应，分离催化裂化反应产物得到干气、液化气、汽油、轻循环油、重循环油和油浆；轻循环油和油浆均外甩出装置，重循环油循环至加氢装置进行加氢处理后返回所述提升管反应器，待生催化剂经汽提后进入再生器，烧焦再生后循环使用。原料油性质列于表1，操作条件列于表4，产品分布列于表5，汽油经抽提精制得到的产品性质见表5。
表4实施例1-A至1-D和对比例1-A至1-C采用的操作条件
Figure PCTCN2018111179-appb-000005
注：“-”表示未测定。
表4(续)实施例1-A至1-D和对比例1-A至1-C采用的操作条件
项目 对比例1-A 对比例1-B 对比例1-C
原料油编号 D A A
催化剂类型 CGP-1 MLC-500 IBA-1
反应器类型 变径提升管 等径提升管 等径提升管
反应温度，℃   530 530
第一反应区 550 - -
第二反应区 530 - -
反应时间，秒 5.2 3.1 3.1
第一反应区 1.2 - -
第二反应区 4.0 - -
剂油比 5.0 5.0 5.0
水油比 0.1 0.1 0.1
床层温度，℃ - - -
床层空速，h -1 - - -
加氢装置      
反应温度，℃ 370 370 370
反应压力，MPa 17.0 17.0 17.0
体积空速，h -1 0.5 0.5 0.5
氢油体积比，Nm 3/m 3 1000 1000 1000
抽提精制装置      
抽提溶剂 环丁砜 环丁砜 环丁砜
抽提温度，℃ 65 65 65
溶剂/油重量比 4 4 4
注：“-”表示未测定。
表5实施例1-A至1-D和对比例1-A至1-C的反应结果
Figure PCTCN2018111179-appb-000006
表5(续)实施例1-A至1-D和对比例1-A至1-C的反应结果
项目 对比例1-A 对比例1-B 对比例1-C
原料油编号 D A A
催化剂类型 CGP-1 MLC-500 IBA-1
反应器类型 变径提升管 等径提升管 等径提升管
产品分布，重量％      
干气 1.72 1.01 1.81
液化气 25.85 17.15 22.03
丙烯 9.2 5.82 7.18
异丁烯 2.71 2.63 3.13
异丁烷 5.88 2.98 7.94
汽油 51.06 55.61 46.25
苯 0.83 0.51 0.89
甲苯 5.77 3.41 4.89
二甲苯 7.95 5.24 6.94
BTX 14.55 9.16 12.72
轻循环油 10.02 14.15 17.91
油浆 5.07 6.19 6.87
焦炭 6.28 5.89 5.13
合计 100.00 100.00 100.00
实施例2
采用图2所示的工艺流程，先对初始原料油B进行加氢饱和得到催化裂化原料油C。预热的催化裂化原料油C进入如图2所示的变径提升管反应器内，在水蒸汽存在下，与热的催化剂IBA-1接触反应，分离催化裂化反应产物得到干气、富含异丁烷的液化气、富含芳烃的汽油、轻循环油和重循环油(无油浆)；轻循环油外甩出装置，重循环油循环至加氢装置进行加氢处理，所得加氢尾油循环回所述提升管反应器进行催化裂化反应，待生催化剂经汽提后进入再生器，烧焦再生后循环使用。原料油性质列于表1，操作条件列于表6，产品分布列于表7，汽油经抽提精制得到的产品性质见表7。
对比例2
采用图1所示的工艺流程，原料油B未加氢，直接将预热的原料油B送入如图1所示的变径提升管反应器内，在水蒸汽存在下，与热的催化剂IBA-1接触反应，分离催化裂化反应产物得到干气、液化气、汽油、轻循环油、重循环油和油浆；轻循环油和油浆外甩出装置，重循环油循环至加氢装置进行加氢处理后返回所述提升管反应器；待生催化剂经汽提后进入再生器，烧焦再生后循环使用。原料油性质列于表1，操作条件列于表6，产品分布列于表7，汽油经抽提精制得到的产品性质见表7。
表6实施例2和对比例2的操作条件
Figure PCTCN2018111179-appb-000007
表7实施例2和对比例2的反应结果
Figure PCTCN2018111179-appb-000008
实施例3
采用图3所示的工艺流程，原料油减压馏分油E经芳烃抽提装置处理后得到减压馏分油抽出油G和减压馏分油抽佘油F，芳烃抽提装置条件：温度为60℃，溶剂与原料重量比为1.5，溶剂为糠醛。减压馏分油抽出油G和氢气混合后进入到加氢装置进行加氢处理，自加氢装置得到加氢尾油G′，将预热的加氢尾油和减压馏分油抽佘油F混合后作为催化裂化原料油进入如图3所示的变径提升管反应器内，在水蒸汽存在下，与热的催化剂IBA-1接触反应，分离催化裂化反应产物得到干气、富含异丁烷的液化气、富含芳烃的汽油、轻循环油和重循环 油(无油浆)；轻循环油外甩出装置，重循环油和减压馏分油抽出油G混合后送至加氢装置进行加氢处理，所得加氢尾油循环回所述提升管反应器进行催化裂化反应；待生催化剂经汽提后进入再生器，烧焦再生后循环使用。原料油性质列于表2，操作参数列于表8，产品分布列于表9，汽油经抽提精制得到的产品性质见表9。
对比例3
采用图1所示的工艺流程，原料油E未经抽提和加氢处理，直接将预热的原料油E送入如图1所示的变径提升管反应器内，在水蒸汽存在下，与热的催化剂IBA-1接触反应，分离催化裂化反应产物得到干气、液化气、汽油、轻循环油、重循环油和油浆；其中轻循环油、油浆外甩出装置，重循环油循环至加氢装置进行加氢处理后返回所述提升管反应器；待生催化剂经汽提后进入再生器，烧焦再生后循环使用。原料性质列于表2，操作参数列于表8，产品分布列于表9，汽油经抽提精制得到的产品性质见表9。
实施例4
采用图3所示的工艺流程，原料油减压渣油H经芳烃抽提装置处理后得到减压渣油抽出油J和减压渣油抽佘油I，芳烃抽提装置条件：温度为60℃，溶剂与原料重量比为1.5，溶剂为糠醛。减压渣油抽出油J和氢气混合后进入到加氢装置进行加氢处理，自加氢装置得到加氢尾油J′，将预热的加氢尾油和减压渣油抽佘油I混合后作为催化裂化原料油进入如图3所示的变径提升管反应器内，在水蒸汽存在下，与热的催化剂IBA-1接触反应，分离催化裂化反应产物得到干气、富含异丁烷的液化气、富含芳烃的汽油、轻循环油和重循环油(无油浆)，轻循环油外甩出装置，重循环油和减压渣油抽出油J混合后送至加氢装置进行加氢处理，所得加氢尾油循环回所述提升管反应器进行催化裂化反应；待生催化剂经汽提后进入再生器，烧焦再生后循环使用。原料油性质列于表2，操作参数列于表8，产品分布列于表9，汽油经抽提精制得到的产品性质见表9。
对比例4
采用图1所示的工艺流程，原料油减压渣油H未经抽提和加氢处理，直接将预热的原料油H送入如图1所示的变径提升管反应器内，在水蒸汽存在下，与热的催化剂IBA-1接触反应，分离催化裂化反应产物得到干气、液化气、汽油、轻循环油、重循环油和油浆；其中轻循环油、油浆外甩出装置，重循环油循环至加氢装置进行加氢处理后返回所述提升管反应器；待生催化剂经汽提后进入再生器，烧焦再生后循环使用。原料油性质列于表2，操作参数列于表8，产品分布列于表9，汽油经抽提精制得到的产品性质见表9。
表8实施例3-4和对比例3-4中采用的操作条件
项目 实施例3 对比例3 实施例4 对比例4
原料油编号 E E H H
催化剂类型 IBA-1 IBA-1 IBA-1 IBA-1
反应器类型 变径提升管 变径提升管 变径提升管 变径提升管
反应温度，℃        
第一反应区 550 550 550 550
第二反应区 530 530 530 530
反应时间，秒        
第一反应区 1.3 1.3 1.3 1.3
第二反应区 4.4 4.4 4.4 4.4
剂油比 5.0 5.0 5.0 5.0
水油比 0.1 0.1 0.1 0.1
加氢装置        
反应温度，℃ 375 375 385 385
反应压力，MPa 14.0 14.0 16.0 16.0
体积空速，h -1 0.7 0.7 0.3 0.3
氢油体积比，Nm 3/m 3 1200 1200 900 900
抽提精制装置        
抽提溶剂 环丁砜 环丁砜 环丁砜 环丁砜
抽提温度，℃ 65 65 65 65
溶剂/油重量比 4 4 4 4
表9实施例3-4和对比例3-4的反应结果
项目 实施例3 对比例3 实施例4 对比例4
原料油编号 E E H H
催化剂类型 IBA-1 IBA-1 IBA-1 IBA-1
反应器类型 变径提升管 变径提升管 变径提升管 变径提升管
产品分布，重量％        
干气 2.49 1.52 2.79 3.14
液化气 36.81 20.71 33.83 16.51
丙烯 12.52 7.10 11.31 5.82
异丁烯 1.31 3.19 1.12 2.52
异丁烷 13.75 4.15 12.44 2.44
汽油 49.04 42.77 45.70 36.16
苯 1.91 0.87 1.86 0.81
甲苯 8.87 5.31 8.56 5.83
二甲苯 12.84 7.35 12.77 7.95
BTX 23.62 13.53 23.19 14.59
轻循环油 6.70 24.87 8.18 23.86
油浆 0.00 6.33 0.00 7.90
焦炭 4.96 3.80 9.50 12.43
合计 100.00 100.00 100.00 100.00
从以上实施例和对比例的结果比较可以看出，本发明的方法能够多产异丁烷和/或轻质芳烃，特别是在采用变径提升管反应器的情况下。
在上文的说明书中，已经参照特定的实施方式描述了本发明的构思。然而，本领域技术人员可以理解，在不脱离所附的权利要求中限定的本发明范围的情况下可以做出各种修改和变更。因此，说明书和附图应认为是说明性的，而不是限制性的，并且所有这类修改和变更应当涵盖在本发明的范围之内。
可以理解，本文为清楚起见以独立的多个实施方式的形式描述的某些特征也可以作为组合提供在单一的实施方式中。相反，为简要起 见以单一实施方式的形式描述的多个不同特征也可以单独地或以任何子组合的形式提供。

Claims (14)

  1. 一种催化裂化方法，包括如下步骤：
    a)提供催化裂化原料油，其中以所述催化裂化原料油的重量计，所述催化裂化原料油中双环以上环烷烃的含量大于约25重量％、优选大于约40重量％；
    b)使所述催化裂化原料油在催化裂化反应器中与催化裂化催化剂接触，并在不同的反应条件下依次进行第一催化裂化反应和第二催化裂化反应，得到催化裂化反应产物；
    c)分离所得的催化裂化反应产物，得到包含异丁烷的液化气馏分和包含轻质芳烃的汽油馏分；以及
    d)任选地，从所述液化气馏分中回收异丁烷，和/或从所述汽油馏分中回收轻质芳烃。
  2. 根据权利要求1所述的方法，其中步骤c)中的所述分离还得到轻循环油馏分、重循环油馏分，和任选的油浆，并且所述方法进一步包括如下步骤：
    e)对步骤c)中得到的轻循环油馏分、重循环油馏分和任选的油浆的至少一部分进行加氢处理得到加氢尾油；以及
    f)将所得的加氢尾油的至少一部分再循环到所述催化裂化反应器中。
  3. 根据权利要求1或2所述的方法，其中步骤a)进一步包括对双环以上环烷烃含量不大于约25重量％的初始原料油进行预处理，以得到所述双环以上环烷烃含量大于约25重量％、优选大于约40重量％的催化裂化原料油；优选地，所述预处理包括芳烃抽提和/或加氢处理。
  4. 根据权利要求2或3所述的方法，其中所述步骤e)的加氢处理和/或作为所述预处理的加氢处理在如下条件下进行：氢分压为约6.0-30.0MPa，优选约8-20MPa；反应温度为约300-450℃，优选约330-430℃；液相体积空速为约0.1-10.0h -1，优选约0.2-5h -1；氢油体积比为约300-3000Nm 3/m 3，优选约500-2500Nm 3/m 3。
  5. 根据权利要求2-4中任一项所述的方法，其中所述步骤e)的加氢处理和/或作为所述预处理的加氢处理在加氢处理催化剂存在下进行，所述加氢处理催化剂包括加氢活性组分和载体，所述加氢活性组 分优选选自第VIB族非贵金属、第VIII族非贵金属，和它们的组合，所述载体优选选自氧化铝、二氧化硅、无定型硅铝，和它们的组合。
  6. 根据权利要求3所述的方法，其中所述芳烃抽提在如下条件下进行：温度为约50-70℃，溶剂与原料重量比为约0.5-2，溶剂选自糠醛、二甲亚砜、二甲基甲酰胺、单乙醇胺、乙二醇、1，2-丙二醇，和它们的组合。
  7. 根据前述权利要求中任一项所述的方法，其中所述催化裂化原料油或者所述初始原料油选自深度加氢轻循环油、延迟焦化装置的重馏分油、催化裂化轻循环油、催化裂化重循环油、催化裂化重馏分油、油浆、加氢裂化柴油、渣油加氢裂化柴油、蜡油加氢裂化柴油、生物柴油、页岩油柴油馏分、煤液化柴油馏分、常压塔顶油、常压塔抽出的馏分油、直馏减压蜡油、加氢蜡油、焦化蜡油、脱沥青油、抽出油、抽余油、常压渣油、减压渣油、上述原料油经加氢得到的加氢尾油，和它们的组合。
  8. 根据前述权利要求中任一项所述的方法，其中：
    所述第一催化裂化反应在如下条件下进行：反应温度为约520-620℃，优选约530-600℃；反应时间为约0.5-3.0秒，优选约0.8-2.0秒；剂油重量比为约3∶1至约15∶1，优选约4∶1至约12∶1；
    所述第二催化裂化反应在如下条件下进行：反应温度为约480-600℃，优选约500-550℃；反应时间为约2-30秒，优选约3-15秒；剂油重量比为约3∶1至约18∶1，优选约4∶1至约15∶1。
  9. 根据前述权利要求中任一项所述的方法，其中所述催化裂化反应器为等直径提升管反应器、等线速提升管反应器、变径提升管反应器、流化床反应器、或等直径提升管与流化床构成的复合反应器。
  10. 根据前述权利要求中任一项所述的方法，其中所述催化裂化反应器为变径提升管反应器，所述变径提升管反应器沿竖直方向由下至上依次设置有同轴且流体连通的预提升段、第一反应区、第二反应区和出口区，所述第一反应区的内径小于第二反应区，所述第二反应区的内径大于出口区，所述催化裂化催化剂送入所述预提升段中，所述催化裂化原料油送入所述第一反应区的下部，所述第一催化裂化反应在第一反应区中进行，所述第二催化裂化反应在第二反应区中进行。
  11. 根据权利要求10所述的方法，其中所述第一反应区和第二反 应区的结合区域设置有至少一个用于输入激冷介质的激冷介质入口；和/或
    所述第二反应区中设置有取热器，所述取热器的高度占第二反应区高度的约50-90％。
  12. 根据权利要求11所述的方法，其中所述激冷介质选自冷激剂、冷却的再生催化剂、冷却的半再生催化剂、新鲜催化剂，和它们的组合；所述冷激剂选自液化气、粗汽油、稳定汽油、轻循环油、重循环油、水，和它们的组合。
  13. 根据前述权利要求中任一项所述的方法，其中所述催化裂化催化剂包括裂化活性组分和载体；以干基计并以所述裂化活性组分的重量为基准，所述裂化活性组分包括约0-100重量％、优选约10-90重量％的FAU型沸石和约0-100重量％、优选约10-90重量％的五元环结构沸石，其中所述FAU型沸石和五元环结构沸石的总量为100重量％；所述FAU型沸石优选选自Y型沸石、HY型沸石、超稳Y型沸石，和它们的组合，所述五元环结构沸石优选选自ZSM-5系列沸石、高硅沸石、镁碱沸石，和它们的组合，所述五元环结构沸石任选地含有稀土和/或磷。
  14. 根据前述权利要求中任一项所述的方法，其中在步骤d)中通过抽提精制从所述汽油馏分中回收轻质芳烃，所述抽提精制优选在如下条件下进行：抽提溶剂选自环丁砜、二甲亚砜、N-甲酰基吗啉、四甘醇、三甘醇、N-甲基吡啶烷酮，和它们的组合，温度为约50-110℃，抽提溶剂与汽油馏分的重量比为约2-6。
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Families Citing this family (5)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
EP3744816A1 (en) * 2019-05-27 2020-12-02 Indian Oil Corporation Limited Process for the production of petrochemical feedstock and high octane gasoline from middle distillates
CN112538384B (zh) * 2019-09-20 2022-04-08 中国石油化工股份有限公司 一种多产异丁烷和轻质芳烃的加氢处理-催化裂化组合工艺方法
CN112745886B (zh) * 2019-10-30 2023-05-05 中国石油化工股份有限公司 一种增产低碳烯烃和轻质芳烃的方法和系统
CN113549470B (zh) * 2020-04-24 2023-02-03 中国石油化工股份有限公司 一种催化裂化方法
CN113897216B (zh) * 2020-06-22 2023-06-09 中国石油化工股份有限公司 一种催化裂解的方法和系统

Citations (2)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN1232069A (zh) * 1999-04-23 1999-10-20 中国石油化工集团公司 一种制取异丁烷和富含异构烷烃汽油的催化转化方法
CN104711022A (zh) * 2013-12-16 2015-06-17 中国石油化工股份有限公司 一种增加汽油收率生产超低硫汽油的两段催化汽油改质方法

Family Cites Families (20)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US2279550A (en) * 1939-03-31 1942-04-14 Universal Oil Prod Co Treatment of cracking stocks
US3684694A (en) * 1970-05-26 1972-08-15 Mobil Oil Corp Method of controlling interaction between pretreating and hydrocracking stages
US5154818A (en) 1990-05-24 1992-10-13 Mobil Oil Corporation Multiple zone catalytic cracking of hydrocarbons
US5582714A (en) * 1995-03-20 1996-12-10 Uop Process for the removal of sulfur from petroleum fractions
JPH08269464A (ja) 1995-04-03 1996-10-15 Japan Energy Corp 炭化水素油の接触分解方法
US5685972A (en) 1995-07-14 1997-11-11 Timken; Hye Kyung C. Production of benzene, toluene, and xylene (BTX) from FCC naphtha
US6210561B1 (en) 1996-08-15 2001-04-03 Exxon Chemical Patents Inc. Steam cracking of hydrotreated and hydrogenated hydrocarbon feeds
CN1081222C (zh) * 1999-06-23 2002-03-20 中国石油化工集团公司 一种降低液化气和汽油中烯烃含量的催化转化方法
US20060231459A1 (en) * 2005-03-28 2006-10-19 Swan George A Iii FCC process combining molecular separation with staged conversion
US20080078695A1 (en) * 2006-09-29 2008-04-03 Marathon Petroleum Company, Llc Method and apparatus for controlling catalytic cracking by near-infrared spectroscopy
WO2009089681A1 (fr) * 2007-12-20 2009-07-23 China Petroleum & Chemical Corporation Procédé intégré amélioré d'hydrogénation et de craquage catalytique d'hydrocarbures liquides
JP5919586B2 (ja) * 2010-01-20 2016-05-18 Ｊｘエネルギー株式会社 炭化水素製造用触媒および炭化水素の製造方法
CN102337154B (zh) * 2010-07-22 2014-12-03 中国石油化工股份有限公司 一种生产丙烯和轻芳烃的催化转化方法
JP5947797B2 (ja) 2010-09-27 2016-07-06 中国石油化工股▲ふん▼有限公司 生成物分布を改善するための触媒変成処理
US9127213B2 (en) * 2011-11-11 2015-09-08 Chevron U.S.A. Inc. Method for predicting catalyst performance
CN104178212B (zh) 2013-05-20 2016-07-27 神华集团有限责任公司 一种煤焦油加氢提质方法
CN104418686B (zh) * 2013-08-30 2017-03-01 中国石油化工股份有限公司 一种生产低碳烯烃和轻芳烃的催化转化方法
CN104560166B (zh) 2013-10-28 2016-05-25 中国石油化工股份有限公司 一种石油烃生产高辛烷值汽油的催化转化方法
CN105567298B (zh) * 2014-10-09 2017-08-29 中国石油化工股份有限公司 利用重碳九改善催化裂化产物分布的方法
WO2017004355A1 (en) * 2015-06-30 2017-01-05 Exxonmobil Research And Engineering Company Fuel production from catalytic slurry oil

Patent Citations (2)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN1232069A (zh) * 1999-04-23 1999-10-20 中国石油化工集团公司 一种制取异丁烷和富含异构烷烃汽油的催化转化方法
CN104711022A (zh) * 2013-12-16 2015-06-17 中国石油化工股份有限公司 一种增加汽油收率生产超低硫汽油的两段催化汽油改质方法

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