WO2022089575A1 - 含烃原料油催化裂解生产低碳烯烃和btx的方法及装置 - Google Patents

含烃原料油催化裂解生产低碳烯烃和btx的方法及装置 Download PDF

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WO2022089575A1
WO2022089575A1 PCT/CN2021/127339 CN2021127339W WO2022089575A1 WO 2022089575 A1 WO2022089575 A1 WO 2022089575A1 CN 2021127339 W CN2021127339 W CN 2021127339W WO 2022089575 A1 WO2022089575 A1 WO 2022089575A1
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catalyst
oil
gas
reaction
light
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PCT/CN2021/127339
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English (en)
French (fr)
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龚剑洪
汪燮卿
杨超
魏晓丽
朱根权
马文明
陈昀
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中国石油化工股份有限公司
中国石油化工股份有限公司石油化工科学研究院
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Priority to KR1020237017715A priority Critical patent/KR20230093312A/ko
Priority to US18/250,767 priority patent/US20230392087A1/en
Priority to CN202180074180.6A priority patent/CN116601268A/zh
Priority to EP21885301.8A priority patent/EP4219664A1/en
Publication of WO2022089575A1 publication Critical patent/WO2022089575A1/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
    • 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
    • C10G55/00—Treatment of hydrocarbon oils, in the absence of hydrogen, by at least one refining process and at least one cracking process
    • C10G55/02—Treatment of hydrocarbon oils, in the absence of hydrogen, by at least one refining process and at least one cracking process plural serial stages only
    • C10G55/06—Treatment of hydrocarbon oils, in the absence of hydrogen, by at least one refining process and at least one cracking process plural serial stages only including at least one catalytic cracking step
    • 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
    • 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
    • C10G51/00—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes only
    • C10G51/06—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes only plural parallel stages 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/10—Feedstock materials
    • C10G2300/1037—Hydrocarbon fractions
    • C10G2300/104—Light gasoline having a boiling range of about 20 - 100 °C
    • 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
    • C10G2300/1037—Hydrocarbon fractions
    • C10G2300/1044—Heavy gasoline or naphtha having a boiling range of about 100 - 180 °C
    • 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
    • C10G2300/70—Catalyst aspects
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
    • C10G2400/20—C2-C4 olefins
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
    • C10G2400/30—Aromatics

Definitions

  • the present application relates to petroleum refining and petrochemical processing processes, in particular, to a method and device for producing light olefins and BTX by catalytic cracking of full-cut hydrocarbon-containing feedstock oil.
  • crude oil is pretreated by solvent deasphalting or hydrorefining, and then directly enters the steam cracking unit to produce chemical materials, but this method is generally limited to light crude oil;
  • the second is to maximize the production of heavy naphtha through the hydrocracking of various crude oil fractions, and then maximize the production of aromatics through the reforming unit;
  • the third is that the light fractions of crude oil enter the steam cracking unit, and the heavy fractions enter the catalytic cracking unit to maximize the production of aromatics.
  • Production of light olefins The above three methods have all been industrialized, and the yield of chemicals is between 35% and 55%.
  • the configuration of the existing chemical refinery mainly relies on the combination of multiple sets of core devices such as steam cracking, reforming, hydrorefining, hydrocracking, and catalytic cracking.
  • the catalytic cracking process has its unique advantages in the production of chemical materials and the adaptability of raw materials, and can produce propylene, ethylene and BTX at the same time.
  • Chinese patent CN1978411B discloses a combined process method for producing small molecular olefins.
  • the catalytic cracking catalyst and the cracking raw materials are mixed and contacted in a reactor to separate the catalyst to be produced and the reaction oil and gas, wherein the catalyst to be produced is sent to the
  • the regenerator performs coke regeneration, and the regenerated hot catalyst is divided into two parts. One part of the regenerated hot catalyst is returned to the above reactor; the other part of the regenerated hot catalyst is first mixed with heavy petroleum hydrocarbons in another reactor.
  • the olefinic raw material rich in C4-C8 is mixed and contacted with the coked catalyst, and a catalytic cracking reaction occurs, and the reaction oil and gas of the catalyst to be separated is separated.
  • Coking and regeneration are carried out in the reactor; the reaction oil and gas are separated to obtain the target products of small molecular olefins such as propylene.
  • the method can convert olefin-rich light feedstocks into small-molecule olefin products such as propylene with high selectivity, while maintaining the thermal balance of the device itself.
  • Chinese patent CN102899078A discloses a catalytic cracking method for producing propylene.
  • the method is based on a combined reactor composed of double risers and a fluidized bed.
  • First, the heavy feedstock oil and the first catalyst are introduced into the first riser reactor for reaction.
  • the oil is separated into the separation system.
  • the cracked heavy oil is introduced into the second riser reactor to contact and react with the catalyst introduced into the second riser reactor, and the light hydrocarbons are introduced into the second riser reactor to contact with the mixture formed by the contact reaction of the cracked heavy oil and the second cracking catalyst , the light hydrocarbons include C4 hydrocarbons or gasoline fractions obtained by the product separation system.
  • the oil and gas reacted in the second riser reactor and the catalyst are introduced into the fluidized bed reactor for reaction.
  • the selective conversion of different feeds has a higher yield of propylene and butene.
  • Chinese patent CN101045667B discloses a combined catalytic conversion method for producing more low-carbon olefins.
  • the heavy oil feedstock is contacted with a regeneration catalyst and an optional coke catalyst in a descending tube reactor, and the separated low-carbon olefins are separated into At least a part of the remaining products are introduced into the riser reactor to contact with the regenerated catalyst.
  • the catalyst is introduced into the catalyst pre-lifting section of the descending pipe reactor, mixed with the regenerated catalyst entering the descending pipe reactor, and then contacted with the heavy oil feedstock.
  • the method adopts the form of a combined reactor in which the heavy oil raw material is reacted in a descending reactor, and the intermediate product olefin is reacted in a riser reactor, so as to improve the yield of light olefins.
  • Chinese patent CN109370644A discloses a method for producing light olefins and aromatics by catalytic cracking of crude oil.
  • the method divides crude oil into light and heavy components, and the cutting point is between 150°C and 300°C.
  • the reaction is carried out in the reaction zone, and the catalyst adopts aluminosilicate composed of silicon dioxide and aluminum oxide as the main component, including alkali metal oxides, alkaline earth metal oxides, titanium, iron oxides, vanadium and nickel oxides .
  • the method is based on the dense phase transported bed reactor for the catalytic cracking of heavy oil to generate light olefins, and is a solution proposed for the catalytic cracking of crude oil to generate light olefins.
  • the purpose of the present disclosure is to propose a method suitable for processing hydrocarbon-containing raw materials for catalytic cracking in view of the characteristics of different hydrocarbon compositions and different cutting temperatures of various hydrocarbon-containing raw materials, so as to maximize the use of hydrocarbon-containing raw materials to produce low Apparatus and method for carbene and BTX.
  • the present disclosure provides a method for producing light olefins and light aromatic hydrocarbons by catalytic cracking of hydrocarbon-containing feedstock oil, the method comprising the following steps:
  • the hydrocarbon-containing feed oil is cut into light distillate oil and heavy distillate oil, and the weight ratio (light distillate oil/heavy distillate oil) of the light distillate oil relative to the heavy distillate oil is X;
  • the light distillate oil and the first catalyst are introduced into the first descending reactor, and the first catalytic cracking is carried out to obtain the material after the first catalytic cracking;
  • Optional S2' introducing the material after the first catalytic cracking into the fluidized bed reactor for the second catalytic cracking to obtain the material after the second catalytic cracking;
  • the continuous catalyst, the heavy distillate oil and the second catalyst are introduced into the second upward reactor, the third catalytic cracking is carried out, and then the gas-solid separation is carried out to obtain the third reaction oil and gas and the third catalyst to be produced;
  • the continuous catalyst is at least a part of the first catalyst or at least a part of the second catalyst; the weight ratio of the second catalyst to the continuous catalyst (second catalyst/continuous catalyst) is R ;
  • T0 is the temperature (unit °C) at which the second strand of catalyst enters step S4, and T3 is the outlet temperature (unit °C) of the second upward reactor.
  • the outlet temperature T3 of the second upward reactor is 530-650°C, preferably 560-640°C, more preferably 580-630°C, still more preferably 600-630°C °C; and/or, the temperature T0 when the second catalyst enters step S4 is 690-750°C, preferably 700-740°C, more preferably 705-730°C, still more preferably 710-725°C.
  • step S1 the hydrocarbon-containing feedstock oil is cut into light distillate oil and heavy distillate oil at any temperature between the cutting point of 100-400°C, so that the light distillate oil is relatively
  • the weight ratio of heavy distillate oil (light distillate oil/heavy distillate oil) is described as X.
  • the conditions for the first catalytic cracking include: the outlet temperature of the first descending reactor is 610-720° C., and the gas The solid residence time is 0.1-3.0 seconds, and the agent-oil ratio is 15-80; and/or, in the fluidized bed reactor, the conditions for the second catalytic cracking include: the reaction in the fluidized bed reactor The temperature is 600-690° C., and the mass space velocity is 2-20h ⁇ 1 ; and/or, in the second upward reactor, the conditions for the third catalytic cracking include: the gas-solid residence time is 0.5-8 seconds , the agent oil ratio is 8-40.
  • the conditions for the first catalytic cracking include: the outlet temperature of the first descending reactor is 650-690° C., and the gas The solid residence time is 0.5-1.5 seconds, and the agent-oil ratio is 25-65; and/or, in the fluidized bed reactor, the conditions for the second catalytic cracking include: the reaction in the fluidized bed reactor The temperature is 640-670° C., and the mass space velocity is 4-12h ⁇ 1 ; and/or, in the second upward reactor, the conditions for the third catalytic cracking include: the gas-solid residence time is 1.5-5 seconds , the agent oil ratio is 10-30.
  • step S4 the continuous catalyst is first mixed with the second strand of catalyst, and then the subsequent catalytic cracking reaction is performed, and/or, when step S2' exists, in step S3 In the gas-solid separation, the separated catalyst is stripped to obtain a second catalyst to be produced; and/or, in step S4, the light olefin fraction from step S5 is prior to the heavy distillate oil and the The mixture of the second strand catalyst and the continuous catalyst is contacted for catalytic cracking; preferably, the light olefins are contacted with the mixture of the second strand catalyst and the continuous catalyst for 0.3-1.0 seconds prior to the heavy distillate oil for catalytic cracking, more preferably The light olefin fraction is contacted with the mixture of the second catalyst and the continuous catalyst for 0.4-0.8 seconds prior to the heavy distillate oil for catalytic cracking; and/or, the method has step S0 before step S1, wherein , the hydrocarbon-containing raw material oil is subjected to des
  • the method further includes: in the gas-solid separation in step S4, stripping the separated catalyst to obtain a third catalyst to be grown; and/or, separating the catalyst
  • Aromatic hydrocarbons, and the light olefin fraction is separated; and/or, when step S2' is not present, in step S5, the light olefin fraction is separated from any one of the first reaction oil and gas, the third reaction oil and gas, or a mixture of the two , and the light olefin fraction is returned to the second upward reactor in step S4; when step S2' exists, in step S5, from either the second reaction oil or the third reaction oil or a mixture of the two The light olefin fraction is separated out, and the light olefin fraction is returned to the fluidized bed reactor of step S2'.
  • the hydrocarbon-containing feedstock oil is one or more of crude oil, coal liquefied oil, synthetic oil, oil sand oil, shale oil, tight oil and animal and vegetable oils and fats.
  • the first strand catalyst and the second strand catalyst each independently include an active component and a carrier, and the active component is selected from ultrastable Y containing or not containing rare earth.
  • the active component is selected from ultrastable Y containing or not containing rare earth.
  • the carrier is selected from alumina, silica, amorphous silica alumina, zirconia, titania, At least one of boron oxide and alkaline earth metal oxide.
  • the first strand of catalyst and the second strand of catalyst each independently include a regenerated catalyst, preferably the first strand of catalyst and the second strand of catalyst are regenerated catalysts, and/ Or, the whole of the first to-be-grown catalyst or the entirety of the second to-be-grown catalyst is used as a continuous catalyst.
  • the present disclosure also provides a device for producing light olefins and light aromatics by catalytic cracking of hydrocarbon-containing feedstock oil, the device comprising the following units:
  • a hydrocarbon-containing feed oil cutting unit in which the hydrocarbon-containing feed oil is cut into light distillate oil and heavy distillate oil, so that the weight ratio of the light distillate oil to the heavy distillate oil (light distillate oil/heavy distillate oil) is X,
  • the first downward reaction unit, the light distillate oil and the first catalyst are introduced from the top of the reaction unit, the first catalytic cracking is carried out, and the material after the first catalytic cracking is obtained below the reaction unit;
  • Optional fluidized bed reaction unit wherein, the material after the first catalytic cracking is introduced, and the second catalytic cracking is performed to obtain the material after the second catalytic cracking;
  • the first gas-solid separation unit wherein the material after the first catalytic cracking is introduced for gas-solid separation to obtain the first reaction oil and gas and the first catalyst to be produced, or the material after the second catalytic cracking is introduced for gas-solid separation Separation to obtain the second reaction oil and gas and the second catalyst to be generated;
  • the continuous catalyst, the second catalyst and the heavy distillate oil are introduced from the bottom of the reaction unit to carry out the third catalytic cracking, and the material after the third catalytic cracking is obtained above the reaction unit, so
  • the continuous catalyst is at least a part of the first catalyst to grow or at least a part of the second catalyst to grow, and the weight ratio of the second catalyst to the continuous catalyst (second catalyst/continuous catalyst) is R,
  • the second gas-solid separation unit wherein the material after the third catalytic cracking is introduced for gas-solid separation to obtain the third reaction oil and gas and the third catalyst to be produced;
  • a separation unit in which any one of the first reaction oil, the second reaction oil and the third reaction oil or a mixture of the first reaction oil and the third reaction oil or the second reaction oil and the third reaction oil is introduced
  • the mixture of oil and gas separates light olefins and light aromatics, and separates light olefin fractions, and returns the light olefin fractions to the second upward reaction unit or the fluidized bed reaction unit;
  • T0 is the temperature (unit °C) at which the second strand of catalyst enters the second upward reaction unit
  • T3 is the outlet temperature (unit °C) of the second upward reaction unit.
  • a regeneration unit is further included, wherein, the third to-be-grown catalyst and the optional first or second to-be-grown catalyst that do not enter the second upward reactor are introduced, Coke regeneration is carried out at a temperature of 690-750°C, preferably 700-740°C, more preferably 705-730°C, and still more preferably 710-725°C to obtain a regenerated catalyst.
  • the first gas-solid separation unit when the device includes a fluidized bed reaction unit, the first gas-solid separation unit further includes a stripping unit, wherein the catalyst obtained by gas-solid separation is stripped to obtain the first gas-solid separation unit. Second, the catalyst to be produced.
  • the second gas-solid separation unit further includes a stripping unit, wherein the catalyst obtained by gas-solid separation is stripped to obtain a third to-be-grown catalyst.
  • the device further includes a dehydration and desalination unit, wherein the hydrocarbon-containing feedstock oil is subjected to desalination and dehydration treatment, and the obtained dehydrated and desalted hydrocarbon-containing feedstock oil is introduced into the hydrocarbon-containing feedstock oil cutting unit for processing. cut.
  • the position where the continuous catalyst and the second catalyst are introduced is upstream of the feed port of the light olefin fraction.
  • the feed port of the light olefin fraction from the separation unit is upstream of the heavy distillate feed port.
  • the hydrocarbon-containing feedstock oil is cut into two parts, light distillate oil and heavy distillate oil, according to the hydrocarbon composition characteristics and cracking reaction characteristics of different fractions of the hydrocarbon-containing feedstock oil, and the light distillate oil is divided into two parts.
  • cracking is carried out at high temperature and short residence time, which can produce light olefins and BTX with high selectivity, and at the same time can significantly reduce the generation of methane.
  • the production of light olefins and BTX can be maximized by using an up-flow reactor.
  • the light olefins in the material after catalytic cracking can be further converted, and the production of light olefins can be maximized.
  • the residence time of the light distillate oil is short, the coke produced by the reaction is low, and the yield of light olefins and BTX is high; in addition, in the fluidized bed reactor, the light olefin fraction is further processed by transform. Therefore, the first catalyst to be grown out of the first descending reactor or the second catalyst to be grown out of the fluidized bed reactor still has a relatively high activity, and the catalyst is loaded with carbon deposits, and the catalyst When used in the catalytic cracking of the heavy distillate oil in the second upward reactor, the yield of light olefins can be improved, and the generation of dry gas and coke can be suppressed.
  • the hydrocarbon content can be determined according to the specific relationship.
  • the cutting ratio is flexibly adjusted, and accordingly, the weight ratio of the second catalyst and the continuous catalyst is adjusted, so that in the second upward reactor, the catalyst activity and the composition of heavy distillate oil are better. Matching can significantly reduce the yield of by-products such as dry gas and coke while maximizing the production of light olefins and BTX.
  • the method for producing light olefins and BTX by catalytic cracking of hydrocarbon-containing feedstock oil provided by the present disclosure can significantly improve the yield and device economy of light olefins and light aromatics.
  • FIG. 1 is a schematic diagram of one embodiment of the apparatus of the present disclosure.
  • FIG. 2 is a schematic diagram of another embodiment of the apparatus of the present disclosure.
  • any specific numerical value disclosed herein, including the endpoints of a numerical range, is not limited to the precise value of the numerical value, but is to be understood to encompass values approximating the precise value, such as within ⁇ 5% of the precise value. all possible values. And, for the disclosed numerical range, between the endpoint values of the range, between the endpoint values and the specific point values in the range, and between the specific point values, one or more new values can be obtained in any combination. Numerical ranges, these new numerical ranges should also be considered to be specifically disclosed herein.
  • any matter or matter not mentioned is directly applicable to those known in the art without any change.
  • any embodiment described herein can be freely combined with one or more other embodiments described herein, and the technical solutions or technical ideas formed thereby are regarded as a part of the original disclosure or original record of the present disclosure, and should not be It is considered to be new content not disclosed or anticipated herein, unless a person skilled in the art considers that the combination is obviously unreasonable.
  • the present disclosure provides a method for producing light olefins and light aromatic hydrocarbons by catalytic cracking of hydrocarbon-containing feedstock oil, the method comprising the following steps:
  • the hydrocarbon-containing feed oil is cut into light distillate oil and heavy distillate oil, and the weight ratio (light distillate oil/heavy distillate oil) of the light distillate oil relative to the heavy distillate oil is X;
  • the light distillate oil and the first catalyst are introduced into the first descending reactor, and the first catalytic cracking is carried out to obtain the material after the first catalytic cracking;
  • Optional S2' introducing the material after the first catalytic cracking into the fluidized bed reactor for the second catalytic cracking to obtain the material after the second catalytic cracking;
  • the continuous catalyst, the heavy distillate oil and the second catalyst are introduced into the second upward reactor, the third catalytic cracking is carried out, and then the gas-solid separation is carried out to obtain the third reaction oil and gas and the third catalyst to be produced;
  • the continuous catalyst is at least a part of the first catalyst or at least a part of the second catalyst; the weight ratio of the second catalyst to the continuous catalyst (second catalyst/continuous catalyst) is R ;
  • T0 is the temperature (unit °C) at which the second strand of catalyst enters step S4, and T3 is the outlet temperature (unit °C) of the second upward reactor.
  • reaction oil and gas any one or a mixture of two or more of the first reaction oil and gas, the second reaction oil and gas, and the third reaction oil and gas are sometimes referred to as reaction oil and gas for short.
  • lower olefins refer to ethylene, propylene, butene and isomers thereof.
  • Light aromatics refers to BTX, ie benzene, toluene and xylene.
  • light olefins can be separated from dry gas, C3 fraction and C4 fraction; light aromatics can be separated from light gasoline and heavy gasoline.
  • the C3 fraction refers to hydrocarbons with 3 carbons in the reaction oil and gas, including propane and propylene; the C4 fraction refers to the hydrocarbons with 4 carbons in the reaction oil and gas, including butane, butene and its isoforms.
  • light gasoline refers to the whole fraction or part of the fraction whose distillation range is in the range of 30-90°C in the reaction oil and gas, wherein "partial fraction” refers to the fraction whose distillation range is part of the temperature range between 30-90°C (For example, fractions with a distillation range of 30-60°C or 40-60°C or 60-90°C, etc.); heavy gasoline refers to the fractions with a distillation range in the range of 30-200°C in the reaction oil and gas except for the fractions other than light gasoline. .
  • light distillate oil and heavy distillate oil refer to the light distillate oil after cutting the hydrocarbon-containing feedstock oil at a certain cutting temperature
  • the cut light distillate oil is called light distillate oil
  • the remaining part is called heavy distillate oil.
  • Those skilled in the art can cut the hydrocarbon-containing feedstock oil according to methods known in the art (including but not limited to fractionation, distillation, etc.) as required, as long as the weight ratio of the light distillate oil to the heavy distillate oil is made ( Light distillate oil/heavy distillate oil) is X, and X may satisfy the following relational expression of the present disclosure.
  • X is selected from the group consisting of 0.1, 0.2, 0.3, 0.4, 0.5, 0.6, 0.7, 0.8, 0.9, 1.0, 2.0, 3.0, 4.0, 5.0, 6.0, 7.0, 8.0, 9.0, 10.0 between any two value ranges. In one embodiment of the present disclosure, X is 0.1-2.0, preferably 0.12-1.0, further preferably 0.15-0.6.
  • step S1 the hydrocarbon-containing feedstock oil is cut into light distillate oil and heavy distillate oil at any temperature between the cutting point of 100-400° C., so that the light distillate oil is relative to the
  • the weight ratio of heavy distillate (light distillate/heavy distillate) is X.
  • the cutting point is, for example, 150°C, 160°C, 170°C, 180°C, 190°C, 200°C, 210°C, 220°C, 230°C, 240°C, 250°C, 260°C, 270°C °C, 280 °C, 290 °C, 300 °C, 310 °C, 320 °C, 330 °C, 340 °C, 350 °C, 360 °C, 370 °C, 380 °C, 390 °C, 400 °C.
  • the hydrocarbon-containing feedstock oil can be various types of feedstock oils known in the art (in the present invention, the hydrocarbon-containing feedstock oil is sometimes referred to as feedstock oil for short), for example, it can be crude oil, coal liquefied oil, synthetic oil , oil sands oil, shale oil, tight oil and animal and vegetable oils and fats, or a mixture of two or more thereof, or their respective partial fractions, their respective heavy fractions of hydro-upgraded oil.
  • the hydrocarbon-containing feedstock oil is preferably crude oil, a partial fraction of crude oil, or hydro-upgraded oil of heavy oil from crude oil.
  • the "partial fraction" can be obtained by subjecting the feedstock oil to conventional treatments in the art, including but not limited to atmospheric distillation, vacuum distillation, and the like. Those skilled in the art can determine the manner of this conventional treatment as needed.
  • crude oil can be used as the hydrocarbon-containing feedstock oil of the present disclosure, and the crude oil can also be subjected to atmospheric distillation or vacuum distillation as required, and the remaining fractions (partial fractions of crude oil) after extracting part of the fractions can be used as The hydrocarbon-containing feedstock oil of the present disclosure, or a product obtained by hydro-upgrading heavy oil derived from crude oil (hydro-upgraded oil of heavy oil) as required, is used as the hydrocarbon-containing feedstock oil of the present disclosure.
  • hydroupgrading includes, but is not limited to, hydrodesulfurization, hydrodenitrogenation, hydrodemetallization, hydrosaturation, and the like.
  • the method includes step S0 before step S1, wherein the hydrocarbon-containing feedstock oil is subjected to desalination and dehydration treatment, and the obtained dehydrated and desalted hydrocarbon-containing feedstock oil is introduced into step S1 for cutting.
  • the conditions for the first catalytic cracking include: the outlet temperature of the first down-type reactor is 610-720° C., preferably 650° C. -690°C.
  • the conditions for the first catalytic cracking further include: the gas-solid residence time is 0.1-3.0 seconds, preferably 0.5-1.5 seconds.
  • the catalyst-oil ratio between the catalyst and the light distillate oil can be the catalyst-oil ratio (in terms of the weight ratio of catalyst/light distillate oil) commonly used in catalytic cracking, for example, it can be 15-80. , preferably 25-65.
  • the light distillate oil and the first catalyst are introduced into the first descending reactor, as long as the light distillate and the first catalyst are introduced at the upper end of the first descending reactor, that is, Can.
  • the light distillate oil and the first catalyst are respectively introduced from different feed ports of the first descending reactor.
  • the first strand catalyst is not limited, and may be a catalyst known in the art that can be used for catalytic cracking of crude oil.
  • the first strand catalyst may include an active component and a carrier, and the active component is selected from ultra-stable Y-type zeolite with or without rare earth, ZSM-5 series zeolite, high silicon with a five-membered ring structure At least one of zeolite and beta zeolite.
  • the support is selected from at least one of alumina, silica, amorphous silica-alumina, zirconia, titania, boron oxide and alkaline earth metal oxides.
  • the structure of the first descending reactor is not particularly limited, as long as the upper part of the reactor can be fed and the lower part can be discharged, for example, it can be of equal diameter or variable diameter Downpipe reactor.
  • the outlet temperature of the first descending reactor reflects the reaction temperature in the reactor.
  • the catalysis of light distillate oil in the first descending reactor can be adjusted by adjusting the temperature of the first catalyst, the gas-solid residence time in the reactor, the outlet temperature of the first descending reactor, etc. degree of cracking.
  • the first strand of catalyst is fresh catalyst.
  • the first stream of catalyst comprises regenerated catalyst from a regenerator.
  • the first stream of catalyst is a regenerated catalyst from a regenerator.
  • the temperature of the first stream of catalyst entering the down-type reactor is not particularly limited, as long as it can be catalytically cracked when contacted with light distillate oil and satisfies the above-mentioned first catalytic cracking conditions of the present disclosure.
  • the first stream of catalyst is directly fed from the regenerator via the first stream of catalyst (regenerated catalyst) delivery pipe, due to the delivery pipe between the regenerator and the first descending reactor Therefore, the temperature of the first stream of catalyst can be regarded as the temperature of the regenerator or the temperature of the regenerated catalyst as it leaves the regenerator (the temperature of the outlet of the regenerator).
  • the temperature of the first stream of catalyst entering the down-type reactor is the temperature of the regenerator or the temperature when the regenerated catalyst leaves the regenerator (the temperature at the outlet of the regenerator), which can usually be 690-750°C, Preferably it is 700-740 degreeC, More preferably, it is 705-730 degreeC, More preferably, it is 710-725 degreeC.
  • the catalyst from the regenerator may be further heated or cooled before being fed into the first down-type reactor.
  • fresh catalyst can be heated to the desired temperature before being introduced into the first descending reactor; thereafter, the regenerated catalyst from the regenerator can be used directly.
  • the first stream of catalyst is preferably fed directly from the regenerator without further heating or cooling.
  • the light distillate oil when the light distillate oil is introduced into the first descending reactor, the light distillate oil may also be preheated first as required.
  • the temperature of the preheated light distillate oil is, for example, 30-100°C.
  • steam atomization of the light distillate oil can also be performed first, and then the light distillate oil can be introduced into the first descending reactor by using the steam as a carrier.
  • the first catalytically cracked material includes the first reacted oil and gas obtained by catalytically cracking the light distillate oil, and the first coked (carbonized) first prepared catalyst.
  • the first to-be-grown catalyst still has relatively high activity, and when the catalyst is loaded with carbon deposits, when it is introduced into the subsequent second ascending reactor as a continuous catalyst, it is helpful for the catalytic cracking of heavy distillate oil, Improve the yield of light olefins and inhibit the generation of dry gas and coke.
  • step S3 the material after the first catalytic cracking is subjected to gas-solid separation to obtain the first reacted oil and gas and the first catalyst to be produced.
  • the method of gas-solid separation is not particularly limited, and methods known in the art can be used, for example, a settler and a cyclone separator are used to separate the catalyst and the first reaction oil and gas.
  • the first reaction oil and gas are separated to obtain dry gas, C3 fraction, C4 fraction, light gasoline, heavy gasoline, diesel oil and oil slurry, and light olefins and light aromatics are separated therefrom and separated.
  • a light olefin fraction is produced.
  • the C4 fraction and/or light gasoline is the light olefin fraction.
  • the first reaction oil and gas is introduced into a fractionation device or a gas separation device for fractionation, so as to achieve the above separation.
  • the light olefin fraction is introduced into the second upstream reactor in step S4 described below.
  • At least a portion of the first as-grown catalyst is introduced as a continuous catalyst into the second ascending reactor described below.
  • the first to-be-grown catalyst that does not enter the second ascending reactor described below is introduced into a regeneration step, in which regeneration of the catalyst is performed.
  • the whole of the first growing catalyst is introduced as a continuous catalyst into the second ascending reactor described below, where the amount of the first growing catalyst as a continuous catalyst substantially corresponds to the amount of the first stream of catalyst.
  • the material after the first catalytic cracking is subjected to gas-solid separation, and the separated catalyst is further stripped to remove the adsorbed hydrocarbon products to obtain the first catalyst to be produced.
  • step S2' may also be included after step S2 and before step S3, wherein the material after the first catalytic cracking is introduced into a fluidized bed reactor for second catalytic cracking to obtain The material after the second catalytic cracking, thereby, the light olefin fraction can be further converted, and the production of light olefins can be maximized.
  • fluidized bed reactor is also referred to as “fluidized reactor”, and its catalyst density is between 150-450 kg/ m3 .
  • the conditions for the second catalytic cracking include: the reaction temperature in the fluidized bed reactor is 600-690° C., preferably 640-670° C. °C.
  • the conditions for the second catalytic cracking further include: the mass space velocity is 2-20h -1 , preferably 4-12h -1 .
  • the material after the first catalytic cracking into the fluidized bed for catalytic cracking without introducing a new catalyst.
  • no additional heat source is applied to the fluidized bed, and the heat of the first catalytically cracked material can be directly utilized.
  • the introduced first catalytically cracked material includes the first reacted oil and gas obtained by catalytically cracking the light distillate oil, and the first coked (carbonized) first prepared catalyst.
  • the first to-be-grown catalyst still has relatively high activity, and the degree of catalytic cracking can be further deepened in the fluidized bed reactor, and the light olefin fraction can be further converted into light olefins.
  • a light olefin fraction is separated from the reaction oil and gas of the present disclosure, and the light olefin fraction is returned to the fluidized bed reactor for further conversion into light olefins. More specifically, the reaction oil and gas are separated to obtain dry gas, C3 fraction, C4 fraction, light gasoline, heavy gasoline, diesel oil and oil slurry, from which light olefins and light aromatics are separated, and light olefin fractions are separated. Wherein, the C4 fraction and/or light gasoline is the light olefin fraction.
  • the reaction oil and gas are introduced into a fractionation unit or gas separation unit to achieve the above separation.
  • the material after the second catalytic cracking includes the second reacted oil and gas and the second catalyst to be produced.
  • the second to-be-grown catalyst still has relatively high activity, and when the catalyst is loaded with carbon deposits, when it is introduced into the subsequent second upward reactor as a continuous catalyst, it is helpful for the catalytic cracking of heavy distillate oil, Improve the yield of light olefins and inhibit the generation of dry gas and coke.
  • step S3 the material after the second catalytic cracking is subjected to gas-solid separation to obtain the second reacted oil and gas and the second catalyst to be produced.
  • the method of gas-solid separation is not particularly limited, and methods known in the art can be used, such as a settler and a cyclone to separate the catalyst and the second reaction oil and gas.
  • the second reaction oil and gas are separated to obtain dry gas, C3 fraction, C4 fraction, light gasoline, heavy gasoline, diesel oil, and oil slurry, from which light olefins and light aromatics are separated and separated.
  • a light olefin fraction is produced.
  • C4 fractions and/or light gasoline are light olefin fractions.
  • the second reaction oil and gas is introduced into a fractionation device or a gas separation device to achieve the above separation.
  • the material after the second catalytic cracking is subjected to gas-solid separation, and the separated catalyst is further stripped to remove the adsorbed hydrocarbon products to obtain the second catalyst to be grown.
  • at least a portion of the second as-grown catalyst is introduced into the second ascending reactor described below as a continuous catalyst.
  • the second to-be-grown catalyst that does not enter the second ascending reactor described below is introduced into a regeneration step, in which regeneration of the catalyst is performed.
  • the whole of the second as-grown catalyst is introduced as a continuous catalyst into the second ascending reactor described below, where the amount of the second as-grown catalyst as a continuous catalyst substantially corresponds to the amount of the first stream of catalyst.
  • step S4 the continuous catalyst, the heavy distillate oil and the second catalyst are introduced into the second ascending reactor, the third catalytic cracking is performed, and then the gas-solid separation is performed to obtain the third reaction oil and gas and the third A growing catalyst; the continuous catalyst is at least a part of the first growing catalyst or at least a part of the second growing catalyst.
  • the conditions for the third catalytic cracking include: the outlet temperature T3 of the second upward reactor is 530-650° C., preferably 560-640° C. °C, more preferably 580-630 °C, still more preferably 600-630 °C.
  • the conditions of the third catalytic cracking further include: the gas-solid residence time is 0.5-8 seconds, preferably 1.5-5 seconds.
  • the catalyst-oil ratio of the catalyst to the heavy distillate oil can be the commonly used catalyst-oil ratio (in the weight ratio of catalyst/heavy distillate oil) in catalytic cracking, for example, it can be 8-40, It is preferably 10-30.
  • step S4 the continuous catalyst is first mixed with the second strand of catalyst, and then the subsequent catalytic cracking reaction is performed. More specifically, in one embodiment of the present disclosure, the continuous catalyst and the second stream of catalyst are independently fed to the bottom of the second upward reactor, mixed, and the mixed catalyst (below , sometimes referred to as catalyst mixture or mixed catalyst) for the catalytic cracking reaction in the second ascending reactor. In one embodiment of the present disclosure, after mixing the continuous catalyst with the second stream of catalyst in the bottom region of the second upward reactor, the mixed catalyst is lifted in the second upward reactor using a pre-lift medium , for the downstream catalytic cracking reaction. In one embodiment of the present disclosure, the pre-lift medium may be dry gas, water vapor, or a mixture thereof.
  • the second catalyst in step S4, is not limited, and may be a catalyst known in the art that can be used for catalytic cracking of crude oil.
  • the second strand catalyst includes an active component and a carrier, and the active component is selected from ultra-stable Y-type zeolite with or without rare earth, ZSM-5 series zeolite, high silica zeolite with a five-membered ring structure and at least one of beta zeolite.
  • the support is selected from at least one of alumina, silica, amorphous silica-alumina, zirconia, titania, boron oxide and alkaline earth metal oxides.
  • the structure of the second upward reactor is not particularly limited, as long as the material can be fed from the bottom and discharged from the top, for example, it can be of equal diameter or variable diameter.
  • the second strand of catalyst is fresh catalyst.
  • the second stream of catalyst comprises regenerated catalyst from a regenerator.
  • the second stream of catalyst is regenerated catalyst from a regenerator.
  • the catalyst needs to be preheated, so that the temperature of the fresh catalyst when entering step S4 satisfies the relational expression of the present disclosure.
  • the second stream of catalyst is a regenerated catalyst from a regenerator.
  • the weight ratio of the second strand catalyst to the continuous catalyst is R, and the R and X are made to satisfy the following relational formula:
  • T0 is the temperature (unit °C) at which the second strand of catalyst enters step S4, and T3 is the outlet temperature (unit °C) of the second upward reactor.
  • the inventors of the present disclosure have surprisingly found that by making the cut ratio (weight ratio of light distillate oil/heavy distillate oil) of the hydrocarbon-containing feedstock oil in step S1 of light distillate oil and heavy distillate oil to the second strand catalyst and
  • the weight ratio of the continuous catalyst (second catalyst/continuous catalyst) satisfies the above relationship, which can make the hydrocarbon-containing feedstock oil composition, cutting ratio, catalyst activity (especially the catalyst activity in the second up-flow reactor) more stable.
  • a good match can significantly reduce the yield of dry gas and coke while maximizing the production of light olefins and BTX.
  • the inventors of the present disclosure speculate that the catalyst that is the continuous catalyst comes from the first or second growing catalyst, due to the generation in the first descending reactor and the fluidized bed reactor.
  • the coke is low, so that the first catalyst and the second catalyst have high catalytic activity, and a certain amount of coke is loaded.
  • This catalyst is in a certain proportion with the second catalyst (fresh catalyst or from the regenerator).
  • the regenerated catalyst) is mixed, and the mixing ratio is matched with the cutting ratio of the hydrocarbon-containing feedstock oil to satisfy the above-mentioned relational formula of the present disclosure, and the obtained catalyst maintains excellent catalytic activity without causing excessive catalyst activity.
  • the ratio of cutting the hydrocarbon-containing feedstock oil into light distillate oil and heavy distillate oil and the mixing ratio of the continuous catalyst and the second catalyst satisfy a specific relationship, and the first step can be adjusted according to the composition of the hydrocarbon-containing feedstock oil, the cutting ratio, etc.
  • the activity of the mixed catalyst in the two ascending reactors maximizes the yield of light olefins and BTX from heavy distillates.
  • (4.84 ⁇ T0-3340)/(780+5 ⁇ T0-6 ⁇ T3) is greater than 0.
  • T0 is greater than T3.
  • T0 is the temperature at which the second strand of catalyst enters step S4. Specifically, it refers to the temperature at which the second stream of catalyst (fresh catalyst or regenerated catalyst) enters the second ascending reactor, ie the temperature at which it enters the bottom of the second ascending reactor, before mixing with the continuous catalyst.
  • the temperature of the regenerator or the temperature of the catalyst when the regenerated catalyst exits the regenerator is due to the short conveying pipe between the regenerator and the second upward reactor. It can be regarded as the temperature when the second catalyst enters step S4.
  • the outlet temperature T3 of the second upward reactor is 530-650°C, preferably 560-640°C, more preferably 580-630°C, and still more preferably 600-630°C; And/or, the temperature T0 of the second strand of catalyst entering step S4 is 690-750°C, preferably 700-740°C, more preferably 705-730°C, still more preferably 710-725°C.
  • the third reaction oil and gas are separated to obtain dry gas, C3 fraction, C4 fraction, light gasoline, heavy gasoline, diesel oil, and oil slurry, from which light olefins and light aromatics are separated and separated.
  • a light olefin fraction is produced.
  • C4 fraction and/or light gasoline are light olefin fractions.
  • the third reaction oil and gas is introduced into a fractionation device or a gas separation device to achieve the above separation.
  • step S5 when step S2' does not exist, in step S5, light olefins and light aromatics are separated from any one of the first reaction oil and the third reaction oil or a mixture of the two, and the separated The light olefin fraction is returned to the second ascending reactor.
  • step S5 when step S2' exists, in step S5, light olefins and light aromatics are separated from any one of the second reaction oil and the third reaction oil or a mixture of the two, and the separated light The olefin fraction is returned to the fluidized bed reactor.
  • step S4 the light olefin fraction from the following step S5 is contacted with the catalyst mixture prior to the heavy distillate oil to undergo a catalytic cracking reaction, and then the heavy distillate oil is then contacted with the catalyst mixture , a catalytic cracking reaction occurs.
  • the light olefin fraction is contacted with the catalyst mixture 0.3-1.0 seconds prior to the heavy fraction. More preferably, the light olefin fraction is contacted with the catalyst mixture 0.4-0.8 seconds prior to the heavy distillate.
  • gas-solid separation is performed on the product of the third catalytic cracking to obtain the third reacted oil and gas and the third catalyst to be produced.
  • the method of gas-solid separation is not particularly limited, and methods known in the art can be used, for example, a settler and a cyclone separator are used to separate the catalyst and the third reaction oil and gas.
  • the material after the third catalytic cracking is subjected to gas-solid separation, and the separated catalyst is further stripped to remove the hydrocarbon products adsorbed therein to obtain the third to-be-grown catalyst.
  • the third catalyst to be grown is entered into a regenerator for catalyst regeneration.
  • the temperature of the regenerator is a temperature commonly used in the field, which may be 690-750°C, preferably 700-740°C, more preferably 705-730°C, and still more preferably 710-725°C °C.
  • the temperature of the regenerator or the temperature of the catalyst when the regenerated catalyst leaves the regenerator can be regarded as the temperature when the second stream of catalyst enters step S4. Therefore, in an embodiment of the present disclosure, the temperature T0 when the second catalyst enters step S4 may be 690-750°C, preferably 700-740°C, more preferably 705-730°C, and still more preferably 710- 725°C.
  • the regenerated catalyst is used as a first strand of catalyst and a second strand of catalyst.
  • step S5 from any one of the first reaction oil, the second reaction oil and the third reaction oil, or the mixture of the first reaction oil and the third reaction oil or the second reaction oil
  • the mixture of oil and gas and the third reaction oil and gas is separated to obtain light olefins and light aromatic hydrocarbons, and the light olefin fraction is separated, and the light olefin fraction is returned to the second upward reactor in step S4 or in step S2'. in the fluidized bed reactor. More specifically, the reaction oil and gas are separated to obtain dry gas, C3 fraction, C4 fraction, light gasoline, heavy gasoline, diesel oil and oil slurry, from which light olefins and light aromatics are separated, and light olefin fractions are separated.
  • the C4 fraction and/or light gasoline is the light olefin fraction.
  • the reaction oil and gas are introduced into a fractionation unit or a gas separation unit to achieve the above separation.
  • the first reaction oil and gas and the third reaction oil and gas can be separated separately, or the two can be combined and separated in a unified manner; or, the second reaction oil and gas and the third reaction oil and gas can be separated. Oil and gas are separated separately, or they can be combined and separated together.
  • the method for separating the light olefin fraction from the reaction oil and gas is not limited, and the separation can be carried out in a manner known in the art, including but not limited to the following methods, after the reaction oil and gas enters the fractionation, absorption and stabilization unit , separate liquefied gas and stable gasoline, the liquefied gas enters the subsequent gas separation device to separate C3 fraction and C4 fraction, and the stable gasoline enters the light and heavy gasoline splitting tower to separate light gasoline and heavy gasoline.
  • the C4 fraction and/or light gasoline are the light olefin fractions. From it, light olefins and light aromatics can be separated.
  • a method for producing light olefins and light aromatics by catalytic cracking of hydrocarbon-containing feedstock oil comprising the following steps:
  • the hydrocarbon-containing feed oil is cut into light distillate oil and heavy distillate oil, and the weight ratio (light distillate oil/heavy distillate oil) of the light distillate oil relative to the heavy distillate oil is X;
  • the light distillate oil and the first catalyst are introduced into the first descending reactor, and the first catalytic cracking is carried out to obtain the material after the first catalytic cracking;
  • the continuous catalyst, the heavy distillate oil and the second catalyst are introduced into the second upward reactor, the third catalytic cracking is carried out, and then the gas-solid separation is carried out to obtain the third reaction oil and gas and the third catalyst to be produced;
  • the continuous catalyst is at least a part of the first catalyst to be grown; the weight ratio of the second catalyst to the continuous catalyst (second catalyst/continuous catalyst) is R;
  • T0 is the temperature (unit °C) at which the second strand of catalyst enters step S4, and T3 is the outlet temperature (unit °C) of the second upward reactor.
  • a method for producing light olefins and light aromatics by catalytic cracking of hydrocarbon-containing feedstock oil comprising the steps of:
  • the hydrocarbon-containing feed oil is cut into light distillate oil and heavy distillate oil, and the weight ratio (light distillate oil/heavy distillate oil) of the light distillate oil relative to the heavy distillate oil is X;
  • the light distillate oil and the first catalyst are introduced into the first descending reactor, and the first catalytic cracking is carried out to obtain the material after the first catalytic cracking;
  • the continuous catalyst, the heavy distillate oil and the second catalyst are introduced into the second upward reactor, the third catalytic cracking is carried out, and then the gas-solid separation is carried out to obtain the third reaction oil and gas and the third catalyst to be produced; the described
  • the continuous catalyst is at least a part of the second catalyst to be grown; the weight ratio of the second catalyst to the continuous catalyst (second catalyst/continuous catalyst) is R;
  • step S5 separating light olefins and light aromatics from either or the mixture of the second reaction oil and gas and the third reaction oil and gas, and separating out a light olefin fraction, and returning the light olefin fraction
  • step S2' separating light olefins and light aromatics from either or the mixture of the second reaction oil and gas and the third reaction oil and gas, and separating out a light olefin fraction, and returning the light olefin fraction
  • T0 is the temperature (unit °C) at which the second strand of catalyst enters step S4, and T3 is the outlet temperature (unit °C) of the second upward reactor.
  • the light olefin fraction is the C4 fraction in the reaction oil and/or the light gasoline.
  • the method for producing light olefins and light aromatic hydrocarbons by catalytic cracking of hydrocarbon-containing feedstock oil comprises the steps:
  • the light distillate oil and the first catalyst are introduced into the first descending reactor, and the first catalytic cracking is carried out to obtain the material after the first catalytic cracking;
  • Optional S2' sending the material after the first catalytic cracking into the fluidized bed reactor for the second catalytic cracking to obtain the material after the second catalytic cracking;
  • the continuous catalyst is the first catalyst or the second catalyst; the weight ratio of the second catalyst to the continuous catalyst is 0.2-5:1;
  • step S1 The method according to A1, wherein, in step S1, the cutting point of the cutting is any temperature between 200-380°C.
  • step S4 the weight ratio of the second catalyst to the continuous catalyst is 0.5-3:1.
  • the conditions for the first catalytic cracking include: the outlet temperature of the first down-type reactor is 610-720° C., and the gas-solid residence time is 0.1-3.0 seconds;
  • the conditions for the second catalytic cracking include: the reaction temperature in the fluidized bed reactor is 600-670° C., and the mass space velocity is 2-20 h ⁇ 1 ;
  • the conditions for the third catalytic cracking include: the outlet temperature of the second ascending reactor is 530-650° C., and the gas-solid residence time is 0.5-8 seconds.
  • the conditions for the first catalytic cracking include: the outlet temperature of the first down-type reactor is 650-690° C., and the gas-solid residence time is 0.5-1.5 seconds;
  • the conditions for the second catalytic cracking include: the reaction temperature in the fluidized bed reactor is 620-640° C., and the mass space velocity is 4-12 h ⁇ 1 ;
  • the conditions for the third catalytic cracking include: the outlet temperature of the second ascending reactor is 560-640° C., and the gas-solid residence time is 1.5-5 seconds.
  • the light olefin fraction is catalytically cracked with the second catalyst for 0.3-1.0 seconds prior to the heavy distillate oil; preferably, the light olefin fraction is prior to the heavy distillate oil for 0.4-0.8 seconds with the second catalyst.
  • the first oil and gas and the second oil and gas are separated to obtain dry gas, C3 fraction, C4 fraction, light gasoline, heavy gasoline, diesel oil and oil slurry;
  • the light olefin fractions are C4 fractions in the first gas and second gas and/or fractions in the range of 30-90° C. in the first gas and second gas.
  • hydrocarbon-containing feedstock oil is one or more of conventional mineral oil, coal liquefied oil, synthetic oil, oil sand oil, shale oil, tight oil and animal and vegetable oils and fats mixture.
  • first strand catalyst and the second strand catalyst each independently comprise an active component and a carrier, and the active component is selected from ultra-stable rare earth-containing or non-rare earth-containing components At least one of Y-type zeolite, ZSP series zeolite, high silica zeolite having a five-membered ring structure and beta zeolite.
  • A10 The method of A1, wherein the first strand of catalyst and the second strand of catalyst each independently comprise the regenerated catalyst.
  • the present disclosure also provides a device for producing light olefins and light aromatics by catalytic cracking of hydrocarbon-containing feedstock oil, the device comprising the following units:
  • a hydrocarbon-containing feed oil cutting unit in which the hydrocarbon-containing feed oil is cut into light distillate oil and heavy distillate oil, so that the weight ratio of the light distillate oil to the heavy distillate oil (light distillate oil/heavy distillate oil) is X,
  • the first downward reaction unit, the light distillate oil and the first catalyst are introduced from the top of the reaction unit, the first catalytic cracking is carried out, and the material after the first catalytic cracking is obtained below the reaction unit;
  • Optional fluidized bed reaction unit wherein, the material after the first catalytic cracking is introduced, and the second catalytic cracking is performed to obtain the material after the second catalytic cracking;
  • the first gas-solid separation unit wherein the material after the first catalytic cracking is introduced for gas-solid separation to obtain the first reaction oil and gas and the first catalyst to be produced, or the material after the second catalytic cracking is introduced for gas-solid separation Separation to obtain the second reaction oil and gas and the second catalyst to be generated;
  • the continuous catalyst, the second catalyst and the heavy distillate oil are introduced from the bottom of the reaction unit to carry out the third catalytic cracking, and the material after the third catalytic cracking is obtained above the reaction unit, so
  • the continuous catalyst is at least a part of the first catalyst to grow or at least a part of the second catalyst to grow, and the weight ratio of the second catalyst to the continuous catalyst (second catalyst/continuous catalyst) is R,
  • the second gas-solid separation unit wherein the material after the third catalytic cracking is introduced for gas-solid separation to obtain the third reaction oil and gas and the third catalyst to be produced;
  • a separation unit in which any one of the first reaction oil, the second reaction oil and the third reaction oil or a mixture of the first reaction oil and the third reaction oil or the second reaction oil and the third reaction oil is introduced
  • the mixture of oil and gas separates light olefins and light aromatics, and separates light olefin fractions, and returns the light olefin fractions to the second upward reaction unit or the fluidized bed reaction unit;
  • T0 is the temperature (unit °C) at which the second strand of catalyst enters the second upward reaction unit
  • T3 is the outlet temperature (unit °C) of the second upward reaction unit.
  • T0 is the temperature (unit °C) at which the second strand of catalyst enters the second upward reaction unit. Specifically, it refers to the temperature at which the second strand of catalyst enters the bottom of the second upward reaction unit before being mixed with the continuous catalyst.
  • the apparatus further includes a regeneration unit, wherein the third undeveloped catalyst and optionally the first or second undeveloped catalyst not entering the second ascending reactor are introduced , carry out coke regeneration to obtain a regenerated catalyst.
  • the temperature of the regeneration unit is a temperature commonly used in the art, which may be 690-750°C, preferably 700-740°C, more preferably 705-730°C, and still more preferably 710-725°C .
  • the outlet temperature T3 of the second upward reaction unit is 530-650°C, preferably 560-640°C, more preferably 580-630°C, still more preferably 600-630°C.
  • the temperature T0 of the second catalyst entering the second upward reaction unit is 690-750°C, preferably 700-740°C, more preferably 705-730°C, and still more preferably 710°C -725°C.
  • the device further includes a dehydration and desalination unit, wherein the hydrocarbon-containing feedstock oil is subjected to desalination and dehydration treatment, and the obtained dehydrated and desalted hydrocarbon-containing feedstock oil is introduced into the hydrocarbon-containing feedstock oil cutting unit for cutting.
  • the structure of the first descending reaction unit is not particularly limited, as long as the upper part of the first down-type reaction unit can realize the feeding and the lower part can realize the discharging, for example, it can be of equal diameter or variable diameter Downpipe reactor.
  • light olefins and light olefins are separated from any one of the first reaction oil and gas, the third reaction oil and gas, or a mixture of the two. aromatic hydrocarbons, and the separated light olefin fraction is returned to the second upward reaction unit.
  • the separation unit when there is a fluidized bed reaction unit, in the separation unit, light olefins and light aromatics are separated from any one of the second reaction oil and gas, the third reaction oil and gas, or a mixture of the two, and will be separated out
  • the light olefin fraction is returned to the fluidized bed reactor.
  • the first gas-solid separation unit and the second gas-solid separation unit include equipment known in the art that can realize gas-solid separation, such as a settler or a cyclone.
  • the device further includes at least one stripping unit, which can be arranged in the gas-solid separation unit, wherein the catalyst obtained by gas-solid separation is stripped to remove the hydrocarbon products adsorbed therein. .
  • the first gas-solid separation unit when the device includes a fluidized bed reaction unit, the first gas-solid separation unit further includes a stripping unit, wherein the catalyst obtained by gas-solid separation is stripped , in order to remove the hydrocarbon products adsorbed therein to obtain the second catalyst to be grown.
  • the second gas-solid separation unit further includes a stripping unit, wherein the catalyst obtained by gas-solid separation is stripped to remove the hydrocarbon products adsorbed therein, and the third to-be-generated catalyst is obtained. catalyst.
  • the continuous catalyst and the second strand of catalyst are introduced into the bottom of the second upward reaction unit, and after mixing, the mixed catalyst is used for the subsequent catalytic cracking reaction.
  • the position where the continuous catalyst and the second catalyst are introduced is upstream of the light olefin fraction feed port.
  • the light olefin fraction feed port is upstream of the heavy distillate oil feed port.
  • the structure of the second upward reactor is not particularly limited, as long as the material can be fed from the bottom and discharged from the top, for example, it can be of equal diameter or variable diameter.
  • the present disclosure provides a device for producing light olefins and light aromatics by catalytic cracking of hydrocarbon-containing feedstock oil, the device comprising the following units:
  • a hydrocarbon-containing feed oil cutting unit in which the hydrocarbon-containing feed oil is cut into light distillate oil and heavy distillate oil, so that the weight ratio of the light distillate oil to the heavy distillate oil (light distillate oil/heavy distillate oil) is X,
  • the first downward reaction unit, the light distillate oil and the first catalyst are introduced from the top of the reaction unit, the first catalytic cracking is carried out, and the material after the first catalytic cracking is obtained below the reaction unit;
  • a first gas-solid separation unit wherein the material after the first catalytic cracking is introduced for gas-solid separation to obtain the first reacted oil and gas and the first catalyst to be produced;
  • the continuous catalyst, the second catalyst and the heavy distillate oil are introduced from the bottom of the reaction unit to carry out the third catalytic cracking, and the material after the third catalytic cracking is obtained above the reaction unit, so
  • the continuous catalyst is at least a part of the first catalyst to be grown, and the weight ratio of the second catalyst to the continuous catalyst (second catalyst/continuous catalyst) is R,
  • the second gas-solid separation unit wherein the material after the third catalytic cracking is introduced for gas-solid separation to obtain the third reaction oil and gas and the third catalyst to be produced;
  • a separation unit in which either or a mixture of the first and third reacted oil and gas is introduced, light olefins and light aromatics are separated, and a light olefin fraction is separated and returned to the light olefin fraction to the second upward reaction unit;
  • T0 is the temperature (unit °C) at which the second strand of catalyst enters the second upward reaction unit
  • T3 is the outlet temperature (unit °C) of the second upward reaction unit.
  • the present disclosure provides a device for producing light olefins and light aromatics by catalytic cracking of hydrocarbon-containing feedstock oil, the device comprising the following units:
  • a hydrocarbon-containing feed oil cutting unit in which the hydrocarbon-containing feed oil is cut into light distillate oil and heavy distillate oil, so that the weight ratio of the light distillate oil to the heavy distillate oil (light distillate oil/heavy distillate oil) is X,
  • the first downward reaction unit, the light distillate oil and the first catalyst are introduced from the top of the reaction unit, the first catalytic cracking is carried out, and the material after the first catalytic cracking is obtained below the reaction unit;
  • a fluidized bed reaction unit wherein the material after the first catalytic cracking is introduced, and the second catalytic cracking is performed to obtain the material after the second catalytic cracking;
  • the first gas-solid separation unit wherein the material after the second catalytic cracking is introduced to carry out gas-solid separation to obtain the second reaction oil and gas and the second catalyst to be produced;
  • the continuous catalyst, the second catalyst and the heavy distillate oil are introduced from the bottom of the reaction unit to carry out the third catalytic cracking, and the material after the third catalytic cracking is obtained above the reaction unit, so
  • the continuous catalyst is at least a part of the second catalyst to be grown, and the weight ratio of the second catalyst to the continuous catalyst (second catalyst/continuous catalyst) is R,
  • the second gas-solid separation unit wherein the material after the third catalytic cracking is introduced for gas-solid separation to obtain the third reaction oil and gas and the third catalyst to be produced;
  • a separation unit in which either or a mixture of the second reaction oil and the third reaction oil is introduced, light olefins and light aromatics are separated, and a light olefin fraction is separated and returned to the light olefin fraction to the fluidized bed reaction unit;
  • T0 is the temperature (unit °C) at which the second strand of catalyst enters the second upward reaction unit
  • T3 is the outlet temperature (unit °C) of the second upward reaction unit.
  • the device for producing light olefins and light aromatics by catalytic cracking of hydrocarbon-containing feedstock oil of the present disclosure is used to implement the method for producing light olefins and light aromatics by catalytic cracking of hydrocarbon-containing feedstock oil of the present disclosure.
  • the hot first stream of catalyst (regenerated catalyst) is transported to the first descending reactor 1 through a first stream of catalyst conveying pipe (regenerated catalyst conveying pipe) 12 .
  • the light distillate oil is injected into the first down-type reactor 1 through the feed nozzle 11, and contacts with the first stream of catalyst and undergoes a catalytic cracking reaction.
  • the obtained first reaction oil and gas is introduced into the separation device (not shown in the figure) through the oil and gas outlet 22 of the descending reactor, and the first catalyst to be produced is introduced into the second ascending reactor 3 as a continuous catalyst through the continuous catalyst delivery pipe 31
  • the second catalyst regenerated catalyst
  • the second catalyst conveying pipe regenerated catalyst conveying pipe
  • the catalyst is lifted upward by a pre-lift medium.
  • the light olefin fraction is injected into the second upward reactor 3 through the light olefin fraction feed nozzle 21, and contacts and reacts with the catalyst.
  • the heavy distillate oil is sprayed into the second upward reactor 3 through the heavy distillate oil feed nozzle 33 to contact and react with the oil-agent mixture from the bottom, and after the reaction, the third catalytically cracked material is obtained, which is allowed to enter the settler 4, and then settles.
  • the separation of the third catalyst to be produced and the third reaction oil and gas is carried out, the third reaction oil and gas enters the separation device (not shown in the figure) through the oil and gas outlet 41 of the third reactor, and the third catalyst to be produced enters the stripper 5.
  • reaction oil and gas (the first reaction oil and gas, the third reaction oil and gas) are separated by a separation device (preferably a fractionation device, a gas separation device) to obtain dry gas, C3 fraction, C4 fraction, light gasoline, heavy gasoline, diesel oil, oil slurry, from which Light olefins and light aromatics are obtained by separation.
  • a separation device preferably a fractionation device, a gas separation device
  • the light olefin fraction is separated from the reaction oil and gas, and the light olefin fraction is introduced into the second upward reactor 3 through the light olefin fraction feed nozzle 21 .
  • the hot first stream of catalyst (regenerated catalyst) is transported to the first down-type reactor 1 through the first stream of catalyst conveying pipe (regenerated catalyst conveying pipe) 12 .
  • the light distillate oil is injected into the first down-type reactor 1 through the feed nozzle 11, contacted with the first stream of catalyst and undergoes a catalytic cracking reaction, and the reacted first catalytically cracked material passes through the first down-type reactor.
  • the outlet mushroom head distributor 13 is introduced into the fluidized bed reactor 2, and the cracking reaction continues in the fluidized bed reactor. After the reaction, the second catalytic cracked material is obtained, and the material is cyclone separated to obtain the second reaction oil and gas. and the second living catalyst.
  • the second reaction oil and gas is introduced into the separation device (not shown in the figure) from the second reaction oil and gas outlet 22.
  • the second catalyst to be produced enters the first stripper 51 to strip off the adsorbed hydrocarbon products, it is passed through the continuous catalyst as a continuous catalyst.
  • the catalyst conveying pipe 31 is introduced into the bottom of the second ascending reactor 3
  • the second catalyst (regenerated catalyst) is introduced into the bottom of the second ascending reactor 3 through the second catalyst conveying pipe (regenerated catalyst conveying pipe) 32, and the second to-be-generated
  • the mixed catalyst (continuous catalyst) and the second stream of catalyst are lifted upwards by a pre-lift medium.
  • the heavy distillate oil is sprayed into the second up-flow reactor 3 through the heavy distillate oil feed nozzle 33 to contact and react with the catalyst, and after the reaction, the third catalytically cracked material is obtained, which enters the settler 4, and the first step is carried out in the settler 4.
  • the third reaction oil and gas is introduced into the separation device (not shown in the figure) through the oil and gas outlet 41 of the third reactor, the third catalyst to be produced enters the second stripper 52, and is stripped out
  • the adsorbed hydrocarbon products are sent to the regenerator 6 by the third catalyst pipeline 53 for regeneration, and the regenerated catalyst is returned to the first down-type reactor and the second up-type reactor for reuse. .
  • the reaction oil and gas (the second reaction oil and gas, the third reaction oil and gas) are separated to obtain dry gas, C3 fraction, C4 fraction, light gasoline, heavy gasoline, diesel oil, and oil slurry through a separation device (preferably a fractionation device, a gas separation device), and therefrom.
  • a separation device preferably a fractionation device, a gas separation device
  • Light olefins and light aromatics are obtained by separation.
  • the light olefin fraction is separated from the reaction oil and gas, and the light olefin fraction is returned to the fluidized bed reactor 2 through the light olefin fraction feed nozzle 21 .
  • the present disclosure is further described in detail below by means of examples.
  • the raw materials used in the examples can be obtained through commercial sources.
  • the catalytic cracking catalyst used in the examples and comparative examples of the present disclosure is industrially produced by the Catalyst Qilu Branch of China Petrochemical Corporation, and the trade name is DMMC-2.
  • the catalyst contains ZSM-5 zeolite and ultra-stable Y-type zeolite with an average pore size of less than 0.7 nanometers.
  • the catalyst was hydrothermally aged with saturated steam at a temperature of 800 °C for 17 hours before use.
  • the main physicochemical properties of the catalyst are shown in Table 1.
  • the hydrocarbon-containing feedstock oil used in the Examples and Comparative Examples is crude oil from Jiangsu Oilfield, and its properties are listed in Table 2.
  • catalyst catalyst physical properties Specific surface/m 2 ⁇ g -1 125 Pore volume/cm -3 ⁇ g -1 0.197 Apparent density g ⁇ cm -3 0.86 chemical components Al 2 O 3 /% 56.8 SiO 2 /% 42.9 Micro-reactive/% 68
  • the light and heavy distillate oil cutting point of crude oil A processed in this example is 320° C., and the cutting ratio is (weight ratio of light distillate oil/heavy distillate oil) 0.4.
  • a modified continuous reaction-regeneration operation medium-sized device was used to carry out the test, and the flow chart is shown in FIG. 1 .
  • the high temperature regenerated catalyst at 720°C is introduced into the top of the descending tubular reactor 1 from the regenerator through the regenerating inclined tube, and the light distillate oil preheated to 45 °C is atomized by steam, and then enters the descending tubular reactor 1 and
  • the first strand of catalyst is contacted for catalytic cracking reaction, the ratio of agent to oil is 40, the outlet temperature of the reactor is 665 °C, the gas-solid residence time is 0.8s, and the material after the first catalytic cracking is separated by cyclone to separate the first reaction oil and gas and the second reaction gas.
  • the first reacted oil and gas enters the separation system, and all the first catalyst to be produced is introduced into the bottom of the riser reactor 3 .
  • the regenerated catalyst (the second stream of catalyst) with a temperature of 720° C. was introduced into the bottom of the riser reactor 3 by the regenerator via the regenerated catalyst delivery pipe 32 .
  • the weight ratio of the second catalyst to the first catalyst (the second catalyst/the first catalyst) is 0.25. After the first catalyst and the second catalyst are mixed at the bottom of the riser reactor 3, they are mixed.
  • the catalyst flows upward under the action of pre-lift steam, and at the same time, the light olefin fraction enters the lower part of the riser reactor 3 through the light olefin fraction feed nozzle under the atomized water vapor medium, and reacts with the mixed catalyst.
  • the heavy distillate oil is atomized by water vapor, it is sprayed into the riser through the heavy distillate oil feed nozzle to react, and the catalytic cracking reaction occurs.
  • the ratio of agent to oil is 20, and the outlet temperature T3 of the reactor is At 610°C, the gas-solid residence time in the reactor is 1.5s, and the material after catalytic cracking is introduced into the settler for oil separation, which is separated into the third reaction oil and gas and the third to-be-generated catalyst, and the third reaction oil and gas is introduced into the separation system.
  • the first reaction oil and gas and the third reaction oil and gas are separated into cracked gas, light gasoline, heavy gasoline, diesel oil and oil slurry in the separation system.
  • the light gasoline partial fraction (with a distillation range of 30-60° C.) is returned to the riser reactor 3 as a light olefin fraction through a light olefin fraction feed nozzle.
  • the third to-be-grown catalyst enters the stripper, and after stripping the hydrocarbon products adsorbed on the third to-be-grown catalyst, it enters the regenerator through the inclined pipe of the unborn catalyst, and is scorched and regenerated by contacting with air at 720°C.
  • the regenerated catalyst is returned to the reactor through the regeneration inclined pipe for recycling.
  • Medium-sized units use electrical heating to maintain the temperature of the reaction-regeneration system.
  • Example 2 The same equipment and reaction steps as in Example 1 were used, except that the cutting point of the light and heavy components of the processed crude oil A was 250°C, and the cutting ratio was (weight ratio of light distillate oil/heavy distillate oil) 0.195.
  • the weight ratio of the second strand of catalyst to the first unborn catalyst (second strand of catalyst/first unborn catalyst) is 0.03, and the temperature of the regenerator is 700°C (that is, the temperature T0 when the second strand of catalyst enters step S4 is 700°C °C), the outlet temperature T3 of the riser reactor was 570 °C.
  • Example 2 The same equipment and reaction steps as in Example 1 were used, except that the cutting point of the light and heavy components of the processed crude oil A was 350°C, and the cutting ratio was (weight ratio of light distillate oil/heavy distillate oil) 0.529.
  • the weight ratio of the second strand of catalyst to the first unborn catalyst (the second strand of catalyst/the first unborn catalyst) is 0.6, and the temperature of the regenerator is 740° C. (that is, the temperature T0 when the second strand of catalyst enters step S4 is 740° C. °C), the outlet temperature T3 of the riser reactor is 630 °C.
  • the light and heavy distillate oil cutting point of crude oil A processed in this example is 250° C., and the cutting ratio is (weight ratio of light distillate oil/heavy distillate oil) 0.195.
  • the light and heavy distillate oil cutting point of crude oil A processed in this example is 350°C, and the cutting ratio is (weight ratio of light distillate oil/heavy distillate oil) 0.529,
  • the light and heavy distillate oil cutting point of the processed crude oil A was 320° C., and the cutting ratio (weight ratio of light distillate oil/heavy distillate oil) was 0.4.
  • a modified continuous reaction-regeneration operation medium-sized device was used to conduct the test, and the flow chart is shown in FIG. 2 .
  • the high-temperature regenerated catalyst at 720°C is introduced into the top of the descending tubular reactor 1 from the regenerator through the regeneration inclined tube, and the light distillate oil preheated to 45 °C is atomized by steam, and then enters the descending tubular reactor 1 through the feed nozzle.
  • the ratio of agent to oil is 40
  • the outlet temperature of the reactor is 670 ° C
  • the gas-solid residence time is 0.6s
  • the material after the first catalytic cracking enters the fluidized bed reaction through the outlet distributor Device 2
  • the reaction temperature is 655 ° C
  • the mass space velocity is 4h ⁇ 1 ;
  • the light olefin fraction enters the bottom of the fluidized bed reactor 2 through the feed nozzle 21 after being atomized by water vapor, and is heated with heat.
  • the catalyst is contacted and reacted, and the material after the second catalytic cracking is subjected to cyclone separation to obtain the second reaction oil and gas and the second catalyst.
  • the regenerated catalyst (the second stream of catalyst) with a temperature of 720° C. was introduced into the bottom of the second ascending pipe reactor 3 by the regenerator via the regenerated catalyst delivery pipe 32 .
  • the weight ratio of the second catalyst to the second catalyst (second catalyst/second catalyst) is 0.25.
  • the mixed catalyst is Under the action of pre-lifting steam, it flows upward.
  • the heavy distillate oil is atomized by steam, it is sprayed into the riser reactor 3 through the heavy distillate oil nozzle, and it contacts with the catalyst for catalytic cracking reaction.
  • the ratio of agent to oil is 20.
  • the outlet temperature of the reactor T3 is 610°C
  • the gas-solid residence time in the reactor is 1.5s
  • the material after catalytic cracking is introduced into the settler for oil separation, separated into the third reaction oil and gas and the third catalyst to be produced, and the reaction oil and gas is introduced into the separation system.
  • the second reaction oil and gas and the third reaction oil and gas are separated into cracked gas, light gasoline, heavy gasoline, diesel oil and oil slurry in the separation system.
  • the light gasoline partial fraction (with a distillation range of 30-60° C.) is returned to the fluidized bed reactor 2 as a light olefin fraction.
  • the third to-be-grown catalyst enters the stripper, and after stripping the hydrocarbon products adsorbed by the third to-be-grown catalyst, it enters the regenerator through the inclined pipe of the unborn catalyst, and is in contact with air for coke regeneration at 720°C.
  • the regenerated catalyst is returned to the reactor through the regeneration inclined pipe for recycling.
  • the medium-sized device adopts electric heating to maintain the temperature of the reaction and regeneration system. After the device runs stably (the product composition remains basically unchanged), the composition of the cracked gas and gasoline obtained from the reaction oil and gas is analyzed to obtain the yields of triene and BTX in the product.
  • Example 6 The same equipment and reaction steps as in Example 6 were used, except that the cutting point of the light and heavy components of the processed crude oil A was 250°C, and the cutting ratio was (weight ratio of light distillate oil/heavy distillate oil) 0.195.
  • the weight ratio of the second strand of catalyst to the second unborn catalyst is 0.03, and the temperature of the regenerator is 700°C (that is, the temperature T0 when the second strand of catalyst enters step S4 is 700°C °C), the outlet temperature T3 of the riser reactor was 570 °C.
  • Example 6 The same equipment and reaction steps as in Example 6 were used, except that the cutting point of the light and heavy components of the processed crude oil A was 350°C, and the cutting ratio was (weight ratio of light distillate oil/heavy distillate oil) 0.529.
  • the weight ratio of the second strand of catalyst to the second unborn catalyst (second strand of catalyst/second unborn catalyst) is 0.6, and the temperature of the regenerator is 740°C (that is, the temperature T0 when the second strand of catalyst enters step S4 is 740°C °C), the outlet temperature T3 of the riser reactor is 630 °C.
  • Example 2 The same device, reaction steps, and reaction conditions are used as in Example 1, the difference is that the light olefin fraction separated from the reaction oil and gas is not returned to the riser reactor 3 .
  • Example 2 The same device and reaction steps as in Example 1 were used, except that the weight ratio of the second catalyst to the first catalyst (the second catalyst/the first catalyst) was 0.1, and the temperature of the regenerator was 740 °C (that is, the temperature T0 when the second catalyst enters step S4 is 740 °C), and the outlet temperature T3 of the riser reactor is 610 °C.
  • Example 2 The same device and reaction steps as in Example 1 were used, except that the weight ratio of the second catalyst to the first catalyst (the second catalyst/the first catalyst) was 0.5, and the temperature of the regenerator was 700 °C (that is, the temperature T0 at which the second stream of catalyst enters step S4 is 700 °C), and the outlet temperature T3 of the riser reactor is 610 °C.
  • Example 1 The same device, reaction step, and reaction conditions are adopted as in Example 1, except that the second catalyst is not introduced at the lower part of the riser reactor 3, but only the first catalyst to be produced is used.
  • Example 1 The same device, reaction steps, and reaction conditions were used as in Example 1, except that the first catalyst to be grown was not introduced into the lower part of the riser reactor 3, and only the second catalyst was used.
  • Example 2 The same device, reaction steps, and reaction conditions are adopted as in Example 1, except that the light distillate oil enters the riser reactor, that is, the first descending pipe reactor 1 becomes the riser reactor (the descending pipe in Table 3).
  • the parameters of the type reactor are shown in this Comparative Example 6 for the parameters of the ascending tubular reactor).
  • the method for producing light olefins and BTX by catalytic cracking of hydrocarbon-containing feedstock oil provided by the present disclosure can significantly improve the yield of light olefins and light aromatic hydrocarbons. The yield of the product is suppressed.

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Abstract

一种含烃原料油催化裂解生产低碳烯烃和轻芳烃的方法和装置，该方法包括：将含烃原料油切割为轻馏分油和重馏分油；将轻馏分油与第一股催化剂引入下行式反应器进行催化裂解，得到第一催化裂解后的物料；将第一催化裂解后的物料进行气固分离，得到第一反应油气和第一待生催化剂；或者，将第一催化裂解后的物料送入流化床反应器进行催化裂解后，再进行气固分离，得到第二反应油气和第二待生催化剂；将连续催化剂、重馏分油与第二股催化剂引入上行式反应器进行催化裂解后气固分离，得到第三油气和第三待生催化剂；从反应油气中分离出低碳烯烃和轻芳烃，并且分离出轻烯烃馏分，并将轻烯烃馏分返回流化床反应器中或上行式反应器中。该方法可以明显提高低碳烯烃和轻芳烃的产率和装置经济性。

Description

含烃原料油催化裂解生产低碳烯烃和BTX的方法及装置 技术领域
本申请涉及石油炼制及石油化工加工过程，具体地，涉及全馏分含烃原料油催化裂解生产低碳烯烃和BTX的方法及装置。
背景技术
随着成品油消费量增速的持续放缓，以低碳烯烃以及芳烃等为代表的基本有机原料需求量的高速增长，化工料型炼厂成为未来的发展趋势。目前化工型炼厂的构型主要包括以下三种，一是原油通过溶剂脱沥青或加氢精制等预处理后，直接进入蒸汽裂解单元生产化工料，但这种方式一般仅限于轻质原油；二是通过原油各馏分加氢裂化后，最大化生成重质石脑油，然后通过重整单元最大化生产芳烃；三是原油的轻馏分进入蒸汽裂解单元，重馏分进入催化裂解单元，最大化生产低碳烯烃。以上三种方式均已实现工业化，化学品的收率在35％-55％之间。可以看出，现有化工型炼厂的构型主要依托蒸汽裂解、重整、加氢精制、加氢裂化、催化裂解等多套核心装置的组合。其中，催化裂解工艺在生产化工料和原料适应性方面具有其独特的优势，可同时生产丙烯、乙烯和BTX。
中国专利CN1978411B公开了一种制取小分子烯烃的组合工艺方法，该方法中，将催化裂解催化剂和裂解原料在一个反应器中混合接触，分离待生催化剂和反应油气，其中待生催化剂送入再生器进行烧焦再生，再生后的热催化剂分为两部分，其中一部分再生后的热催化剂返回上述反应器；另一部分再生后的热催化剂先和重质石油烃类在另外一个反应器中混合接触，进行预结焦，富含C4-C8的烯烃原料再和结焦的催化剂混合接触，发生催化裂解反应，分离待生催化剂反应油气，该待生催化剂与上步中所述待生催化剂送入再生器进行烧焦再生；分离反应油气得到丙烯等小分子烯烃目的产品。该方法可将富含烯烃的轻质原料高选择性地转化为丙烯等小分子烯烃产物，同时维持装置自身的热平衡。
中国专利CN102899078A公开了一种生产丙烯的催化裂解方法，该方法基于双提升管与流化床构成的组合反应器，首先将重质原料油 与第一股催化剂引入第一提升管反应器进行反应，油剂分离后进入分离系统。将裂解重油引入第二提升管反应器与引入第二提升管反应器的催化剂接触反应，将轻质烃引入第二提升管反应器，与裂解重油和第二股裂解催化剂接触反应形成的混合物接触，所述轻质烃包括产品分离系统得到的C4烃或汽油馏分。然后将第二提升管反应器反应后的油气与催化剂引入流化床反应器反应。通过工艺方案的优化，配备合适的催化剂，对不同进料进行选择性转化，具有较高的丙烯和丁烯产率。
中国专利CN101045667B公开了一种多产低碳烯烃的组合式催化转化方法，该方法将重油原料在下行管反应器内与再生催化剂和任选的积炭催化剂接触，将分离的除低碳烯烃的其余产物中至少一部分引入提升管反应器与再生催化剂接触反应，提升管反应后催化剂引入下行管反应器催化剂预提升段，与进入下行管反应器的再生催化剂混合后和重油原料接触。该方法采用重油原料在下行式反应器反应，中间产物烯烃在提升管反应器内反应的组合反应器形式，提高低碳烯烃的产率。
中国专利CN109370644A公开了一种原油催化裂解制低碳烯烃和芳烃的方法，该方法将原油分成轻重组分，切割点在150℃-300℃之间，轻馏分和重馏分在同一个反应器不同反应区进行反应，催化剂采用以二氧化硅和三氧化二铝组成的硅铝酸盐为主要组分，包括碱金属氧化物，碱土金属氧化物，钛、铁氧化物，钒和镍的氧化物。该方法是基于重油催化裂解生成低碳烯烃的密相输送床反应器的基础上，针对原油催化裂解生成低碳烯烃提出的解决方案。
以上方法从新型反应器结构开发、新型催化材料研制以及控制反应深度提高丙烯选择性等方面进行研究，提出了催化裂解制取低碳烯烃和芳烃的方法，但目前仍没有针对原油最大化生产化工料的制备方法和反应器结构。
发明内容
本公开的目的在于，针对各类含烃原料油的烃类组成不同、切割温度不同的特点，提出一种适于加工含烃原料油进行催化裂解，以最大限度的利用含烃原料油生产低碳烯烃和BTX的装置及方法。
为了实现上述目的，本公开提供了一种含烃原料油催化裂解生产低碳烯烃和轻芳烃的方法，该方法包括如下步骤：
S1、将含烃原料油切割为轻馏分油和重馏分油，所述轻馏分油相对于所述重馏分油的重量比(轻馏分油/重馏分油)为X；
S2、将所述轻馏分油与第一股催化剂引入第一下行式反应器，进行第一催化裂解，得到第一催化裂解后的物料；
任选的S2′、将所述第一催化裂解后的物料引入流化床反应器进行第二催化裂解，得到第二催化裂解后的物料；
S3、将所述第一催化裂解后的物料进行气固分离，得到第一反应油气和第一待生催化剂，或者将所述第二催化裂解后的物料进行气固分离，得到第二反应油气和第二待生催化剂；
S4、将连续催化剂、所述重馏分油与第二股催化剂引入第二上行式反应器，进行第三催化裂解，然后进行气固分离，得到第三反应油气和第三待生催化剂；所述连续催化剂为所述第一待生催化剂的至少一部分或所述第二待生催化剂的至少一部分；所述第二股催化剂与所述连续催化剂的重量比(第二股催化剂/连续催化剂)为R；
S5、从所述第一反应油气、所述第二反应油气和所述第三反应油气的任一者或者第一反应油气和第三反应油气的混合物或者第二反应油气和第三反应油气的混合物中分离出低碳烯烃和轻芳烃，并且分离出轻烯烃馏分，并将所述轻烯烃馏分返回步骤S4的所述第二上行式反应器中或者步骤S2′的所述流化床反应器中，
所述R和X满足以下关系式：
(4.84×T0-3340)/(780+5×T0-6×T3)＜R/X＜(0.968×T0-630)/(668+0.2×T0-1.2×T3)
T0为所述第二股催化剂进入步骤S4时的温度(单位℃)，T3为所述第二上行式反应器的出口温度(单位℃)。
可选地，本公开的方法中，所述第二上行式反应器的出口温度T3为530-650℃，优选为560-640℃，进一步优选为580-630℃，更进一步优选为600-630℃；和/或，所述第二股催化剂进入步骤S4时的温度T0为690-750℃，优选700-740℃，进一步优选为705-730℃，更进一步优选为710-725℃。
可选地，本公开的方法中，步骤S1中，在切割点100-400℃之间 的任意温度将含烃原料油切割为轻馏分油和重馏分油，使所述轻馏分油相对于所述重馏分油的重量比(轻馏分油/重馏分油)为X。
可选地，本公开的方法中，所述第一下行式反应器中，所述第一催化裂解的条件包括：所述第一下行式反应器的出口温度为610-720℃，气固停留时间为0.1-3.0秒，剂油比为15-80；和/或，所述流化床反应器中，所述第二催化裂解的条件包括：所述流化床反应器中的反应温度为600-690℃，质量空速为2-20h -1；和/或，所述第二上行式反应器中，所述第三催化裂解的条件包括：气固停留时间为0.5-8秒，剂油比为8-40。
可选地，本公开的方法中，所述第一下行式反应器中，所述第一催化裂解的条件包括：所述第一下行式反应器的出口温度为650-690℃，气固停留时间为0.5-1.5秒，剂油比为25-65；和/或，所述流化床反应器中，所述第二催化裂解的条件包括：所述流化床反应器中的反应温度为640-670℃，质量空速为4-12h -1；和/或，所述第二上行式反应器中，所述第三催化裂解的条件包括：气固停留时间为1.5-5秒，剂油比为10-30。
可选地，本公开的方法中，在步骤S4中，首先使所述连续催化剂与所述第二股催化剂混合，再进行后续催化裂解反应，和/或，存在步骤S2′时，在步骤S3的气固分离中，将分离的催化剂进行汽提，得到第二待生催化剂；和/或，在步骤S4中，将来自S5步骤的所述轻烯烃馏分先于所述重馏分油与所述第二股催化剂和连续催化剂的混合物接触进行催化裂解；优选将所述轻烯烃先于所述重馏分油0.3-1.0秒与所述第二股催化剂和连续催化剂的混合物接触进行催化裂解，更优选将所述轻烯烃馏分先于所述重馏分油0.4-0.8秒与所述第二股催化剂和连续催化剂的混合物接触进行催化裂解；和/或，所述方法在步骤S1之前具有步骤S0，其中，将含烃原料油进行脱盐脱水处理，将得到的经脱水脱盐的含烃原料油引入步骤S1进行切割。
可选地，本公开的方法中，该方法还包括：在步骤S4的气固分离中，将分离的的催化剂进行汽提，得到第三待生催化剂；和/或，将所述第三待生催化剂和任选的未进入第二上行式反应器的第一待生催化剂或第二待生催化剂，在690-750℃、优选700-740℃、进一步优选705-730℃、更进一步优选710-725℃的温度下进行烧焦再生，得到再 生催化剂；和/或，将所述第一反应油气、所述第二反应油气和所述第三反应油气的任一者或者第一反应油气和第三反应油气的混合物或者第二反应油气和第三反应油气的混合物进行分离，得到干气、C3馏分、C4馏分、轻汽油、重汽油、柴油和油浆，从中分离得到低碳烯烃、轻芳烃，并分离出轻烯烃馏分；和/或，不存在步骤S2′时，在步骤S5中，从第一反应油气、第三反应油气的任一者或者两者的混合物中分离出轻烯烃馏分，并将所述轻烯烃馏分返回步骤S4的第二上行式反应器中；存在步骤S2′时，在步骤S5中，从第二反应油气、第三反应油气的任一者或者两者的混合物中分离出轻烯烃馏分，并将所述轻烯烃馏分返回步骤S2′的流化床反应器中。
可选地，本公开的方法中，所述含烃原料油为原油、煤液化油、合成油、油砂油、页岩油、致密油和动植物油脂中的一种或其两种以上的混合物，，或其各自的部分馏分、其各自的重质馏分的加氢改质油。
可选地，本公开的方法中，所述第一股催化剂和所述第二股催化剂各自独立地包括活性组分和载体，所述活性组分为选自含或不含稀土的超稳Y型沸石、ZSM-5系列沸石、具有五元环结构的高硅沸石和β沸石中的至少一种，所述载体为选自氧化铝、氧化硅、无定形硅铝、氧化锆、氧化钛、氧化硼和碱土金属氧化物中的至少一种。
可选地，本公开的方法中，所述第一股催化剂和所述第二股催化剂各自独立地包括再生催化剂，优选所述第一股催化剂和所述第二股催化剂为再生催化剂，和/或，将所述第一待生催化剂的全部或所述第二待生催化剂的全部作为连续催化剂。
本公开还提供一种含烃原料油催化裂解生产低碳烯烃和轻芳烃的装置，该装置包括以下单元：
含烃原料油切割单元，在其中将含烃原料油切割为轻馏分油和重馏分油，使得所述轻馏分油相对于所述重馏分油的重量比(轻馏分油/重馏分油)为X，
第一下行式反应单元，从该反应单元的上方引入所述轻馏分油与第一股催化剂，进行第一催化裂解，在该反应单元的下方得到第一催化裂解后的物料；
任选的流化床反应单元，其中，引入所述第一催化裂解后的物料，进行第二催化裂解，得到第二催化裂解后的物料；
第一气固分离单元，其中引入所述第一催化裂解后的物料进行气固分离，得到第一反应油气和第一待生催化剂，或者其中引入所述第二催化裂解后的物料进行气固分离，得到第二反应油气和第二待生催化剂；
第二上行式反应单元，从该反应单元的下方引入连续催化剂、第二股催化剂与所述重馏分油，进行第三催化裂解，在该反应单元的上方得到第三催化裂解后的物料，所述连续催化剂为所述第一待生催化剂的至少一部分或所述第二待生催化剂的至少一部分，所述第二股催化剂与所述连续催化剂的重量比(第二股催化剂/连续催化剂)为R，
第二气固分离单元，其中引入所述第三催化裂解后的物料进行气固分离，得到第三反应油气和第三待生催化剂；
分离单元，其中引入所述第一反应油气、所述第二反应油气和所述第三反应油气的任一者或者第一反应油气和第三反应油气的混合物或者第二反应油气和第三反应油气的混合物，分离出低碳烯烃和轻芳烃，并且分离出轻烯烃馏分，并将轻烯烃馏分返回至所述第二上行式反应单元或者所述流化床反应单元；
其中，所述R和X满足以下关系式：
(4.84×T0-3340)/(780+5×T0-6×T3)＜R/X＜(0.968×T0-630)/(668+0.2×T0-1.2×T3)
T0为所述第二股催化剂进入第二上行式反应单元时的温度(单位℃)，T3为所述第二上行式反应单元的出口温度(单位℃)。
可选地，本公开的装置中，还包括再生单元，其中，引入所述第三待生催化剂和任选的未进入第二上行式反应器的第一待生催化剂或第二待生催化剂，在690-750℃、优选700-740℃、进一步优选705-730℃、更进一步优选710-725℃的温度下进行烧焦再生，得到再生催化剂。
可选地，本公开的装置中，所述装置在包括流化床反应单元时，第一气固分离单元中还包括汽提单元，其中，将气固分离得到的催化剂进行汽提，得到第二待生催化剂。
所述第二气固分离单元中还包括汽提单元，其中，将气固分离得到的催化剂进行汽提，得到第三待生催化剂。
可选地，本公开的装置中，所述装置还包括脱水脱盐单元，其中，将含烃原料油进行脱盐脱水处理，将得到的经脱水脱盐的含烃原料油 引入含烃原料油切割单元进行切割。
可选地，本公开的装置中，第二上行式反应单元中，引入连续催化剂和第二股催化剂的位置在轻烯烃馏分的进料口的上游。
可选地，本公开的装置中，第二上行式反应单元中，来自分离单元的轻烯烃馏分的进料口在重馏分油进料口的上游。
技术效果
本公开中，通过上述特定的方法，根据含烃原料油不同馏分段的烃类组成特点和裂解反应特性，将含烃原料油切割为轻馏分油和重馏分油两部分，将轻馏分油在下行式反应器中高温下、短停留时间进行裂解，可以高选择性生产低碳烯烃和BTX，同时可以明显降低甲烷生成。同时，对于重馏分油，通过采用上行式反应器，可以最大化生产低碳烯烃和BTX。
另外，本公开中，通过在第一下行式反应器下部设置流化床反应器，可以将催化裂解后的物料中的轻烯烃进一步转化，可以最大化生产低碳烯烃。
本公开中，在第一下行式反应器中轻馏分油停留时间短、反应的生焦低、低碳烯烃和BTX收率高；另外，在流化床反应器中，轻烯烃馏分进一步被转化。由此，出第一下行式反应器的第一待生催化剂或者出流化床反应器的第二待生催化剂仍具有较高的活性，并且在该催化剂上负载有积炭，将该催化剂用于第二上行式反应器中的重馏分油的催化裂解时，可以提高低碳烯烃产率，抑制干气和焦炭的生成。
更为重要是，本公开中，通过使切割得到的轻馏分油和重馏分油的重量比(X)与第二股催化剂和连续催化剂的重量比(R)满足特定的关系，可以根据含烃原料油种类的不同，对于切割比例进行灵活调节，与此相应地，对于第二股催化剂和连续催化剂的重量比进行调节，使得第二上行式反应器中，催化剂活性与重馏分油的组成更加匹配，可以在最大化生产低碳烯烃和BTX的同时，可以明显降低干气和焦炭等副产物的产率。
另外，通过上述技术方案，本公开提供的含烃原料油催化裂解生产低碳烯烃和BTX的方法，可以明显提高低碳烯烃和轻芳烃的产率和装置经济性。
本公开的其他特征和优点将在随后的具体实施方式部分予以详细说明。
附图说明
附图是用来提供对本公开的进一步理解，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公开，但并不构成对本公开的限制。在附图中：
图1是本公开的装置的一个实施方式的示意图。
图2是本公开的装置的另一个实施方式的示意图。
附图标记说明
1.下行式反应器                 32.第二股催化剂(再生催化剂)输
                               送管
11.轻馏分油进料喷嘴            33.重馏分油进料喷嘴
12.第一股催化剂(再生催化剂)输  4.沉降器
送管
13.蘑菇头分布器                41.第三反应器油气出口
2.流化床反应器                 5.汽提器
                               51.第一汽提器
21.轻烯烃馏分进料喷嘴          52.第二汽提器
22.第一反应油气出口/第二反应   53.第三待生催化剂输送管
油气出口
3.上行式反应器                 6.再生器
31.连续催化剂输送管            7.气固分离器
具体实施方式
以下对本公开的具体实施方式进行详细说明。应当理解的是，此处所描述的具体实施方式仅用于说明和解释本公开，并不用于限制本公开。
在本文中所披露的任何具体数值(包括数值范围的端点)都不限于该数值的精确值，而应当理解为还涵盖了接近该精确值的值，例如在该精确值±5％范围内的所有可能的数值。并且，对于所披露的数值范围而言，在该范围的端点值之间、端点值与范围内的具体点值之间，以 及各具体点值之间可以任意组合而得到一个或多个新的数值范围，这些新的数值范围也应被视为在本文中具体公开。
除非另有说明，本文所用的术语具有与本领域技术人员通常所理解的相同的含义，如果术语在本文中有定义，且其定义与本领域的通常理解不同，则以本文的定义为准。
本申请中，除了明确说明的内容之外，未提到的任何事宜或事项均直接适用本领域已知的那些而无需进行任何改变。而且，本文描述的任何实施方式均可以与本文描述的一种或多种其他实施方式自由结合，由此形成的技术方案或技术思想均视为本公开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。
本公开提供了一种含烃原料油催化裂解生产低碳烯烃和轻芳烃的方法，该方法包括如下步骤：
S1、将含烃原料油切割为轻馏分油和重馏分油，所述轻馏分油相对于所述重馏分油的重量比(轻馏分油/重馏分油)为X；
S2、将所述轻馏分油与第一股催化剂引入第一下行式反应器，进行第一催化裂解，得到第一催化裂解后的物料；
任选的S2′、将所述第一催化裂解后的物料引入流化床反应器进行第二催化裂解，得到第二催化裂解后的物料；
S3、将所述第一催化裂解后的物料进行气固分离，得到第一反应油气和第一待生催化剂，或者将所述第二催化裂解后的物料进行气固分离，得到第二反应油气和第二待生催化剂；
S4、将连续催化剂、所述重馏分油与第二股催化剂引入第二上行式反应器，进行第三催化裂解，然后进行气固分离，得到第三反应油气和第三待生催化剂；所述连续催化剂为所述第一待生催化剂的至少一部分或所述第二待生催化剂的至少一部分；所述第二股催化剂与所述连续催化剂的重量比(第二股催化剂/连续催化剂)为R；
S5、从所述第一反应油气、所述第二反应油气和所述第三反应油气的任一者或者第一反应油气和第三反应油气的混合物或者第二反应油气和第三反应油气的混合物中分离出低碳烯烃和轻芳烃，并且分离出轻烯烃馏分，并将所述轻烯烃馏分返回步骤S4的所述第二上行式反应器中或者步骤S2′的所述流化床反应器中，
所述R和X满足以下关系式：
(4.84×T0-3340)/(780+5×T0-6×T3)＜R/X＜(0.968×T0-630)/(668+0.2×T0-1.2×T3)
T0为所述第二股催化剂进入步骤S4时的温度(单位℃)，T3为所述第二上行式反应器的出口温度(单位℃)。
本公开中，有时将所述第一反应油气、所述第二反应油气和所述第三反应油气的任一者或两者以上的混合物简称为反应油气。
本公开中，低碳烯烃是指乙烯、丙烯、丁烯及其异构体。轻芳烃是指BTX，即苯、甲苯和二甲苯。本公开中，低碳烯烃可以从干气、C3馏分和C4馏分中分离得到；轻芳烃可以从轻汽油和重汽油分离得到。
本公开中，C3馏分是指反应油气中的具有3个碳的烃类，包括丙烷、丙烯；C4馏分是指反应油气中的具有4个碳的烃类，包括丁烷、丁烯及其异构体；轻汽油是指反应油气中的馏程在30-90℃范围内的全部馏分或部分馏分，其中“部分馏分”是指馏程为30-90℃之间的一部分的温度范围的馏分(例如馏程为30-60℃或40-60℃或60-90℃等的馏分)；重汽油是指反应油气中的馏程在30-200℃范围内的馏分中除去轻汽油以外的馏分。
本公开中，轻馏分油、重馏分油是指对于含烃原料油在一定切割温度下进行切割后，切割出的轻质馏分称为轻馏分油、剩余部分称为重馏分油。本领域技术人员可以根据需要按照本领域公知的方法(包括但不限于分馏、蒸馏等方式)对含烃原料油进行切割，只要使得所述轻馏分油相对于所述重馏分油的重量比(轻馏分油/重馏分油)为X，并使X满足本公开的下述关系式即可。在本公开的一个实施方式中，X在选自0.1、0.2、0.3、0.4、0.5、0.6、0.7、0.8、0.9、1.0、2.0、3.0、4.0、5.0、6.0、7.0、8.0、9.0、10.0中任意两个数值范围之间。在本公开的一个实施方式中，X为0.1-2.0、优选为0.12-1.0、进一步优选为0.15-0.6。
在本公开的一个实施方式中，步骤S1中，在切割点100-400℃之间的任意温度将含烃原料油切割为轻馏分油和重馏分油，使得所述轻馏分油相对于所述重馏分油的重量比(轻馏分油/重馏分油)为X。在本公开的一个实施方式中，切割点例如为150℃、160℃、170℃、180℃、190℃、200℃、210℃、220℃、230℃、240℃、250℃、260℃、270℃、280℃、290℃、300℃、310℃、320℃、330℃、340℃、350℃、360℃、 370℃、380℃、390℃、400℃。
本公开中，所述含烃原料油可以为本领域已知的各类原料油(本发明中，有时将含烃原料油简称为原料油)，例如，可以为原油、煤液化油、合成油、油砂油、页岩油、致密油和动植物油脂中的一种或其两种以上的混合物，或其各自的部分馏分、其各自的重质馏分的加氢改质油。在本公开的一个实施方式中，含烃原料油优选为原油、原油的部分馏分、或来自原油的重油的加氢改质油。本领域技术人员已知的是，“部分馏分”可以通过对原料油进行本领域的常规处理，包括但不限于常压蒸馏、减压蒸馏等，来获得。本领域技术人员可以根据需要确定该常规处理的方式。在本公开的一个实施方式中，可以将原油作为本公开的含烃原料油，还可以根据需要将原油进行常压蒸馏或减压蒸馏，抽取部分馏分后的剩余馏分(原油的部分馏分)作为本公开的含烃原料油，或者根据需要将来自原油的重油进行加氢改质后的产物(重油的加氢改质油)作为本公开的含烃原料油。本领域中已知的是，加氢改质包括但不限于加氢脱硫、加氢脱氮、加氢脱金属、加氢饱和等处理。
在本公开的一个实施方式中，所述方法在步骤S1之前具有步骤S0，其中，将含烃原料油进行脱盐脱水处理，将得到的经脱水脱盐的含烃原料油引入步骤S1进行切割。
根据本公开，步骤S2中，在所述第一下行式反应器中，所述第一催化裂解的条件包括：所述第一下行式反应器的出口温度为610-720℃，优选650-690℃。所述第一催化裂解的条件还包括：气固停留时间为0.1-3.0秒，优选0.5-1.5秒。所述第一下行式反应器中，催化剂与轻馏分油的剂油比可以为催化裂解中的常用的剂油比(以催化剂/轻馏分油的重量比计)，例如可以为15-80，优选为25-65。
本公开中，对于轻馏分油与第一股催化剂引入第一下行式反应器的方式，没有任何限定，只要是在第一下行式反应器的上端引入轻馏分油与第一股催化剂即可。优选的是，由第一下行式反应器的不同进料口分别引入轻馏分油与第一股催化剂。
本公开中，对于第一股催化剂没有限定，可以为本领域已知的可用于原油催化裂解的催化剂即可。例如，所述第一股催化剂可以包括活性组分和载体，所述活性组分为选自含或不含稀土的超稳Y型沸石、 ZSM-5系列沸石、具有五元环结构的高硅沸石和β沸石中的至少一种。所述载体选自氧化铝、氧化硅、无定形硅铝、氧化锆、氧化钛、氧化硼和碱土金属氧化物中的至少一种。
在本公开的一个实施方式中，所述第一下行式反应器的结构没有特别限定，只要可以从其上部实现进料、下部实现出料即可，例如其可以是等径或变径的下行管反应器。
本公开中，由于在第一下行式反应器中不使用额外的热源，因此第一下行式反应器的出口温度即反映了反应器中的反应温度。本公开中，可以通过调节第一股催化剂的温度、反应器中的气固停留时间、第一下行式反应器的出口温度等条件来调节第一下行式反应器中轻馏分油的催化裂解程度。
在本公开的一个实施方式中，所述第一股催化剂为新鲜催化剂。在本公开的一个实施方式中，所述第一股催化剂包括来自再生器的再生催化剂。优选的是，所述第一股催化剂为来自再生器的再生催化剂。
本公开中，对于进入下行式反应器的第一股催化剂的温度没有特别限定，只要其可以与轻馏分油接触时发生催化裂解、并且满足本公开上述的第一催化裂解条件即可。在使用再生催化剂作为第一股催化剂时，经由第一股催化剂(再生催化剂)输送管，由再生器直接进料第一股催化剂，由于再生器与第一下行式反应器之间的输送管短，因此，所述第一股催化剂的温度可以视为再生器的温度或再生催化剂离开再生器时的温度(再生器出口的温度)。本公开的一个实施方式中，进入下行式反应器的第一股催化剂的温度为再生器的温度或再生催化剂离开再生器时的温度(再生器出口的温度)，通常可以为690-750℃，优选为700-740℃，进一步优选为705-730℃，更进一步优选为710-725℃。另外，此时还可以根据需要，对于来自再生器的催化剂进一步加热或冷却后再使其进料到第一下行式反应器中。在本公开的方法中，在开车时，可以将新鲜催化剂加热到所需温度，再引入第一下行式反应器；此后，可以直接采用来自再生器的再生催化剂。本公开的一个实施方式中，优选由再生器直接进料第一股催化剂，而不对其进行进一步加热或冷却。
本公开中，在将轻馏分油引入第一下行式反应器时，还可以根据需要先对轻馏分油进行预热。预热后的轻馏分油的温度例如为30-100℃。 另外，还可以先对轻馏分油进行水蒸气雾化，然后再以水蒸气作为载体将轻馏分油引入第一下行式反应器。
本公开中，第一催化裂解后的物料包含将轻馏分油进行催化裂解后的得到的第一反应油气和第一股催化剂被焦化(碳化后的)第一待生催化剂。该第一待生催化剂仍具有较高的活性，并且在该催化剂上负载有积炭，将其作为连续催化剂引入后续的第二上行式反应器中时，有助于重馏分油的催化裂解，提高低碳烯烃产率，抑制干气和焦炭的生成。
本公开中，步骤S3中，将第一催化裂解后的物料进行气固分离，得到第一反应油气和第一待生催化剂。作为气固分离的方式，没有特别限定，可以采用本领域中公知的方式，例如采用沉降器、旋风分离器实现催化剂与第一反应油气的分离。
在本公开的一个实施方式中，将第一反应油气进行分离，得到干气、C3馏分、C4馏分、轻汽油、重汽油、柴油和油浆，从中分离得到低碳烯烃和轻芳烃，并分离出轻烯烃馏分。其中，C4馏分和/或轻汽油即为所述轻烯烃馏分。在本公开的一个实施方式中，将第一反应油气引入分馏装置或气分装置进行分馏，以实现上述分离。在本公开的一个实施方式中，将所述轻烯烃馏分引入下述步骤S4中的第二上行式反应器。
在本公开的一个实施方式中，将第一待生催化剂的至少一部分作为连续催化剂引入下述第二上行式反应器。在本公开的一个实施方式中，将未进入下述第二上行式反应器的第一待生催化剂引入再生步骤，在其中进行催化剂的再生。优选的是，将全部第一待生催化剂作为连续催化剂引入下述第二上行式反应器，此时作为连续催化剂的第一待生催化剂的量基本上对应于第一股催化剂的量。
在本公开的一个实施方式中，将第一催化裂解后的物料进行气固分离，对分离的催化剂，进一步进行汽提除去其中吸附的烃类产物，得到第一待生催化剂。
在本公开的一个实施方式中，还可以在步骤S2之后、步骤S3之前，包括步骤S2′，其中，将所述第一催化裂解后的物料引入流化床反应器进行第二催化裂解，得到第二催化裂解后的物料，由此，可以将轻烯烃馏分进一步转化，可以最大化生产低碳烯烃。
在本公开中，“流化床反应器”也称为“流态化反应器”，其催化剂密 度在150-450kg/m 3之间。
根据本公开，步骤S2′中，在所述流化床反应器中，所述第二催化裂解的条件包括：所述流化床反应器中的反应温度为600-690℃，优选640-670℃。第二催化裂解的条件还包括：质量空速为2-20h -1，优选4-12h -1。
根据本公开的一个实施方式，在流化床中不用引入新的催化剂，而直接引入第一催化裂解后的物料进行催化裂解即可。根据本公开的一个实施方式，不对流化床施加额外的热源，而直接利用第一催化裂解后的物料的热量即可。所引入的第一催化裂解后的物料包含将轻馏分油进行催化裂解后的得到的第一反应油气和第一股催化剂被焦化(碳化后的)第一待生催化剂。该第一待生催化剂仍具有较高的活性，可以在流化床反应器中继续加深催化裂解的程度，将轻烯烃馏分进一步转化为低碳烯烃。
根据本公开的一个实施方式，从本公开的反应油气中分离出轻烯烃馏分，并将所述轻烯烃馏分返回所述流化床反应器中，以将其进一步转化为低碳烯烃。更具体而言，将反应油气进行分离，得到干气、C3馏分、C4馏分、轻汽油、重汽油、柴油和油浆，从中分离得到低碳烯烃和轻芳烃，并分离出轻烯烃馏分。其中，C4馏分和/或轻汽油即为所述轻烯烃馏分。在本公开的一个实施方式中，将反应油气引入分馏装置或气分装置，以实现上述分离。
本公开中，第二催化裂解后的物料包含第二反应油气和第二待生催化剂。该第二待生催化剂仍具有较高的活性，并且在该催化剂上负载有积炭，将其作为连续催化剂引入后续的第二上行式反应器中时，有助于重馏分油的催化裂解，提高低碳烯烃产率，抑制干气和焦炭的生成。
本公开中，步骤S3中，将第二催化裂解后的物料进行气固分离，得到第二反应油气和第二待生催化剂。作为气固分离的方式，没有特别限定，可以采用本领域中公知的方式，例如采用沉降器、旋风分离器实现催化剂与第二反应油气的分离。
在本公开的一个实施方式中，将第二反应油气进行分离，得到干气、C3馏分、C4馏分、轻汽油、重汽油、柴油和油浆，从中分离得到低碳烯烃和轻芳烃，并分离出轻烯烃馏分。其中，C4馏分和/或轻汽油即为 轻烯烃馏分。在本公开的一个实施方式中，将第二反应油气引入分馏装置或气分装置，以实现上述分离。
本公开的一个实施方式中，将第二催化裂解后的物料进行气固分离，对分离的催化剂，进一步进行汽提除去其中吸附的烃类产物，得到第二待生催化剂。本公开中，第二待生催化剂的至少一部分作为连续催化剂引入下述第二上行式反应器。
在本公开的一个实施方式中，将未进入下述第二上行式反应器的第二待生催化剂引入再生步骤，在其中进行催化剂的再生。优选的是，将全部第二待生催化剂作为连续催化剂引入下述第二上行式反应器，此时作为连续催化剂的第二待生催化剂的量基本上对应于第一股催化剂的量。
本公开中，步骤S4中，将连续催化剂、所述重馏分油与第二股催化剂引入第二上行式反应器，进行第三催化裂解，然后进行气固分离，得到第三反应油气和第三待生催化剂；所述连续催化剂为所述第一待生催化剂的至少一部分或所述第二待生催化剂的至少一部分。
本公开的一个实施方式中，所述第二上行式反应器中，所述第三催化裂解的条件包括：所述第二上行式反应器的出口温度T3为530-650℃，优选560-640℃，进一步优选为580-630℃，更进一步优选600-630℃。第三催化裂解的条件还包括：气固停留时间为0.5-8秒，优选1.5-5秒。所述第二上行式反应器中，催化剂与重馏分油的剂油比可以为催化裂解中的常用的剂油比(以催化剂/重馏分油的重量比计)，例如可以为8-40，优选为10-30。
本公开的一个实施方式中，在步骤S4中，首先使所述连续催化剂与所述第二股催化剂混合，再进行后续催化裂解反应。更具体而言，本公开的一个实施方式中，使所述连续催化剂与所述第二股催化剂各自独立地进料到第二上行式反应器的底部，将其混合，混合后的催化剂(以下，有时也称为催化剂混合物或混合催化剂)用于第二上行式反应器中的催化裂解反应。在本公开的一个实施方式中，使所述连续催化剂与所述第二股催化剂在第二上行式反应器的底部区域混合后，使用预提升介质使混合催化剂在第二上行式反应器中提升，进行下游的催化裂解反应。在本公开的一个实施方式中，所述预提升介质可以为干气、水蒸气或它们的混合物。
本公开的一个实施方式中，在步骤S4中，对于第二股催化剂没有限定，可以为本领域已知的可用于原油催化裂解的催化剂即可。例如，所述第二股催化剂包括活性组分和载体，所述活性组分为选自含或不含稀土的超稳Y型沸石、ZSM-5系列沸石、具有五元环结构的高硅沸石和β沸石中的至少一种。所述载体选自氧化铝、氧化硅、无定形硅铝、氧化锆、氧化钛、氧化硼和碱土金属氧化物中的至少一种。
在本公开的一个实施方式中，所述第二上行式反应器的结构没有特别限定，只要可以从其底部实现进料、从上部实现出料即可，例如其可以为等径或变径的提升管反应器，等径或变径的提升管反应器和流化床复合反应器。
在本公开的一个实施方式中，所述第二股催化剂为新鲜催化剂。在本公开的一个实施方式中，所述第二股催化剂包括来自再生器的再生催化剂。在本公开的一个实施方式中，所述第二股催化剂为来自再生器的再生催化剂。当使用新鲜催化剂作为第二股催化剂时，需要对该催化剂进行预热，使得该新鲜催化剂进入步骤S4时的温度满足本公开的关系式。优选的是，所述第二股催化剂为来自再生器的再生催化剂。
本公开中，所述第二股催化剂与所述连续催化剂的重量比(第二股催化剂/连续催化剂)为R，并且，使所述R和X满足以下关系式：
(4.84×T0-3340)/(780+5×T0-6×T3)＜R/X＜(0.968×T0-630)/(668+0.2×T0-1.2×T3)
T0为所述第二股催化剂进入步骤S4时的温度(单位℃)，T3为所述第二上行式反应器的出口温度(单位℃)。
本公开的发明人惊讶地发现，通过使步骤S1中的含烃原料油的轻馏分油和重馏分油的切割比例(轻馏分油/重馏分油的重量比)与所述第二股催化剂和所述连续催化剂的重量比(第二股催化剂/连续催化剂)满足上述关系式，可以使得含烃原料油组分、切割比例、催化剂活性(特别是第二上行式反应器中的催化剂活性)更好地匹配，可以在最大化生产低碳烯烃和BTX的同时，可以明显降低干气和焦炭的产率。不受任何理论限定地，本公开的发明人推测，作为连续催化剂的催化剂来自第一待生催化剂或第二待生催化剂，由于在第一下行式反应器中和流化床反应器中生焦低，使得第一待生催化剂和第二待生催化剂具有较高的催化活性的同时，并负载一定量的积炭，这种催化剂以一定比例 与第二股催化剂(新鲜催化剂或来自再生器的再生催化剂)进行混合，并使混合比例与含烃原料油切割比例相匹配，满足本公开的上述关系式，所得的催化剂保持优异的催化活性的同时，不会引发由于催化剂活性过高导致的重馏分油的过度焦化，也不会引发由于催化剂活性过低导致的重馏分油的催化裂化不充分。本公开中，使含烃原料油切割为轻馏分油和重馏分油的比例与连续催化剂和第二股催化剂的混合比例满足特定关系，可以根据含烃原料油的组成、切割比例等来调节第二上行式反应器中的混合催化剂的活性，使得来自重馏分油的低碳烯烃和BTX的产率最大化。
本公开的一个实施方式中，(4.84×T0-3340)/(780+5×T0-6×T3)大于0。本公开中，T0大于T3。
本公开中，T0为所述第二股催化剂进入步骤S4时的温度。具体而言，是指第二股催化剂(新鲜催化剂或再生催化剂)进入第二上行式反应器时的温度，即，其进入第二上行式反应器底部时、在与连续催化剂混合前的温度。当使用再生催化剂作为第二股催化剂时，由于再生器和第二上行式反应器之间的输送管短，因此再生器的温度或再生催化剂出再生器时催化剂的温度(再生器出口的温度)可以视为第二股催化剂进入步骤S4时的温度。
本公开的一个实施方式中，所述第二上行式反应器的出口温度T3为530-650℃，优选为560-640℃，进一步优选为580-630℃，更进一步优选为600-630℃；和/或，所述第二股催化剂进入步骤S4时的温度T0为690-750℃，优选700-740℃，进一步优选为705-730℃，更进一步优选为710-725℃。
在本公开的一个实施方式中，将第三反应油气进行分离，得到干气、C3馏分、C4馏分、轻汽油、重汽油、柴油和油浆，从中分离得到低碳烯烃和轻芳烃，并分离出轻烯烃馏分。其中，C4馏分和/或轻汽油即为轻烯烃馏分。在本公开的一个实施方式中，将第三反应油气引入分馏装置或气分装置，以实现上述分离。
本公开中，不存在步骤S2′时，在步骤S5中，从第一反应油气、第三反应油气的任一者或者两者的混合物中分离得到低碳烯烃和轻芳烃，并将分离出的轻烯烃馏分返回所述第二上行式反应器中。
本公开中，存在步骤S2′时，在步骤S5中，从第二反应油气、第 三反应油气的任一者或者两者的混合物中分离得到低碳烯烃和轻芳烃，并将分离出的轻烯烃馏分返回所述流化床反应器中。
本公开的一个实施方式中，在步骤S4中，来自下述S5步骤的轻烯烃馏分先于所述重馏分油与所述催化剂混合物接触，发生催化裂解反应，而后重馏分油再与催化剂混合物接触，发生催化裂解反应。优选的是，所述轻烯烃馏分先于所述重馏分油0.3-1.0秒与所述催化剂混合物接触。更优选的是，所述轻烯烃馏分先于所述重馏分油0.4-0.8秒与所述催化剂混合物接触。
本公开的一个实施方式中，对第三催化裂解的产物进行气固分离，得到第三反应油气和第三待生催化剂。作为气固分离的方式，没有特别限定，可以采用本领域中公知的方式，例如采用沉降器、旋风分离器实现催化剂与第三反应油气的分离。
在本公开的一个实施方式中，将第三催化裂解后的物料进行气固分离，对分离的催化剂，进一步进行汽提除去其中吸附的烃类产物，得到第三待生催化剂。在本公开的一个实施方式中，使第三待生催化剂进入再生器进行催化剂的再生。
在本公开的一个实施方式中，再生器的温度为本领域常用的温度，其可以为690-750℃，优选为700-740℃，进一步优选为705-730℃，更进一步优选为710-725℃。在本公开的一个实施方式中，再生器的温度或再生催化剂出再生器时催化剂的温度(再生器出口的温度)可以视为第二股催化剂进入步骤S4时的温度。因此，在本公开的一个实施方式中，第二股催化剂进入步骤S4时的温度T0可以为690-750℃，优选为700-740℃，进一步优选为705-730℃，更进一步优选为710-725℃。
在本公开的一个实施方式中，再生后的催化剂用作第一股催化剂和第二股催化剂。
在本公开中，步骤S5中，从所述第一反应油气、所述第二反应油气和所述第三反应油气的任一者或者第一反应油气和第三反应油气的混合物或者第二反应油气和第三反应油气的混合物中分离得到低碳烯烃和轻芳烃，并且分离出轻烯烃馏分，并将所述轻烯烃馏分返回步骤S4的所述第二上行式反应器中或者步骤S2′的所述流化床反应器中。更具体而言，将反应油气进行分离，得到干气、C3馏分、C4馏分、轻汽油、重汽油、柴油和油浆，从中分离得到低碳烯烃和轻芳烃，并分离 出轻烯烃馏分。其中，C4馏分和/或轻汽油即为所述轻烯烃馏分。优选的是，将反应油气引入分馏装置或气分装置，以实现上述分离。步骤S5中，可以将所述第一反应油气、所述第三反应油气各自进行分离，也可以将二者合并后统一进行分离；或者，可以将所述第二反应油气、所述第三反应油气各自进行分离，也可以将二者合并后统一进行分离。
在本公开的一个实施方式中，从反应油气中分离出轻烯烃馏分的方法没有限定，可以采用本领域已知的方式进行分离，包括但不限于以下方法，反应油气进入分馏、吸收稳定单元后，分离出液化气和稳定汽油，液化气进入后续气分装置分离出C3馏分和C4馏分，稳定汽油进入轻重汽油分割塔，分离出轻汽油和重汽油。C4馏分和/或轻汽油即为所述轻烯烃馏分。从中可以分离得到低碳烯烃和轻芳烃。
更详细而言，在本公开的一个实施方式中，提供一种含烃原料油催化裂解生产低碳烯烃和轻芳烃的方法，该方法包括如下步骤：
S1、将含烃原料油切割为轻馏分油和重馏分油，所述轻馏分油相对于所述重馏分油的重量比(轻馏分油/重馏分油)为X；
S2、将所述轻馏分油与第一股催化剂引入第一下行式反应器，进行第一催化裂解，得到第一催化裂解后的物料；
S3、将所述第一催化裂解后的物料进行气固分离，得到第一反应油气和第一待生催化剂；
S4、将连续催化剂、所述重馏分油与第二股催化剂引入第二上行式反应器，进行第三催化裂解，然后进行气固分离，得到第三反应油气和第三待生催化剂；所述连续催化剂为所述第一待生催化剂的至少一部分；所述第二股催化剂与所述连续催化剂的重量比(第二股催化剂/连续催化剂)为R；
S5、从所述第一反应油气和所述第三反应油气的任一者或者两者的混合物中分离出低碳烯烃和轻芳烃，并且分离出轻烯烃馏分，并将所述轻烯烃馏分返回步骤S4的所述第二上行式反应器中，
所述R和X满足以下关系式：
(4.84×T0-3340)/(780+5×T0-6×T3)＜R/X＜(0.968×T0-630)/(668+0.2×T0-1.2×T3)
T0为所述第二股催化剂进入步骤S4时的温度(单位℃)，T3为所述第二上行式反应器的出口温度(单位℃)。
更详细而言，在本公开的一个实施方式中，提供含烃原料油催化裂解生产低碳烯烃和轻芳烃的方法，该方法包括如下步骤：
S1、将含烃原料油切割为轻馏分油和重馏分油，所述轻馏分油相对于所述重馏分油的重量比(轻馏分油/重馏分油)为X；
S2、将所述轻馏分油与第一股催化剂引入第一下行式反应器，进行第一催化裂解，得到第一催化裂解后的物料；
S2′、将所述第一催化裂解后的物料引入流化床反应器进行第二催化裂解，得到第二催化裂解后的物料；
S3、将所述第二催化裂解后的物料进行气固分离，得到第二反应油气和第二待生催化剂；
S4、将连续催化剂、所述重馏分油与第二股催化剂引入第二上行式反应器，进行第三催化裂解，然后进行气固分离，得到第三反应油气和第三待生催化剂；所述连续催化剂为所述第二待生催化剂的至少一部分；所述第二股催化剂与所述连续催化剂的重量比(第二股催化剂/连续催化剂)为R；
S5、从所述第二反应油气和所述第三反应油气的任一者或者两者的混合物中分离出低碳烯烃和轻芳烃，并且分离出轻烯烃馏分，并将所述轻烯烃馏分返回步骤S2′的所述流化床反应器中，
所述R和X满足以下关系式：
(4.84×T0-3340)/(780+5×T0-6×T3)＜R/X＜(0.968×T0-630)/(668+0.2×T0-1.2×T3)
T0为所述第二股催化剂进入步骤S4时的温度(单位℃)，T3为所述第二上行式反应器的出口温度(单位℃)。
在本公开的一个实施方式中，所述轻烯烃馏分为所述反应油气中的C4馏分和/或所述轻汽油。
本公开还提供以下技术方案：
A1、含烃原料油催化裂化生产低碳烯烃和轻芳烃的方法，该方法包括如下步骤：
S1、将脱盐脱水的含烃原料油切割为轻馏分油和重馏分油；所述切割的切割点为100-400℃之间的任意温度；
S2、将所述轻馏分油与第一股催化剂引入第一下行式反应器，进行第一催化裂解，得到第一催化裂解后的物料；
任选的S2′、将所述第一催化裂解后的物料送入流化床反应器进行第二催化裂解，得到第二催化裂解后的物料；
S3、将所述第一催化裂解后的物料进行气固分离，得到第一反应油气和第一待生催化剂，或者将所述第二催化裂解后的物料进行气固分离，得到第二反应油气和第二待生催化剂；
S4、将连续催化剂、所述重馏分油与第二股催化剂引入第二上行式反应器，进行第三催化裂解，然后进行气固分离，得到第三反应油气和第三待生催化剂；所述连续催化剂为所述第一待生催化剂或所述第二待生催化剂；所述第二股催化剂与所述连续催化剂的重量比为0.2-5∶1；
S5、从所述第一反应油气和所述第二反应油气中分离出轻烯烃馏分，并将所述轻烯烃馏分返回所述流化床反应器中或所述第二上行式反应器中。
A2、根据A1所述的方法，其中，步骤S1中，所述切割的切割点为200-380℃之间的任意温度。
A3、根据A1所述的方法，其中，步骤S4中，所述第二股催化剂与所述连续催化剂的重量比为0.5-3∶1。
A4、根据权利要求A1所述的方法，其中，
所述第一下行式反应器中，所述第一催化裂化的条件包括：所述第一下行式反应器的出口温度为610-720℃，气固停留时间为0.1-3.0秒；
所述流化床反应器中，所述第二催化裂化的条件包括：所述流化床反应器中的反应温度为600-670℃，质量空速为2-20h -1；
所述第二上行式反应器中，所述第三催化裂化的条件包括：所述第二上行式反应器的出口温度为530-650℃，气固停留时间为0.5-8秒。
A5、根据A4所述的方法，其中，
所述第一下行式反应器中，所述第一催化裂化的条件包括：所述第一下行式反应器的出口温度为650-690℃，气固停留时间为0.5-1.5秒；
所述流化床反应器中，所述第二催化裂化的条件包括：所述流化床反应器中的反应温度为620-640℃，质量空速为4-12h -1；
所述第二上行式反应器中，所述第三催化裂化的条件包括：所述 第二上行式反应器的出口温度为560-640℃，气固停留时间为1.5-5秒。
A6、根据A1所述的方法，其中，
将所述轻烯烃馏分先于所述重馏分油0.3-1.0秒与所述第二股催化剂进行催化裂解；优选将所述轻烯烃馏分先于所述重馏分油0.4-0.8秒与所述第二股催化剂进行催化裂解。
A7、根据A1所述的方法，其中，该方法还包括：
将所述第三待生催化剂进行烧焦再生，得到再生催化剂；
将所述第一油气和所述第二油气进行分离，得到干气、C3馏分、C4馏分、轻汽油、重汽油、柴油和油浆；
所述轻烯烃馏分为所述第一油气和所述第二油气中的C4馏分和/或所述第一油气和所述第二油气中30-90℃范围内的馏分。
A8、根据A1所述的方法，其中，所述含烃原料油为常规矿物油、煤液化油、合成油、油砂油、页岩油、致密油和动植物油脂中的一种或几种的混合物。
A9、根据A1所述的方法，其中，所述第一股催化剂和所述第二股催化剂各自独立地包括活性组分和载体，所述活性组分为选自含或不含稀土的超稳Y型沸石、ZSP系列沸石、具有五元环结构的高硅沸石和β沸石中的至少一种。
A10、根据A1所述的方法，其中，所述第一股催化剂和所述第二股催化剂各自独立地包括所述再生催化剂。
本公开还提供一种含烃原料油催化裂解生产低碳烯烃和轻芳烃的装置，该装置包括以下单元：
含烃原料油切割单元，在其中将含烃原料油切割为轻馏分油和重馏分油，使得所述轻馏分油相对于所述重馏分油的重量比(轻馏分油/重馏分油)为X，
第一下行式反应单元，从该反应单元的上方引入所述轻馏分油与第一股催化剂，进行第一催化裂解，在该反应单元的下方得到第一催化裂解后的物料；
任选的流化床反应单元，其中，引入所述第一催化裂解后的物料，进行第二催化裂解，得到第二催化裂解后的物料；
第一气固分离单元，其中引入所述第一催化裂解后的物料进行气固 分离，得到第一反应油气和第一待生催化剂，或者其中引入所述第二催化裂解后的物料进行气固分离，得到第二反应油气和第二待生催化剂；
第二上行式反应单元，从该反应单元的下方引入连续催化剂、第二股催化剂与所述重馏分油，进行第三催化裂解，在该反应单元的上方得到第三催化裂解后的物料，所述连续催化剂为所述第一待生催化剂的至少一部分或所述第二待生催化剂的至少一部分，所述第二股催化剂与所述连续催化剂的重量比(第二股催化剂/连续催化剂)为R，
第二气固分离单元，其中引入所述第三催化裂解后的物料进行气固分离，得到第三反应油气和第三待生催化剂；
分离单元，其中引入所述第一反应油气、所述第二反应油气和所述第三反应油气的任一者或者第一反应油气和第三反应油气的混合物或者第二反应油气和第三反应油气的混合物，分离出低碳烯烃和轻芳烃，并且分离出轻烯烃馏分，并将轻烯烃馏分返回至所述第二上行式反应单元或者所述流化床反应单元；
其中，所述R和X满足以下关系式：
(4.84×T0-3340)/(780+5×T0-6×T3)＜R/X＜(0.968×T0-630)/(668+0.2×T0-1.2×T3)
T0为所述第二股催化剂进入第二上行式反应单元时的温度(单位℃)，T3为所述第二上行式反应单元的出口温度(单位℃)。
本公开中，T0为所述第二股催化剂进入第二上行式反应单元时的温度(单位℃)。具体而言，是指第二股催化剂进入第二上行式反应单元底部时、在与连续催化剂混合前的温度。
本公开的一个实施方式中，所述装置还包括再生单元，其中，引入所述第三待生催化剂和任选的未进入第二上行式反应器的第一待生催化剂或第二待生催化剂，进行烧焦再生，得到再生催化剂。优选的是，向再生单元中只引入所述第三待生催化剂。本公开的一个实施方式中，再生单元的温度为本领域常用的温度，其可以为690-750℃，优选为700-740℃，进一步优选为705-730℃，更进一步优选为710-725℃。
本公开的一个实施方式中，所述第二上行式反应单元的出口温度T3为530-650℃，优选为560-640℃，进一步优选为580-630℃，更进一步优选为600-630℃。
本公开的一个实施方式中，所述第二股催化剂进入第二上行式反应单元时的温度T0为690-750℃，优选700-740℃，进一步优选为705-730℃，更进一步优选为710-725℃。
本公开的一个实施方式中，所述装置还包括脱水脱盐单元，其中，将含烃原料油进行脱盐脱水处理，将得到的经脱水脱盐的含烃原料油引入含烃原料油切割单元进行切割。
在本公开的一个实施方式中，所述第一下行式反应单元的结构没有特别限定，只要可以从其上部实现进料、下部实现出料即可，例如其可以是等径或变径的下行管反应器。
在本公开的一个实施方式中，不存在流化床反应单元时，在分离单元中，从第一反应油气、第三反应油气的任一者或者两者的混合物中分离得到低碳烯烃和轻芳烃，并将分离出的轻烯烃馏分返回所述第二上行式反应单元中。
本公开中，存在流化床反应单元时，在分离单元中，从第二反应油气、第三反应油气的任一者或者两者的混合物中分离得到低碳烯烃和轻芳烃，并将分离出的轻烯烃馏分返回所述流化床反应器中。
本公开的一个实施方式中，所述第一气固分离单元、所述第二气固分离单元包括本领域公知的可以实现气固分离的设备，例如可以包括沉降器或旋风分离器。
本公开的一个实施方式中，所述装置还包括至少一个汽提单元，其可以设置在气固分离单元中，其中，将气固分离得到的催化剂进行汽提，以除去其中吸附的烃类产物。
更具体而言，本公开的一个实施方式中，所述装置在包括流化床反应单元时，第一气固分离单元中还包括汽提单元，其中，将气固分离得到的催化剂进行汽提，以除去其中吸附的烃类产物，得到第二待生催化剂。本公开的一个实施方式中，所述第二气固分离单元中还包括汽提单元，其中，将气固分离得到的催化剂进行汽提，以除去其中吸附的烃类产物，得到第三待生催化剂。
本公开的一个实施方式中，将连续催化剂和第二股催化剂引入第二上行式反应单元的底部，进行混合后，混合催化剂用于后续的催化裂解反应。
本公开的一个实施方式中，第二上行式反应单元中，引入连续催化 剂和第二股催化剂的位置在轻烯烃馏分进料口的上游。
本公开的一个实施方式中，第二上行式反应单元中，轻烯烃馏分进料口在重馏分油进料口的上游。
在本公开的一个实施方式中，所述第二上行式反应器的结构没有特别限定，只要可以从其底部实现进料、从上部实现出料即可，例如其可以为等径或变径的提升管反应器，等径或变径的提升管反应器和流化床复合反应器。
更详细而言，本公开提供一种含烃原料油催化裂解生产低碳烯烃和轻芳烃的装置，该装置包括以下单元：
含烃原料油切割单元，在其中将含烃原料油切割为轻馏分油和重馏分油，使得所述轻馏分油相对于所述重馏分油的重量比(轻馏分油/重馏分油)为X，
第一下行式反应单元，从该反应单元的上方引入所述轻馏分油与第一股催化剂，进行第一催化裂解，在该反应单元的下方得到第一催化裂解后的物料；
第一气固分离单元，其中引入所述第一催化裂解后的物料进行气固分离，得到第一反应油气和第一待生催化剂；
第二上行式反应单元，从该反应单元的下方引入连续催化剂、第二股催化剂与所述重馏分油，进行第三催化裂解，在该反应单元的上方得到第三催化裂解后的物料，所述连续催化剂为所述第一待生催化剂的至少一部分，所述第二股催化剂与所述连续催化剂的重量比(第二股催化剂/连续催化剂)为R，
第二气固分离单元，其中引入所述第三催化裂解后的物料进行气固分离，得到第三反应油气和第三待生催化剂；
分离单元，其中引入所述第一反应油气和所述第三反应油气的任一者或两者的混合物，分离出低碳烯烃和轻芳烃，并且分离出轻烯烃馏分，并将轻烯烃馏分返回至所述第二上行式反应单元；
其中，所述R和X满足以下关系式：
(4.84×T0-3340)/(780+5×T0-6×T3)＜R/X＜(0.968×T0-630)/(668+0.2×T0-1.2×T3)
T0为所述第二股催化剂进入第二上行式反应单元时的温度(单位℃)，T3为所述第二上行式反应单元的出口温度(单位℃)。
更详细而言，本公开提供一种含烃原料油催化裂解生产低碳烯烃和轻芳烃的装置，该装置包括以下单元：
含烃原料油切割单元，在其中将含烃原料油切割为轻馏分油和重馏分油，使得所述轻馏分油相对于所述重馏分油的重量比(轻馏分油/重馏分油)为X，
第一下行式反应单元，从该反应单元的上方引入所述轻馏分油与第一股催化剂，进行第一催化裂解，在该反应单元的下方得到第一催化裂解后的物料；
流化床反应单元，其中，引入所述第一催化裂解后的物料，进行第二催化裂解，得到第二催化裂解后的物料；
第一气固分离单元，其中引入所述第二催化裂解后的物料进行气固分离，得到第二反应油气和第二待生催化剂；
第二上行式反应单元，从该反应单元的下方引入连续催化剂、第二股催化剂与所述重馏分油，进行第三催化裂解，在该反应单元的上方得到第三催化裂解后的物料，所述连续催化剂为所述第二待生催化剂的至少一部分，所述第二股催化剂与所述连续催化剂的重量比(第二股催化剂/连续催化剂)为R，
第二气固分离单元，其中引入所述第三催化裂解后的物料进行气固分离，得到第三反应油气和第三待生催化剂；
分离单元，其中引入所述第二反应油气和所述第三反应油气的任一者或者两者的混合物，分离出低碳烯烃和轻芳烃，并且分离出轻烯烃馏分，并将轻烯烃馏分返回至所述流化床反应单元；
其中，所述R和X满足以下关系式：
(4.84×T0-3340)/(780+5×T0-6×T3)＜R/X＜(0.968×T0-630)/(668+0.2×T0-1.2×T3)
T0为所述第二股催化剂进入第二上行式反应单元时的温度(单位℃)，T3为所述第二上行式反应单元的出口温度(单位℃)。
本公开的含烃原料油催化裂解生产低碳烯烃和轻芳烃的装置用于实施本公开的含烃原料油催化裂解生产低碳烯烃和轻芳烃的方法。
以下，参照图1和图2，分别详细说明本公开的两种实施方式，但本公开不限于此。
本公开的一种具体的实施方式如图1所示，热的第一股催化剂(再 生催化剂)通过第一股催化剂输送管(再生催化剂输送管)12向第一下行式反应器1输送。轻馏分油通过进料喷嘴11喷入第一下行式反应器1中，与第一股催化剂接触并进行催化裂解反应，反应后的第一催化裂解后的物料在气固分离器7进行催化剂与油气分离，所得第一反应油气通过下行式反应器油气出口22引入分离装置(图中未示出)，第一待生催化剂通过连续催化剂输送管31作为连续催化剂引入第二上行式反应器3底部，第二股催化剂(再生催化剂)通过第二股催化剂输送管(再生催化剂输送管)32引入第二上行式反应器3底部，第一待生催化剂(连续催化剂)和第二股催化剂混合后的催化剂通过预提升介质向上提升。轻烯烃馏分通过轻烯烃馏分进料喷嘴21喷入第二上行式反应器3，与催化剂接触并发生反应。重馏分油通过重馏分油进料喷嘴33喷入第二上行式反应器3与来自底部的油剂混合物接触反应，反应后得到第三催化裂解后的物料，使其进入沉降器4，在沉降器4中进行第三待生催化剂与第三反应油气的分离，第三反应油气通过第三反应器油气出口41进入分离装置(图中未示出)，第三待生催化剂进入汽提器5，汽提出吸附的烃类产物，由输送管53将第三待生催化剂送至再生器6进行再生，再生后的催化剂返回第一下行式反应器和第二上行式反应器重复使用。将反应油气(第一反应油气、第三反应油气)经分离装置(优选为分馏装置、气分装置)分离得到干气、C3馏分、C4馏分、轻汽油、重汽油、柴油、油浆，从中分离得到低碳烯烃和轻芳烃。另外，从反应油气中分离出轻烯烃馏分，并将所述轻烯烃馏分由轻烯烃馏分进料喷嘴21引入所述第二上行式反应器3中。
本公开的另一种具体的实施方式如图2所示，热的第一股催化剂(再生催化剂)通过第一股催化剂输送管(再生催化剂输送管)12向第一下行式反应器1输送。轻馏分油通过进料喷嘴11喷入第一下行式反应器1中，与第一股催化剂接触并进行催化裂解反应，反应后的第一催化裂解后的物料通过第一下行式反应器的出口蘑菇头分布器13引入流化床反应器2中，在流化床反应器内继续发生裂解反应，反应后得到第二催化裂解后物料，将该物料进行旋风分离后得到第二反应油气和第二待生催化剂。所述第二反应油气由第二反应油气出口22引入分离装置(图中未示出)，第二待生催化剂进入第一汽提器51汽提出吸附的烃类产物后，作为连续催化剂通过连续催化剂输送管31引入第二上行 式反应器3底部，第二股催化剂(再生催化剂)通过第二股催化剂输送管(再生催化剂输送管)32引入第二上行式反应器3底部，第二待生催化剂(连续催化剂)和第二股催化剂混合后的催化剂通过预提升介质向上提升。重馏分油通过重馏分油进料喷嘴33喷入第二上行式反应器3与催化剂接触反应，反应后得到第三催化裂解后的物料，使其进入沉降器4，在沉降器4中进行第三待生催化剂与第三反应油气的分离，第三反应油气通过第三反应器油气出口41引入分离装置(图中未示出)，第三待生催化剂进入第二汽提器52，汽提出吸附的烃类产物，由第三待生催化剂输送管53将第三待生催化剂送至再生器6进行再生，再生后的催化剂返回第一下行式反应器和第二上行式反应器重复使用。将反应油气(第二反应油气、第三反应油气)经分离装置(优选为分馏装置、气分装置)分离得到干气、C3馏分、C4馏分、轻汽油、重汽油、柴油、油浆，从中分离得到低碳烯烃和轻芳烃。另外，从反应油气中分离出轻烯烃馏分，并将所述轻烯烃馏分由轻烯烃馏分进料喷嘴21返回流化床反应器2中。
实施例
以下通过实施例进一步详细说明本公开。实施例中所用到的原材料均可通过商购途径获得。本公开的实施例和对比例中所使用的催化裂解催化剂由中国石油化工股份有限公司催化剂齐鲁分公司工业生产，商品牌号为DMMC-2。该催化剂含有平均孔径小于0.7纳米的ZSM-5沸石和超稳Y型沸石，催化剂在使用前在800℃的温度下经饱和蒸汽水热老化17小时，催化剂的主要物化性质见表1。实施例和对比例中所用的含烃原料油为来自江苏油田的原油，其性质列于表2。
表1
催化剂 催化剂
物理性质  
比表面/m 2·g -1 125
孔体积/cm -3·g -1 0.197
表观密度g·cm -3 0.86
化学组成  
Al 2O 3/％ 56.8
SiO 2/％ 42.9
微反活性/％ 68
表2
项目 原油A
密度(20℃)/(g·cm -3) 0.849
凝固点/℃ 35
运动粘度(80℃)/(mm 2/s) 6.8
残炭/％ 3.5
胶质含量/％ 8.4
沥青质含量/％ 0.2
小于250℃馏分质量分数/％ 16.3
小于320℃馏分质量分数/％ 28.6
小于350℃馏分质量分数/％ 34.6
实施例1
本实施例中所加工的原油A轻重馏分油切割点为320℃，切割比例为(轻馏分油/重馏分油的重量比)0.4。
采用改造的连续反应-再生操作的中型装置进行试验，其流程如附图1所示。720℃的高温再生催化剂经再生斜管由再生器引入下行管式反应器1顶部，预热到45℃的轻馏分油经水蒸气雾化后，通过进料喷嘴进入下行管式反应器1与第一股催化剂接触进行催化裂解反应，剂油比为40，反应器的出口温度为665℃，气固停留时间为0.8s，第一催化裂解后的物料经旋风分离出第一反应油气和第一待生催化剂，第一 反应油气进入分离系统，全部第一待生催化剂引入提升管反应器3底部。同时由再生器经由再生催化剂输送管32将温度为720℃的再生催化剂(第二股催化剂)引入提升管反应器3底部。第二股催化剂与第一待生催化剂的重量比(第二股催化剂/第一待生催化剂)为0.25，第一待生催化剂和第二股催化剂在提升管反应器3的底部混合后，混合催化剂在预提升蒸汽作用下向上流动，同时轻烯烃馏分在雾化水蒸气介质下通过轻烯烃馏分进料喷嘴进入提升管反应器3下部，与混合催化剂接触发生反应，重馏分油喷嘴在轻烯烃馏分进料喷嘴的上方800毫米处，重馏分油经水蒸气雾化后，经重馏分油进料喷嘴喷入提升管反应发生催化裂解反应，剂油比为20，反应器的出口温度T3为610℃，反应器中的气固停留时间为1.5s，催化裂解后的物料引入沉降器进行油剂分离，分离为第三反应油气和第三待生催化剂，第三反应油气引入分离系统。第一反应油气、第三反应油气在分离系统中分离成裂化气、轻汽油、重汽油、柴油、油浆。将轻汽油部分馏分(馏程范围30-60℃)作为轻烯烃馏分通过轻烯烃馏分进料喷嘴返回提升管反应器3。第三待生催化剂进入汽提器，汽提出第三待生催化剂上吸附的烃类产物后，通过待生剂斜管进入再生器，与空气接触在720℃下烧焦再生。再生后的催化剂经再生斜管返回反应器中循环使用。中型装置采用电加热维持反应-再生系统温度。待该装置运行稳定(产物组成基本保持不变)后，对从反应油气中得到的裂化气和汽油组成进行分析，得到产物中低碳烯烃(以下简称为三烯)和轻芳烃(以下简称为BTX)产率。
主要操作条件和结果列于表3。
实施例2
采用与实施例1相同的装置和反应步骤，所不同的是所加工原油A的轻重组分切割点为250℃，切割比例为(轻馏分油/重馏分油的重量比)0.195，另外，第二股催化剂与第一待生催化剂的重量比(第二股催化剂/第一待生催化剂)为0.03，再生器的温度为700℃(即、第二股催化剂进入步骤S4时的温度T0为700℃)，提升管反应器的出口温度T3为570℃。
其余主要操作条件和结果列于表3。
实施例3
采用与实施例1相同的装置和反应步骤，所不同的是所加工原油A的轻重组分切割点为350℃，切割比例为(轻馏分油/重馏分油的重量比)0.529，另外，第二股催化剂与第一待生催化剂的重量比(第二股催化剂/第一待生催化剂)为0.6，再生器的温度为740℃(即、第二股催化剂进入步骤S4时的温度T0为740℃)，提升管反应器的出口温度T3为630℃。
其余主要操作条件和结果列于表3。
实施例4
采用与实施例1相同的装置和反应步骤。
本实施例中所加工的原油A轻重馏分油切割点为250℃，切割比例为(轻馏分油/重馏分油的重量比)0.195。
除表3所列的条件之外，采用与实施例1相同的条件。
结果列于表3。
实施例5
采用与实施例1相同的装置和反应步骤。
本实施例中所加工的原油A轻重馏分油切割点为350℃，切割比例为(轻馏分油/重馏分油的重量比)0.529，
除表3所列的条件之外，采用与实施例1相同的条件。
结果列于表3。
实施例6
所加工的原油A轻重馏分油切割点为320℃，切割比例(轻馏分油/重馏分油的重量比)为0.4。
采用改造的连续反应-再生操作的中型装置进行试验，其流程如附图2所示。720℃的高温再生催化剂经再生斜管由再生器引入下行管式反应器1顶部，预热到45℃的轻馏分油经水蒸气雾化后，通过进料喷嘴进入下行管式反应器1，与第一股催化剂接触进行催化裂解反应，剂油比为40，反应器的出口温度为670℃，气固停留时间为0.6s，第一催化裂解后的物料通过出口分布器进入流化床反应器2，进一步进行催化 裂解反应，反应温度为655℃，质量空速为4h -1；另外，轻烯烃馏分经水蒸气雾化后通过进料喷嘴21进入流化床反应器2底部，与热的催化剂接触发生反应，第二催化裂解后的物料进行旋风分离得到第二反应油气和第二待生催化剂，第二反应油气引入后续分离系统，分离的第二待生催化剂经汽提后全部引入提升管反应器3底部。同时由再生器经由再生催化剂输送管32将温度为720℃的再生催化剂(第二股催化剂)引入第二上行管反应器3底部。第二股催化剂与第二待生催化剂的重量比(第二股催化剂/第二待生催化剂)为0.25，第二待生催化剂和第二股催化剂在反应器3的底部混合后，混合催化剂在预提升蒸汽作用下向上流动，重馏分油经水蒸气雾化后，经重馏分油喷嘴喷入提升管反应器3，与催化剂接触发生催化裂解反应，剂油比为20，反应器的出口温度T3为610℃，反应器中的气固停留时间为1.5s，催化裂解后的物料引入沉降器进行油剂分离，分离为第三反应油气和第三待生催化剂，反应油气引入分离系统。第二反应油气、第三反应油气在分离系统中分离成裂化气、轻汽油、重汽油、柴油、油浆。将轻汽油部分馏分(馏程范围30-60℃)作为轻烯烃馏分返回流化床反应器2。第三待生催化剂进入汽提器，汽提出第三待生催化剂吸附的烃类产物后，通过待生剂斜管进入再生器，与空气接触在720℃下烧焦再生。再生后的催化剂经再生斜管返回反应器中循环使用。中型装置采用电加热维持反应、再生系统温度。待该装置运行稳定(产物组成基本保持不变)后，对从反应油气中得到的裂化气和汽油组成进行分析，得到产物中三烯和BTX产率。
主要操作条件和结果列于表3。
实施例7
采用与实施例6相同的装置和反应步骤，所不同的是所加工原油A的轻重组分切割点为250℃，切割比例为(轻馏分油/重馏分油的重量比)0.195，另外，第二股催化剂与第二待生催化剂的重量比(第二股催化剂/第二待生催化剂)为0.03，再生器的温度为700℃(即、第二股催化剂进入步骤S4时的温度T0为700℃)，提升管反应器的出口温度T3为570℃。
其余主要操作条件和结果列于表3。
实施例8
采用与实施例6相同的装置和反应步骤，所不同的是所加工原油A的轻重组分切割点为350℃，切割比例为(轻馏分油/重馏分油的重量比)0.529，另外，第二股催化剂与第二待生催化剂的重量比(第二股催化剂/第二待生催化剂)为0.6，再生器的温度为740℃(即、第二股催化剂进入步骤S4时的温度T0为740℃)，提升管反应器的出口温度T3为630℃。
其余主要操作条件和结果列于表3。
对比例1
采用与实施例1相同的装置和反应步骤、反应条件，所不同的是由反应油气中分离得到的轻烯烃馏分不回流到提升管反应器3。
其余主要操作条件和结果列于表3。
对比例2
采用与实施例1相同的装置和反应步骤，所不同的是第二股催化剂与第一待生催化剂的重量比(第二股催化剂/第一待生催化剂)为0.1，再生器的温度为740℃(即、第二股催化剂进入步骤S4时的温度T0为740℃)，提升管反应器的出口温度T3为610℃。
其余主要操作条件和结果列于表3。
对比例3
采用与实施例1相同的装置和反应步骤，所不同的是第二股催化剂与第一待生催化剂的重量比(第二股催化剂/第一待生催化剂)为0.5，再生器的温度为700℃(即、第二股催化剂进入步骤S4时的温度T0为700℃)，提升管反应器的出口温度T3为610℃。
其余主要操作条件和结果列于表3。
对比例4
采用与实施例1相同的装置和反应步骤、反应条件，所不同的是在提升管反应器3下部不引入第二股催化剂，而只使用第一待生催化 剂。
其余主要操作条件和结果列于表3。
对比例5
采用与实施例1相同的装置和反应步骤、反应条件，所不同的是在提升管反应器3下部不引入第一待生催化剂，只使用第二股催化剂。
其余主要操作条件和结果列于表3。
对比例6
采用与实施例1相同的装置和反应步骤、反应条件，所不同的是轻馏分油进入提升管反应器，即第一下行管反应器1变为提升管反应器(表3中的下行管式反应器的各参数在该对比例6中表示上行管式反应器的各参数)。
其余主要操作条件和结果列于表3。
Figure PCTCN2021127339-appb-000001
Figure PCTCN2021127339-appb-000002
由表3的数据可以看出，使用本公开提供的含烃原料油催化裂解生产低碳烯烃和BTX的方法，可以明显提高低碳烯烃和轻芳烃的产率，同时，干气和焦炭等副产物的产率得到抑制。
以上详细描述了本公开的优选实施方式，但是，本公开并不限于上述实施方式中的具体细节，在本公开的技术构思范围内，可以对本公开的技术方案进行多种简单变型，这些简单变型均属于本公开的保护范围。
另外需要说明的是，在上述具体实施方式中所描述的各个具体技术特征，在不矛盾的情况下，可以通过任何合适的方式进行组合。为了避免不必要的重复，本公开对各种可能的组合方式不再另行说明。
此外，本公开的各种不同的实施方式之间也可以进行任意组合，只要其不违背本公开的思想，其同样应当视为本公开所公开的内容。

Claims (17)

  1. 含烃原料油催化裂解生产低碳烯烃和轻芳烃的方法，该方法包括如下步骤：
    S1、将含烃原料油切割为轻馏分油和重馏分油，所述轻馏分油相对于所述重馏分油的重量比(轻馏分油/重馏分油)为X；
    S2、将所述轻馏分油与第一股催化剂引入第一下行式反应器，进行第一催化裂解，得到第一催化裂解后的物料；
    任选的S2′、将所述第一催化裂解后的物料引入流化床反应器进行第二催化裂解，得到第二催化裂解后的物料；
    S3、将所述第一催化裂解后的物料进行气固分离，得到第一反应油气和第一待生催化剂，或者将所述第二催化裂解后的物料进行气固分离，得到第二反应油气和第二待生催化剂；
    S4、将连续催化剂、所述重馏分油与第二股催化剂引入第二上行式反应器，进行第三催化裂解，然后进行气固分离，得到第三反应油气和第三待生催化剂；所述连续催化剂为所述第一待生催化剂的至少一部分或所述第二待生催化剂的至少一部分；所述第二股催化剂与所述连续催化剂的重量比(第二股催化剂/连续催化剂)为R；
    S5、从所述第一反应油气、所述第二反应油气和所述第三反应油气的任一者或者第一反应油气和第三反应油气的混合物或者第二反应油气和第三反应油气的混合物中分离出低碳烯烃和轻芳烃，并且分离出轻烯烃馏分，并将所述轻烯烃馏分返回步骤S4的所述第二上行式反应器中或者步骤S2′的所述流化床反应器中，
    所述R和X满足以下关系式：
    (4.84×T0-3340)/(780+5×T0-6×T3)＜R/X＜(0.968×T0-630)/(668+0.2×T0-1.2×T3)
    T0为所述第二股催化剂进入步骤S4时的温度(单位℃)，T3为所述第二上行式反应器的出口温度(单位℃)。
  2. 根据权利要求1所述的方法，其中，所述第二上行式反应器的出口温度T3为530-650℃，优选为560-640℃，进一步优选为580-630℃，更进一步优选为600-630℃；和/或
    所述第二股催化剂进入步骤S4时的温度T0为690-750℃，优选 700-740℃，进一步优选为705-730℃，更进一步优选为710-725℃。
  3. 根据权利要求1或2所述的方法，其中，步骤S1中，在切割点100-400℃之间的任意温度将含烃原料油切割为轻馏分油和重馏分油，使所述轻馏分油相对于所述重馏分油的重量比(轻馏分油/重馏分油)为X。
  4. 根据权利要求1-3中任一项所述的方法，其中，
    所述第一下行式反应器中，所述第一催化裂解的条件包括：所述第一下行式反应器的出口温度为610-720℃，气固停留时间为0.1-3.0秒，剂油比为15-80；和/或
    所述流化床反应器中，所述第二催化裂解的条件包括：所述流化床反应器中的反应温度为600-690℃，质量空速为2-20h -1；和/或
    所述第二上行式反应器中，所述第三催化裂解的条件包括：气固停留时间为0.5-8秒，剂油比为8-40。
  5. 根据权利要求1-4中任一项所述的方法，其中，
    所述第一下行式反应器中，所述第一催化裂解的条件包括：所述第一下行式反应器的出口温度为650-690℃，气固停留时间为0.5-1.5秒，剂油比为25-65；和/或
    所述流化床反应器中，所述第二催化裂解的条件包括：所述流化床反应器中的反应温度为640-670℃，质量空速为4-12h -1；和/或
    所述第二上行式反应器中，所述第三催化裂解的条件包括：气固停留时间为1.5-5秒，剂油比为10-30。
  6. 根据权利要求1-5中任一项所述的方法，其中，
    存在步骤S2′时，在步骤S3的气固分离中，将分离的催化剂进行汽提，得到第二待生催化剂；和/或
    在步骤S4中，首先使所述连续催化剂与所述第二股催化剂混合，再进行后续催化裂解反应；和/或
    在步骤S4中，将来自S5步骤的所述轻烯烃馏分先于所述重馏分油与所述第二股催化剂和连续催化剂的混合物接触；优选将所述轻烯烃馏分先于所述重馏分油0.3-1.0秒与所述第二股催化剂和连续催化剂的混合物接触，更优选将所述轻烯烃馏分先于所述重馏分油0.4-0.8秒与所述第二股催化剂和连续催化剂的混合物接触；和/或
    所述方法在步骤S1之前具有步骤S0，其中，将含烃原料油进行脱 盐脱水处理，将得到的经脱水脱盐的含烃原料油引入步骤S1进行切割。
  7. 根据权利要求1-6中任一项所述的方法，其中，该方法还包括：
    在步骤S4的气固分离中，将分离的催化剂进行汽提，得到第三待生催化剂；和/或
    将所述第三待生催化剂和任选的未进入第二上行式反应器的第一待生催化剂或第二待生催化剂，在690-750℃、优选700-740℃、进一步优选705-730℃、更进一步优选710-725℃的温度下进行烧焦再生，得到再生催化剂；和/或
    将所述第一反应油气、所述第二反应油气和所述第三反应油气的任一者或者第一反应油气和第三反应油气的混合物或者第二反应油气和第三反应油气的混合物进行分离，得到干气、C3馏分、C4馏分、轻汽油、重汽油、柴油和油浆，从中分离得到低碳烯烃、轻芳烃，并分离出轻烯烃馏分；和/或
    不存在步骤S2′时，在步骤S5中，从第一反应油气、第三反应油气的任一者或者两者的混合物中分离出轻烯烃馏分，并将所述轻烯烃馏分返回步骤S4的第二上行式反应器中；存在步骤S2′时，在步骤S5中，从第二反应油气、第三反应油气的任一者或者两者的混合物中分离出轻烯烃馏分，并将所述轻烯烃馏分返回步骤S2′的流化床反应器中。
  8. 根据权利要求1-7中任一项所述的方法，其中，所述含烃原料油为原油、煤液化油、合成油、油砂油、页岩油、致密油和动植物油脂中的一种或其两种以上的混合物，，或其各自的部分馏分、其各自的重质馏分的加氢改质油。
  9. 根据权利要求1-8中任一项所述的方法，其中，所述第一股催化剂和所述第二股催化剂各自独立地包括活性组分和载体，所述活性组分为选自含或不含稀土的超稳Y型沸石、ZSM-5系列沸石、具有五元环结构的高硅沸石和β沸石中的至少一种，所述载体为选自氧化铝、氧化硅、无定形硅铝、氧化锆、氧化钛、氧化硼和碱土金属氧化物中的至少一种。
  10. 根据权利要求1-9中任一项所述的方法，其中，所述第一股催 化剂和所述第二股催化剂各自独立地包括再生催化剂，优选所述第一股催化剂和所述第二股催化剂为再生催化剂，和/或
    将所述第一待生催化剂的全部或所述第二待生催化剂的全部作为连续催化剂。
  11. 一种含烃原料油催化裂解生产低碳烯烃和轻芳烃的装置，该装置包括以下单元：
    含烃原料油切割单元，在其中将含烃原料油切割为轻馏分油和重馏分油，使得所述轻馏分油相对于所述重馏分油的重量比(轻馏分油/重馏分油)为X，
    第一下行式反应单元，从该反应单元的上方引入所述轻馏分油与第一股催化剂，进行第一催化裂解，在该反应单元的下方得到第一催化裂解后的物料；
    任选的流化床反应单元，其中，引入所述第一催化裂解后的物料，进行第二催化裂解，得到第二催化裂解后的物料；
    第一气固分离单元，其中引入所述第一催化裂解后的物料进行气固分离，得到第一反应油气和第一待生催化剂，或者其中引入所述第二催化裂解后的物料进行气固分离，得到第二反应油气和第二待生催化剂；
    第二上行式反应单元，从该反应单元的下方引入连续催化剂、第二股催化剂与所述重馏分油，进行第三催化裂解，在该反应单元的上方得到第三催化裂解后的物料，所述连续催化剂为所述第一待生催化剂的至少一部分或所述第二待生催化剂的至少一部分，所述第二股催化剂与所述连续催化剂的重量比(第二股催化剂/连续催化剂)为R，
    第二气固分离单元，其中引入所述第三催化裂解后的物料进行气固分离，得到第三反应油气和第三待生催化剂；
    分离单元，其中引入所述第一反应油气、所述第二反应油气和所述第三反应油气的任一者或者第一反应油气和第三反应油气的混合物或者第二反应油气和第三反应油气的混合物，分离出低碳烯烃和轻芳烃，并且分离出轻烯烃馏分，并将轻烯烃馏分返回至所述第二上行式反应单元或者所述流化床反应单元；
    其中，所述R和X满足以下关系式：
    (4.84×T0-3340)/(780+5×T0-6×T3)＜R/X＜(0.968×T0-630)/(668+0.2×T 0-1.2×T3)
    T0为所述第二股催化剂进入第二上行式反应单元时的温度(单位℃)，T3为所述第二上行式反应单元的出口温度(单位℃)。
  12. 根据权利要求11所述的装置，其中，还包括再生单元，其中，引入所述第三待生催化剂和任选的未进入第二上行式反应器的第一待生催化剂或第二待生催化剂，在690-750℃、优选700-740℃、进一步优选705-730℃，更进一步优选710-725℃的温度下进行烧焦再生，得到再生催化剂。
  13. 根据权利要求11或12所述的装置，其中，所述装置在包括流化床反应单元时，第一气固分离单元中还包括汽提单元，其中，将气固分离得到的催化剂进行汽提，得到第二待生催化剂。
  14. 根据权利要求11-13中任一项所述的装置，其中，所述第二气固分离单元中还包括汽提单元，其中，将气固分离得到的催化剂进行汽提，得到第三待生催化剂。
  15. 根据权利要求11-14中任一项所述的装置，其中，所述装置还包括脱水脱盐单元，其中，将含烃原料油进行脱盐脱水处理，将得到的经脱水脱盐的含烃原料油引入含烃原料油切割单元进行切割。
  16. 根据权利要求11-15中任一项所述的装置，其中，第二上行式反应单元中，引入连续催化剂和第二股催化剂的位置在轻烯烃馏分的进料口的上游。
  17. 根据权利要求11-15中任一项所述的装置，其中，第二上行式反应单元中，来自分离单元的轻烯烃馏分的进料口在重馏分油进料口的上游。
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