WO2019228130A1 - 一种石油烃多级流化催化反应方法及反应器 - Google Patents

一种石油烃多级流化催化反应方法及反应器 Download PDF

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WO2019228130A1
WO2019228130A1 PCT/CN2019/085278 CN2019085278W WO2019228130A1 WO 2019228130 A1 WO2019228130 A1 WO 2019228130A1 CN 2019085278 W CN2019085278 W CN 2019085278W WO 2019228130 A1 WO2019228130 A1 WO 2019228130A1
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reaction
reaction zone
catalyst
zone
stage
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PCT/CN2019/085278
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English (en)
French (fr)
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石宝珍
石林
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青岛京润石化设计研究院有限公司
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Priority to US17/059,173 priority Critical patent/US11578274B2/en
Publication of WO2019228130A1 publication Critical patent/WO2019228130A1/zh

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    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G51/00—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes only
    • C10G51/02—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes only plural serial stages only
    • C10G51/026—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes only plural serial stages only only catalytic cracking steps
    • B—PERFORMING OPERATIONS; TRANSPORTING
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    • B01J8/00—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes
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    • B01J8/1827—Feeding of the fluidising gas the fluidising gas being a reactant
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    • B01J8/24—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles according to "fluidised-bed" technique
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    • B01J8/28—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles according to "fluidised-bed" technique with two or more fluidised beds, e.g. reactor and regeneration installations the one above the other
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    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G11/00—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
    • C10G11/14—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts
    • C10G11/18—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts according to the "fluidised-bed" technique
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    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
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    • C10G11/00—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
    • C10G11/14—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts
    • C10G11/18—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts according to the "fluidised-bed" technique
    • C10G11/182—Regeneration
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    • C10G11/00—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
    • C10G11/14—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts
    • C10G11/18—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts according to the "fluidised-bed" technique
    • C10G11/187—Controlling or regulating
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    • C10G51/00—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes 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
    • C10G51/00—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes only
    • C10G51/02—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes only plural serial stages only
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J2208/00—Processes carried out in the presence of solid particles; Reactors therefor
    • B01J2208/00743—Feeding or discharging of solids
    • B01J2208/00769—Details of feeding or discharging
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J2208/00—Processes carried out in the presence of solid particles; Reactors therefor
    • B01J2208/00796—Details of the reactor or of the particulate material
    • B01J2208/00938—Flow distribution elements
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    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/20—Characteristics of the feedstock or the products
    • C10G2300/201—Impurities
    • C10G2300/202—Heteroatoms content, i.e. S, N, O, P
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
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    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/20—Characteristics of the feedstock or the products
    • C10G2300/30—Physical properties of feedstocks or products
    • C10G2300/308—Gravity, density, e.g. API
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    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
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    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/40—Characteristics of the process deviating from typical ways of processing
    • C10G2300/4006—Temperature
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/40—Characteristics of the process deviating from typical ways of processing
    • C10G2300/4018—Spatial velocity, e.g. LHSV, WHSV
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • 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
    • C10G2300/4093—Catalyst stripping

Definitions

  • the invention belongs to the technical field of catalytic conversion of petroleum hydrocarbons, and particularly relates to a multistage fluidized catalytic reaction method and reactor for petroleum hydrocarbons, capable of producing propylene in a large amount.
  • the catalytic cracking unit is the main gasoline production unit and also an important propylene production unit.
  • Propylene is one of the most important petrochemical raw materials. Drawing on the operation and design experience of conventional heavy oil catalytic cracking reaction-regeneration systems, researchers at home and abroad have developed a series of technologies for catalytic cracking of heavy oil to produce propylene.
  • TMP technology based on two-stage riser catalytic cracking technology.
  • This technology uses heavy oil as raw material, and utilizes the process characteristics of two-stage riser catalytic cracking process, catalytic relay, and large agent-to-oil ratio to optimize the combination of feeding methods for different types of reaction materials to control the suitability of different materials. Reaction conditions to achieve the purpose of improving the yield of propylene.
  • Most companies producing propylene in foreign countries also adopt the double riser method.
  • a reactor is divided into two cracking reaction zones connected in series, that is, the riser and the fluidized bed are subjected to a catalytic cracking reaction, and the riser is partially subjected to a raw oil cracking reaction and a fluidized bed.
  • the intermediate product of the riser reaction zone is further subjected to a secondary cracking reaction for increasing the production of propylene; since the catalyst in the fluidized bed reaction zone is a standby agent for the reaction of the raw material oil, the small molecule catalytic cracking in the fluidized bed reaction zone The performance is reduced, and the reaction temperature needs to be increased to increase the thermal cracking to increase the yield of propylene and further reduce the selectivity of propylene.
  • Both DCC and DCC-PLUS use a regenerant cycle to control the reaction zone of the fluidized bed bed with a fixed amount of raw materials.
  • the space velocity can only be controlled by the change of the catalyst level in the fluidized bed.
  • the fluidized bed bed area must be equipped with a catalyst outlet to achieve the level control through the catalyst outflow; due to the requirements of catalyst carrying and gas-solid separation, the fluidized bed There is a dilute phase space between the reaction zone of the bed and the gas-solid separator, and the oil and gas still carry a large amount of catalyst when leaving the fluidized bed.
  • the residence time from the open fluidized bed to the gas-solid separator is more than 20 seconds.
  • the catalyst above the catalyst level and the residence time of the oil and gas will inevitably cause further side reactions to stop, and the propylene will be further thermally cracked. Influencing product distribution and propylene selectivity, the reaction is difficult to terminate in time.
  • the catalysts in the DCC and DCC-PLUS fluidized bed cracking reaction zone are carbon deposition catalysts after the reaction of the feedstock oil. In order to increase the propylene yield, the fluidized bed reaction needs to be further improved. The reaction temperature in the zone will inevitably lead to the limitation of the catalytic cracking reaction, the increase of the thermal reaction, and the selectivity of propylene will be greatly reduced, leading to higher dry gas and coke yields.
  • the prior art focuses on the production of propylene and is divided into two types.
  • the first type is a riser and fluidized bed series reaction
  • the second type is a double riser parallel reaction.
  • the prior art mainly focuses on changing the catalyst and reaction conditions, such as using a higher reaction temperature, agent-to-oil ratio, and steam injection volume than the FCC process to improve cracking reaction depth and propylene selectivity.
  • An object of the present invention is to provide a multi-stage fluidized catalytic reaction method for petroleum hydrocarbons, which can realize the classification reaction of different properties of raw materials according to their respective conditions, and simultaneously realize the control of the classification of the catalyst and the reaction temperature during the reaction process, and improve the selection of propylene. Properties, significantly reducing the yield of by-products such as dry gas and coke.
  • Another object of the present invention is to provide a multistage fluidized catalytic reactor for petroleum hydrocarbons.
  • a multi-stage fluidized catalytic reaction method for petroleum hydrocarbons realizes zoned multi-stage reactions in one reactor, including a first-stage reaction of a feed oil, a second-stage reaction, a light hydrocarbon and / or a cycle oil cracking reaction process; the method
  • the catalyst replacement and two-stage relay reactions are realized during the first and second stage reactions of the medium catalytic cracking feedstock, and the two-stage reactions of light hydrocarbons and / or circulating oils are also realized; these reactions are realized in different areas within a reactor; light hydrocarbons and The first stage reaction of the circulating oil is performed in a separate reaction zone;
  • Catalytic cracking feed oil first enters the first reaction zone (also known as the feed oil reaction zone or the feed oil cracking reaction zone) to perform the first-stage catalytic cracking reaction under the environment of a catalyst (regenerated catalyst), and then the catalyst is separated at the exit of the first reaction zone.
  • the oil and gas enters the second-stage catalytic cracking reaction of the second reaction zone to continue deep cracking through the first-stage reaction zone outlet logistics conveying pipe provided inside the third reaction zone.
  • the first catalyst enters the first reaction zone below the feedstock inlet. Under the action of the pre-lifting medium, it enters the first reaction zone; the stream in the second reaction zone enters the settler from the conveying pipe for gas-solid separation;
  • the first reaction zone adopts the riser reaction conditions, and a catalyst splitter (swirl gas-solid splitter) is set at the outlet;
  • the second catalyst enters the third reaction zone in the third reaction zone (the third reaction zone is arranged above the first reaction zone, and the third reaction zone is a light hydrocarbon, circulating oil reaction zone), and the light hydrocarbon and / or circulating oil enters the third
  • the reaction zone performs an independent first-stage catalytic cracking reaction; then, the catalyst and reaction products in the third reaction zone enter the second reaction zone upwards to continue the second-stage catalytic cracking reaction (that is, the third reaction zone reaction).
  • the latter catalyst is continued to be used in the second reaction zone).
  • the third reaction zone adopts turbulent fluidized bed or circulating fluidized bed reaction conditions; in specific implementation, the second catalyst enters the third reaction zone and performs a catalytic reaction with the reaction raw materials entering the reaction zone.
  • the third reaction zone is arranged around the first reaction outlet stream conveying pipe above the first reaction zone, and the second catalyst entering the third reaction zone from outside the reactor is fluidized with steam in this zone to form a turbulent fluidized bed or cycle. Fluidized bed conditions; the reaction products of the first reaction zone and part of the catalyst react with the catalyst and gas after the reaction in the third reaction zone and enter the second reaction zone above, and perform the second-stage catalytic reaction in the second reaction zone to achieve the Cracking transformation
  • reaction products and catalysts in the third reaction zone and the oil and gas in the first reaction zone enter the second reaction zone to continue the reaction.
  • the first catalyst enters the pre-lifting section (or catalyst conveying section) of the first reaction section from the first catalyst inlet pipe and is fluidized and conveyed to the first reaction section by the pre-lifting medium.
  • the first reaction zone of a reaction section it is mixed with the first catalyst to gasify and perform a first-stage catalytic cracking reaction (specifically, the reaction is completed in a riser reactor, and the raw material oil is atomized into small droplets through a nozzle by steam).
  • the feedstock oil After contacting the catalyst in the riser reaction zone and obtaining heat, the feedstock oil is subjected to a catalytic cracking reaction.
  • the first reaction zone mainly realizes the conversion of the feedstock oil to olefin-rich gasoline and diesel components, although it is not necessary.
  • the reaction time in the reaction zone is controlled within 1.5 seconds, not more than 2.5 seconds, and the reaction temperature is controlled below 540 ° C.)
  • the catalyst split rate in the first reaction zone is 70% to 90%, and 10% to 30% of the catalyst remains in the reaction oil and gas and reacts with the reaction oil.
  • the catalyst separated from the first reaction exit stream settles into the first stripping section below and is stripped out of the reactor from the bioreactor transfer pipe; the oil and gas is transported through the first reaction zone exit stream
  • the tube enters the second reaction zone above to perform the second-stage catalytic cracking reaction (in specific implementation, the second-stage catalytic cracking reaction can further increase the steam ratio and reduce the partial pressure of petroleum hydrocarbon components);
  • the second catalyst enters the tube from the second catalyst into the third reaction zone provided in the third reaction zone between the first stripping section and the second reaction zone, and the light hydrocarbons and / or circulating oil enter the third reaction zone.
  • an independent first-stage catalytic cracking reaction is performed; then, the catalyst and reaction products in the third reaction zone enter the second reaction zone and continue to perform the second-stage (re) catalytic cracking reaction.
  • the light hydrocarbon may be C4 component or petroleum hydrocarbon whose final boiling point is lower than 210 ° C;
  • the catalyst is fluidized with steam, so that the zone is formed into a turbulent fluidized bed or a circulating fluidized bed state, and these steams enter the second reaction zone at the same time;
  • the gas and catalyst after the reaction in the second reaction zone are sent to the cyclone separator in the settler for gas-solid separation, and the reaction products flow out of the settler;
  • the catalyst after the reaction in the second reaction zone enters the first catalyst stripping zone of the first stripping section directly from the catalyst return pipe in the second reaction zone, and is stripped with the catalyst separated from the first reaction zone through the catalyst splitter. It then flows out of the reactor from the standby agent delivery pipe; or a second stripping section is provided outside the second reaction zone, and the second standby catalyst after the stripping flows out of the reactor from the catalyst outlet pipe.
  • the raw materials when a variety of raw materials are reacted in the third reaction zone, the raw materials are layered into the third reaction zone for reaction according to the boiling point or final boiling point, and the components with lower boiling points or final boiling points enter the third layer in the lower layer.
  • components with high boiling points or final boiling points enter the third reaction zone in the upper layer in turn.
  • the light cycle oil enters the third reaction zone above the light hydrocarbons.
  • the first reaction zone adopts a riser, that is, a conveying bed reaction form
  • the third reaction zone adopts a turbulent fluidized bed or circulating fluidized bed reaction form
  • the second reaction zone adopts Circulating fluidized bed or turbulent fluidized bed reaction form.
  • the apparent flow velocity of the turbulent fluidized bed gas is less than 1.2 m / s
  • the apparent flow velocity of the circulating fluidized bed gas is 1.2-4.0 m / s.
  • the third reaction section is vertically divided into a plurality of isolated areas, that is, a partition area and / or a cylinder, through a vertical partition and / or a vertical cylinder.
  • the body region (specifically, the partition plate and the cylinder can be connected or supported on the shell of the third reaction section); the catalyst entering the third reaction section passes through the communication channels at the bottom of each partition region and / or the barrel region, respectively Enter each zone; the partitioned zone and / or barrel zone divided in whole or in part is used as the third reaction zone; light hydrocarbons and / or circulating oil are in one or more of the above-mentioned partitioned zone and / or barrel zone Partial or full zone reaction; when multiple raw materials react in the third reaction zone, different raw materials react in different partition zones or barrel zones according to their respective conditions; especially when the C4 component enters the third reaction zone independently In this case, the reaction is preferably performed in a separate partition region or a cylinder region.
  • the amount of reactants entering the third reaction zone is based on the carbon content of the catalyst after the reaction in the second reaction zone, which is the second reaction zone.
  • the catalyst activity is controlled such that the weight ratio of the carbon content of the catalyst after the reaction in the second reaction zone is not more than 0.8%.
  • the amount of reactants entering the third reaction zone is controlled according to the carbon content of the catalyst after the reaction in the third reaction zone, so that the weight ratio of the carbon content of the catalyst after the reaction in the third reaction zone is not greater than 0.4 %.
  • a catalyst heat pipe is installed in the third reaction zone, and the catalyst in the second reaction zone is used to realize the catalyst in the second reaction zone.
  • the catalyst heat extraction tube can be located in the isolated area as the third reaction area (that is, the heat extraction tube is also located in the isolated reaction area. This reaction area Simultaneously take heat and reaction), or it is located in a separate area that is not the third reaction zone. Heat extraction from catalysts in a fluidized bed is a common technique and will not be described in detail.
  • the second reaction zone adopts a reaction temperature higher than or the same as the first reaction zone; preferably, the reaction temperature in the second reaction zone is higher than the reaction temperature in the first reaction zone. ° C to 70 ° C, more preferably 10 ° C to 70 ° C; the reaction temperature in the second reaction zone is controlled by the amount of the second catalyst.
  • a fourth reaction zone is provided between the second reaction zone and the outer shell of the second reaction zone, and the zone is used for circulating oil or refining reaction.
  • the invention also provides a petroleum hydrocarbon multi-stage fluidized catalytic reactor:
  • the reactor includes a first reaction section, a catalyst splitter, a third reaction section, a second reaction section and a settler from bottom to top,
  • the first reaction section is a riser reactor, which includes a first catalyst inlet tube, a pre-lift section, a feedstock inlet, and a first reaction zone from bottom to top for a feedstock oil reaction; the feedstock reaction zone is the first reaction zone A first stripping section is set around, and the catalyst splitter is disposed at the outlet of the first reaction zone and is located at the top of the first catalyst stripping zone in the first stripping section;
  • a baffle is provided between the first catalyst stripping zone and the third reaction zone, and the baffle is provided with a first-reaction-zone outlet logistics conveying pipe; a second reaction zone is provided with a second reaction zone, and a third reaction zone is provided with There is a third reaction zone; the outlet of the first reaction zone and the delivery pipe exit is located at the top of the third reaction zone and the entrance of the second reaction zone (this first reaction zone exit stream delivery pipe is located inside the third reaction zone); the third reaction The section is provided with a light hydrocarbon or circulating oil inlet and a second catalyst inlet pipe for light hydrocarbon and circulating oil reaction;
  • a second partition is provided between the third reaction zone and the outside of the second reaction zone in the second reaction zone, and a transfer pipe is provided between the second reaction zone and the gas-solid separator in the settler.
  • the first reaction section, the catalyst splitter, the first stripping section, the third reaction section, the second reaction section and the settler are arranged coaxially above and below.
  • an orifice plate is further provided between the third reaction zone and the second reaction zone, and the stream from the first reaction zone and the catalyst and gas from the third reaction zone pass through the hole.
  • the plates are distributed to the second reaction zone.
  • a second reaction zone catalyst return pipe is provided between the second reaction zone and the first stripping zone, so that the catalyst after the reaction in the second reaction zone is directly refluxed.
  • a second reaction zone catalyst return pipe is provided between the second reaction section and the first stripping section, and is outside the second reaction zone (that is, the outer shell of the second reaction zone and the second reaction zone) (2)
  • a second stripping section is provided, so that the catalyst after the reaction in the second reaction zone is stripped in the second stripping section from the catalyst return pipe in the second reaction zone into the first stripping section, and then transported from the standby agent.
  • the tube flows out of the reactor.
  • a catalyst storage zone after the reaction in the second reaction zone.
  • the catalyst separated by the gas-solid separator in the settler first enters the storage zone, and then enters the first stripping section and is stripped and then flows out.
  • a reactor, or a stripping member is provided in the storage area, and the storage area can be used as a second stripping section.
  • the second stripping section is provided with a catalyst outflow pipe (outlet), and the catalyst separated by the gas-solid separator is directly located there. After the stripping in the storage area, it exits the reactor, or it is stripped in the second stripping section before entering the first stripping section.
  • the above-mentioned petroleum hydrocarbon multi-stage fluidized catalytic reactor further includes, if necessary, a fourth reaction zone around the second reaction zone (between the second reaction zone and the outer shell of the second reaction zone) for circulating oil or Independent reaction of refining oil.
  • a fourth reaction zone around the second reaction zone (between the second reaction zone and the outer shell of the second reaction zone) for circulating oil or Independent reaction of refining oil.
  • the fourth reaction zone is located above the second stripping section.
  • the heat and residual activity of the catalyst in the second reaction zone are used in the fourth reaction zone for catalytic conversion of heavy-cycle oil or refining oil.
  • the reaction zone adopts bubbling fluidized bed or turbulent fluidized bed conditions; when C4 is independent During refining, C4 can react in the pre-lift section below the first reaction zone.
  • the space between the third reaction stage shell and the first reaction zone outlet logistics conveying pipe is divided into a plurality of parts by vertical partitions and / or vertical cylinders.
  • the isolated area is the partition area and / or the cylinder area (the partition and the barrel can be connected or supported on the shell of the third reaction section), the vertical partition or the bottom end of the vertical cylinder communicate with each other,
  • the catalyst flows in various sections; light hydrocarbons and steam inlets are provided at the bottom inlets of some or all of the partition zone and / or the barrel zone, so that different raw materials can react in different partition zones or barrel zones according to their respective conditions.
  • steam inlets are provided at the bottom inlets of all the partition areas and the barrel areas, and light hydrocarbon inlets are provided only in the areas where the reaction is to be isolated to introduce light hydrocarbon feedstock.
  • a catalyst reflux tube is arranged between the catalysts to return the catalyst from the third reaction zone or the second stripping section or the first catalyst stripping section to the pre-lifting section;
  • the components of the reactor are arranged up and down integrally, preferably (usually) coaxially arranged up and down, the first reaction section is located at the bottom, and the third reaction section is located between the first reaction section and the second reaction section;
  • the second stripping section is set at the periphery of the second reaction zone, the settler is at the top, and a two-stage gas-solid separator is set in the settler.
  • the second reaction zone is connected to the inlet of the first-stage gas-solid separator through a transfer pipe, and the second reaction The reaction stream in the zone enters the gas-solid separator directly through the conveying pipe, avoiding the reaction of oil and gas into the dilute phase space of the settler, and achieving rapid gas-solid separation and space velocity termination of the reaction;
  • the first reaction section of the reactor of the present invention is a riser tube, which includes a pre-lifter section, a first reaction zone, a first stripping section, a first catalyst inlet tube, and a regenerant delivery tube;
  • the second reaction section includes a second Reaction zone, (out of reaction) catalyst outflow pipe;
  • the third reaction zone includes a third reaction zone, a first reaction zone exit stream delivery pipe, and a second catalyst entry pipe;
  • the third reaction section can be provided with an independent steam inlet, and the steam enters the second reaction zone from the third reaction zone to supplement the steam required in the second reaction zone;
  • the light hydrocarbon according to the present invention is a C4 component, a catalytic light gasoline, gasoline, or naphtha component, and the final boiling point is not higher than 210 ° C;
  • the circulating oil includes a light circulating oil and a heavy circulating oil component;
  • the light cycle oil is a petroleum hydrocarbon or light diesel oil having a boiling point of 180 ° -255 ° C;
  • the heavy cycle oil is a petroleum hydrocarbon or heavy diesel oil having a boiling point of 250 ° -365 ° C;
  • the refining oil is petroleum having a boiling point of 340 ° -520 ° C Hydrocarbon component.
  • the invention realizes the partition reaction of different raw materials, and simultaneously realizes the catalyst replacement, the stepwise control of the reaction temperature and the control of the carbon content of the bioreagent.
  • the reaction conditions are provided for different raw materials, and the catalyst in the reactor is improved.
  • Efficiency which can significantly reduce the yield of dry gas and coke, increase the yield of propylene, and improve economic benefits; because the reactants in the second reaction zone are the gas phase from the first reaction zone and the third reaction zone, and the catalyst in the second reaction zone Carbon content is critical to the selectivity of the reaction;
  • the method of the invention realizes multi-stage control of reaction raw materials, catalysts and reaction temperature in one reactor.
  • Different raw materials or different reaction stages adopt corresponding fluidized reactor forms; the equipment is simple and does not change the arrangement of conventional reaction settler and Installation conditions greatly reduce investment and operating costs.
  • reaction process of the method of the present invention is:
  • the catalytic fluidized feedstock oil heated to 200-350 °C enters the atomization nozzle in several ways, and the steam with the amount of feedstock oil is about 5% to 5% is used to achieve atomization and enter the lowermost first reaction stage riser reactor.
  • the first catalyst (regenerated catalyst) delivered by the pre-lift gas is mixed, and the catalytic cracking reaction is performed in the first reaction zone for about 1.0 to 2.0 seconds after the gasification.
  • the reaction temperature is about 500 ° C, and the reaction heat is provided by the regenerated catalyst.
  • the reaction oil, gas and steam transport the catalyst in the first reaction zone to the outlet of the first reaction zone, and a portion of the catalyst is separated and flowed out by the cyclone, and the oil and gas directly enters the first reaction zone outlet logistics conveying pipe provided in the third reaction zone;
  • the second catalyst was replenished in the third reaction section.
  • the supplementary catalyst was mixed with the oil and gas from the first reaction zone at the outlet of the first reaction zone and the outlet of the delivery pipe and flowed upward into the second reaction zone.
  • the catalyst was replenished in the second reaction zone.
  • the oil and gas generated in the first reaction zone continues to crack under the new catalyst conditions.
  • the weight ratio of the carbon content of the catalyst in the second reaction zone is controlled to 0.6%, not more than 0.8%. According to different product requirements, the reaction time in the second reaction zone is 1.0 seconds to 6.0 seconds. At the same time, the reaction temperature is increased and the conversion efficiency is improved.
  • FIG. 1 is a schematic structural diagram of a device of Embodiment 1 of a method for multistage fluidized catalytic reaction of petroleum hydrocarbons according to the present invention
  • Figure 2 A-A view in Figure 1, that is, a schematic diagram of the layout of the logistics conveyance pipe at the exit of the first reaction zone;
  • Figure 3 A-A view in Figure 1, that is, a schematic diagram of the second arrangement of the outlet conveying pipe of the first reaction zone;
  • FIG. 4 is a schematic structural diagram of a second embodiment of a method for multistage fluidized catalytic reaction of petroleum hydrocarbons according to the present invention.
  • An orifice plate (distribution plate) is provided at the entrance of the second reaction zone;
  • FIG. 5 is a schematic structural diagram of a third embodiment of a method for implementing a multistage fluidized catalytic reaction of petroleum hydrocarbons according to the present invention, and a fourth reaction zone is provided;
  • FIG. 6 is a schematic structural diagram of a fourth embodiment of a multi-stage fluidized catalytic reaction method for petroleum hydrocarbons according to the present invention.
  • the catalyst after the reaction in the second reaction zone returns to the first stripping section;
  • Figure 7 A partially enlarged view of the third reaction section in Figure 6;
  • FIG. 8 is a schematic cross-sectional view of the third reaction section divided into a plurality of isolated regions, viewed from the direction of B-B in FIG. 7.
  • first reaction section 112 pre-lift section, 113 feedstock inlet (nozzle), 114 catalyst splitter, 115 first reaction zone, 116 baffle, 118 first catalyst inlet tube, 121 third reaction section (light hydrocarbon and Circulating oil reaction zone), 122 First reaction zone exit stream conveying pipe, 123 first steam distributor, 124 light hydrocarbon and / or circulating oil inlet, 125 third reaction zone, 125Y baffle zone, 125T cylinder zone, 126 Vertical partition, 127 vertical cylinder, 128 second catalyst inlet tube, 131 first stripping section, 132 second steam distributor, 133 first catalyst stripping zone, 134 bioreactor transfer pipe, 135 stripping Internals, 138 catalyst return pipe, 141 second reaction section, 142 third steam distributor, 143 second stripping section, 144 (after reaction) catalyst outflow pipe, 145 second reaction zone, 146 transfer pipe, 147 holes Plate, 148 second reaction zone catalyst return tube, 149 second baffle, 151 settler, 152 settler dilute phase zone, 153 first stage
  • the multi-stage fluidized catalytic reaction method for petroleum hydrocarbons of the present invention realizes zoned multi-stage reactions in one reactor, including a first-stage reaction, a second-stage reaction, a cracking reaction of light hydrocarbons and / or circulating oil in a reactor; and a catalytic cracking of the raw oil Catalyst replacement and two-stage relay reactions are achieved during the first- and second-stage reactions, and two-stage reactions of light hydrocarbons and / or circulating oils are also achieved; these reactions are realized in different areas within a reactor; light hydrocarbon and / or circulating oil reactions Performed in a separate reaction zone;
  • the reactor includes a first reaction section 110, a catalyst splitter 114, a third reaction section 121, a second reaction section 141, and a settler 151 from bottom to top;
  • the first reaction section 110 is a riser reactor, which adopts the riser reaction conditions, and includes a first catalyst inlet pipe 118, a pre-lift section 112, a feed oil inlet 113, and a first reaction zone 115 for the feed oil Y from bottom to top. Reaction; a first stripping section 131 is set around the first reaction zone 115 of the feedstock oil reaction zone, a stripping internal 135 is provided in the first stripping section 131, and a catalyst splitter 114 is provided at the exit of the first reaction zone 115; And located at the top of the first catalyst stripping zone 133 of the first stripping section 131;
  • a baffle 116 is provided between the first catalyst stripping zone 133 and the third reaction section 121.
  • the baffle 116 is provided with a first reaction zone outlet logistics conveying pipe 122; the first reaction zone outlet logistics conveying pipe 122 is located at the third At the top of the reaction zone 125 is the entrance of the second reaction zone 145; the third reaction section 121 is provided with a light hydrocarbon and / or circulating oil inlet 124 and a second catalyst inlet pipe 128 for light hydrocarbon and circulating oil reactions;
  • the second reaction zone 145 is provided inside the second reaction zone 141, and a second partition plate 149 is provided between the third reaction zone 121 and the outside of the second reaction zone 145 in the second reaction zone 141.
  • the second reaction zone 145 and sedimentation A transfer pipe 146 is provided between the gas-solid separator in the device 151;
  • a stripping member is placed in the catalyst storage area after the reaction in the second reaction area outside the second reaction area, and this storage area serves as the second stripping section 143;
  • the second stripping section 143 is provided with a (out-reacted) catalyst outflow pipe 144,
  • the catalyst separated by the gas-solid separator is directly stripped by the steam S introduced by the third steam distributor 142 in the storage area, and the second stand-by catalyst B2 flows out of the reactor through the (reacted) catalyst outflow pipe 144;
  • the first reaction section 110, the catalyst splitter 114, the first stripping section 131, the third reaction section 121, the third reaction zone 125, the second reaction section 141, and the settler 151 are arranged coaxially above and below; the catalyst is split.
  • the reactor 114 is in the form of a circumferential swirl; a first reaction zone exit logistics conveying pipe 122 may be provided in the third reaction zone 125. As shown in FIG. 2, the first reaction zone exit logistics conveying pipe 122 is located at the center of the third reaction zone 125. Coaxially with the first reaction zone, or a plurality of first reaction zone outlet logistics conveying pipes 122 are provided in the third reaction zone 125, as shown in FIG. 3, four first reaction zone outlet logistics are arranged in parallel. Tube 122;
  • the reaction process goes through the following steps:
  • the first catalyst A1 enters the pre-lifting section 112 of the first reaction section 110 from the first catalyst inlet pipe 118 and is fluidized and conveyed to the first reaction zone 115 by the pre-lifting medium G.
  • the raw material oil Y is atomized by steam through the raw material inlet 113 Entered into the first reaction zone 115 of the first reaction section 110, mixed with the first catalyst A1 to gasify and perform the first-stage catalytic cracking reaction; after completing the first-stage catalytic cracking reaction, a part of the catalyst was separated from the catalyst splitter 114 The separated catalyst settles into the first catalyst stripping zone 133 of the first stripping section 131 below.
  • the second catalyst A2 enters the third reaction zone 125 provided in the third reaction zone 121 between the first stripping section 131 and the second reaction section 141 from the second catalyst inlet pipe 128, and the light hydrocarbon raw material Y1 enters the third reaction zone 125.
  • an independent first-stage catalytic cracking reaction is performed; then, the catalyst P3 in the third reaction zone and the reaction product, that is, the reacted stream P3 and the oil and gas in the first reaction zone enter the second reaction upward.
  • Zone 145 continues the second stage catalytic cracking reaction;
  • the catalyst S is fluidized with the steam S introduced through the first steam distributor 123, so that the zone is formed into a fluidized bed or turbulent fluidized bed state, and these steams simultaneously enter the second reaction zone 145;
  • the gas and catalyst are transferred from the transfer pipe 146 to the settler 151, and the first-stage gas-solid separator 153 and the second-stage gas-solid separator 154 perform gas-solid separation.
  • the reaction product P passes through the product outlet. 156 (does not pass through the dilute phase zone 152 of the settler) out of the settler;
  • the catalyst after the reaction in the second reaction zone is stripped in the second stripping section 143, and the second stand-by catalyst B2 after the stripping flows out of the reactor from the catalyst outflow pipe 144.
  • the catalytic fresh raw material oil Y is atomized by steam, it enters the first reaction zone 115 through the raw material oil inlet 113 and is mixed with the first catalyst A1, which is a regenerant sent through the first catalyst delivery pipe 118. It is gasified, and the catalytic cracking reaction in the first reaction zone is completed in about 1.2 seconds.
  • the raw material reaction product separates about 85% of the catalyst at the catalyst diverter 114 at the exit of the first reaction zone, and the oil and gas enters the first reaction zone.
  • the reaction gas and catalyst in the third reaction zone enter the second reaction zone with the oil and gas from the outlet stream conveying pipe 122 in the first reaction zone to continue the reaction; Flow out of the reactor; the second stand-by catalyst B2 after the reaction in the second reaction zone flows out of the reactor from the catalyst outflow pipe 144 after the second stripping section 143 is stripped;
  • the first reaction zone is designed according to the riser conditions; the second reaction zone is designed based on the circulating fluidized bed conditions, and the apparent gas flow rate is about 2.0m / s; the third reaction zone is designed based on the turbulent fluidized bed conditions, and the apparent gas flow rate is 0.6m /s-1.0m/s, the catalyst space velocity is about 4-6;
  • the reaction temperature in the first reaction zone is about 500 ° C
  • the reaction temperature in the second reaction zone is about 550 ° C
  • the reaction temperature in the third reaction zone is about 590 ° C to 610 ° C;
  • the first catalyst A1 is a regenerant;
  • the second catalyst A2 is a catalyst having a carbon content of 0.1% -0.2%;
  • Steam is replenished from the third reaction zone, and the steam first participates in the reaction of the third reaction zone, and then enters the second reaction zone to reduce the oil and gas partial pressure.
  • FR is a flow recording instrument
  • FRC is a flow control instrument
  • TIC is a temperature control instrument.
  • an orifice plate 147 is provided at the entrance of the second reaction zone between the third reaction zone 121 and the second reaction zone 145.
  • the catalyst and gas in the third reaction zone are distributed to the second reaction zone 145 through the orifice plate 147;
  • the C4 component is introduced into the lower part of the raw material oil inlet 113, and the first catalyst A1 first contacts and reacts with the C4 component, and then flows upward to contact and react with the raw material Y.
  • the other parts of the device structure are the same as those of the first embodiment.
  • the second reaction zone 145 is a riser reaction zone (the upper area of the riser is equivalent to the conveying pipe in FIG. 1), and the riser outlet is directly connected to the first
  • the first-stage gas-solid separator 153 communicates, and a second stripping section and a fourth reaction section 155 are arranged around the reaction zone of the riser, and the fourth reaction section 155 is arranged above the second stripping section 143; light cycle oil raw material Y2 and heavy Circulating oil or refining oil Y3 enters the fourth reaction zone 155;
  • a catalyst return pipe 138 is provided between the third reaction zone 125 and the pre-lifting section 112 of the first reaction section, so that part of the catalyst is refluxed from the third reaction zone to the pre-lifting section 112 to participate in the catalytic reaction of the feed oil Y.
  • the other parts of the device structure are the same as those of the first embodiment.
  • Embodiment 4 is a diagrammatic representation of Embodiment 4:
  • a second reaction zone catalyst return pipe 148 is provided between the second reaction section 141 and the first stripping section 131. After the second reaction zone is reacted, The catalyst is first stripped in the second stripping section 143, and then enters the first catalyst stripping section 133 of the first stripping section 131 from the second reaction zone catalyst return pipe 148, and passes through the catalyst diverter 114 from the first reaction zone. The separated catalyst 115 is stripped together, and after the stripping, it flows out of the reactor from the bioreactor conveying pipe 134;
  • the space between the shell of the third reaction section 121 and the outlet logistics conveying pipe 122 of the first reaction zone is divided into three partitions by two vertical partitions 126 and a vertical cylinder 127.
  • Two vertical partitions 126 are connected to the shell of the third reaction section 121, and the vertical cylinder 127 is supported on the shell of the third reaction section 121.
  • the vertical partition 126 and the vertical cylinder 127 communicate with each other at the bottom to realize the flow of the various sections of the catalyst; a light hydrocarbon and steam inlet is provided at the bottom entrance of the cylinder zone 125T, and the cylinder zone 125T is used as the third reaction zone.
  • the other parts of the device structure are the same as those of the first embodiment.
  • FIG. 1 The device used in this embodiment is shown in FIG. 1.
  • the catalytic feedstock is heavy oil with a processing capacity of 2 million tons / year.
  • the properties of the feedstock are shown in Table 1; the feedstock is preheated at 280 ° C; the reaction time in the first reaction zone is 1.2 seconds, the reaction temperature is 510 ° C, the agent-to-oil ratio is 6.5, and the regenerant A1 The temperature is 680 °C; the amount of atomized steam is 5% of the amount of raw oil;
  • the reaction temperature in the second reaction zone is 550 ° C, the reaction time is 3.5 seconds, the catalyst has a weight hourly space velocity of 8, the agent-to-oil ratio is 7.0, and the carbon content of the regenerant is 0.6%; the amount of supplementary steam is 12% of the amount of raw oil;
  • Catalyst A2 is a semi-regenerant, the temperature is 660 ° C, and the carbon content is 0.15%;
  • the raw material of the third reaction zone is self-produced light gasoline, the temperature is 40 ° C, the amount of light gasoline is 300,000 tons / year, and the reaction temperature is 600 ° C;
  • the first reaction zone is 1.2 meters in diameter and 14 meters in height;
  • the second reaction zone is 3.8 meters in diameter and 9 meters in height;
  • the diameter of the third reaction zone is 5 meters, the diameter is 4 meters, and the diameter of the outlet conveying pipe 122 in the first reaction zone is 1.4 meters;
  • the first stripping section 131 has a diameter of 4.8 meters; the second stripping section 143 has a diameter of 5 meters;
  • the settler part is designed according to the conventional technology, which is controlled by those skilled in the art.
  • reaction conditions are: the reaction temperature in the first reaction zone is 580 ° C, the reaction temperature in the second reaction zone is 560 ° C, the catalyst space velocity in the fluidized bed reaction zone is 4, and the apparent velocity of the gas in the fluidized bed reaction zone is 1.0 m / s.
  • C4 which is a hydrocarbon of four carbon atoms, is a well-known technology in the present invention.
  • the present invention compared with the conventional catalytic process, has significantly lower yields of low-value products such as dry gas and coke, and significantly higher yields of high-value products such as olefins, liquefied gas, and gasoline. Yield decreased by 2.1 percentage points, coke yield decreased by 0.93 percentage points, propylene yield increased by 0.64 percentage points, and selectivity of propylene to dry gas increased by 0.589. It can be seen that the present invention greatly increases the selectivity of propylene, and the dry gas and coke yields are greatly reduced.

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Abstract

一种石油烃多级流化催化反应方法及反应器，该方法在一个反应器内实现分区多级反应，包括原料油一级反应、二级反应、轻烃和/或循环油裂解反应过程，该方法中催化裂化原料油一级和二级反应过程中实现催化剂置换和两级接力反应，也实现轻轻和/或循环油的两级反应；这些反应在一个反应器内不同区域实现：轻烃和/或循环油第一级反应在独立的反应区进行。反应器包括第一反应段、催化剂分流器、第二反应段、第三反应段和沉降器等。通过在反应过程中对催化剂、反应温度的多级梯级选择控制，提高了烯烃产率和选择性，显著降低了焦炭等副产品的收率。

Description

一种石油烃多级流化催化反应方法及反应器 技术领域
本发明属于石油烃类催化转化技术领域，特别涉及一种石油烃多级流化催化反应方法及反应器，能够多产丙烯。
背景技术
催化裂化装置是主要的汽油生产装置，也是丙烯的重要生产装置。丙烯是最重要的石油化工原料之一。借鉴常规重油催化裂化反应-再生系统的操作、设计经验，国内外研究人员开发了一系列重油催化裂化生产丙烯的工艺技术。
中国石油大学(华东)在两段提升管催化裂化技术基础上开发了TMP技术。该技术以重质油为原料，利用两段提升管催化裂化工艺分段反应、催化剂接力和大剂油比的工艺特点，针对不同性质的反应物料进行进料方式的优化组合，控制不同物料适宜的反应条件，以达到提高丙烯产率的目的。国外公司多产丙烯技术也多为采用双提升管方法。
中国石化石油化工科学研究院开发了以重油为原料、丙烯为目的产品的DCC技术，该技术的反应器为提升管加湍流流化床床层串联式反应器。中国石化石油化工科学研究院还在DCC工艺基础上开发了采用新型组合式反应器体系的增强型催化裂解技术(DCC-PLUS)，该技术与DCC工艺相同之处为均采用提升管反应器加流化床床层反应器的形式，不同之处在于DCC-PLUS工艺将再生器来的另外一股再生催化剂引入流化床床层反应器。无论DCC还是DCC-PLUS都将一个反应器分为两个串联的裂化反应区，即提升管和流化床床层都进行催化裂化反应，提升管部分进行原料油裂化反应，流化床床层部分使提升管反应区的中间产物产品进一步进行增产丙烯的二次裂化反应；由于流化床床层反应区的催化剂为原料油反应的待生剂，流化床床层反应区小分子催化裂化性能降低，需要进一步提高反应温度增加热裂化作用来提高丙烯收率，进一步降低了丙烯的选择性；DCC和DCC-PLUS都使用再生剂循环，对固定的原料量流化床床层反应区的空速只能靠流化床床层内的催化剂料位变化控制，流化床床层区必须设置催化剂流出口通过催化剂流出量实现料位控制；由于催化剂携带和气固分离的要求，流化床床层反应区与气固分离器之间全部是稀相空间，并且油气离开流化床床层时仍然携带大量催化剂，而油气离开流化床床层到气固分离器的停留时间20秒以上，采用流化床床层反应时催化剂料位上方催化剂的携带以及油气停留时间都必然造成进一步的副反应止，丙烯进一步热裂解，影响产品分布及丙烯选择性，反应难以及时终 止；另外DCC和DCC-PLUS流化床裂化反应区的催化剂为原料油反应后的积碳催化剂，为了提高丙烯产率需要进一步提高流化床反应区的反应温度，必然导致催化裂化反应受到限制，热反应增加，大幅度降低丙烯的选择性，导致干气和焦炭产率偏高。
已有技术的侧重点均为生产丙烯，分为两类，第一类为提升管加流化床串联式反应，第二类为双提升管并联式反应。现有技术主要集中在通过催化剂和反应条件改变上，如采用比FCC工艺更高的反应温度、剂油比和蒸汽注入量，以提高裂化反应深度和丙烯的选择性。
发明内容
本发明的一个目的在于提供一种石油烃多级流化催化反应方法，实现针对不同性质原料按各自的条件的分级反应，同时实现反应过程中对催化剂和反应温度的分级的控制，提高丙烯选择性，显著降低干气和焦炭等副产品的收率。
本发明的另一目的在于提供了一种石油烃多级流化催化反应器。
本发明技术方案如下：
一种石油烃多级流化催化反应方法，该反应方法在一个反应器内实现分区多级反应，包括原料油一级反应、二级反应、轻烃和/或循环油裂解反应过程；该方法中催化裂化原料油一级和二级反应过程中实现催化剂置换和两级接力反应，也实现轻烃或/和循环油的两级反应；这些反应在一个反应器内不同区域实现；轻烃和/或循环油第一级反应在独立的反应区进行；
催化裂化原料油先进入第1反应区(又称原料油反应区或原料油裂化反应区)在催化剂(再生催化剂)环境下进行第一级催化裂化反应，随即在第1反应区出口进行催化剂分离，油气经设在第三反应段内部的第1反应区出口物流输送管向上进入第2反应区继续深度裂解的第二级催化裂化反应，第一催化剂在原料油入口下方进入第一反应段，在预提升介质作用下进入第1反应区；第2反应区的物流从输送管进入沉降器内进行气固分离；
第1反应区采用提升管反应条件，其出口设置催化剂分流器(旋流气固分流器)；
第二催化剂在第三反应段进入第3反应区(第3反应区设置在第一反应段上方，第3反应区为轻烃、循环油反应区)，轻烃和/或循环油进入第3反应区在第二催化剂环境下，进行独立的第一级催化裂化反应；然后第3反应区的催化剂和反应产物向上进入第2反应区继续进行第二级催化裂化反应(即第3反应区反应后的催化剂继续在第2反应区使用)。
根据本发明的具体实施方案，第3反应区采用湍流流化床或循环流化床反应条件；具体实施时，第二催化剂进入该第3反应区与进入该反应区的反应原料进行催化反应，该第3反应区设置在第1反应区上方的第1反应出口物流输送管周围，从反应器外进入第3反应区的第二催化剂在该区用蒸汽流化，形成湍流流化床或循环流化床条件；第1反应区的反应产物和部分催化剂与第3反应区反应后的催化剂和气体汇合后进入上方的第2反应区，在第2反应区进行第二级催化反应，实现再次裂化转化；
第3反应区的反应产物和催化剂与第1反应区的油气进入第2反应区继续反应。
上述的石油烃多级流化催化反应方法，进一步地，反应过程在一个上下分区的石油烃多级流化催化反应器内完成，经历以下步骤：
第一催化剂从第一催化剂进入管进入第一反应段的预提升段(或催化剂输送区)被预提升介质流化输送到第1反应区，原料油经过原料油入口由蒸汽雾化后进入第一反应段的第1反应区内，与第一催化剂混合接触气化并进行第一级催化裂化反应(具体地，该反应在提升管反应器完成，原料油用蒸汽通过喷嘴雾化成小液滴后在提升管反应区内与催化剂接触并获得热量使原料油进行催化裂化反应，第1反应区主要实现原料油向富含烯烃的汽油和柴油组分转化，尽管并非必须，一般情况下第1反应区的反应时间控制在1.5秒以内，不宜大于2.5秒，反应温度控制在540℃以下)；完成第一级催化裂化反应后，直接从设在第1反应区出口的催化剂分流器分离出部分催化剂(具体实施时，第1反应区的催化剂分流率为70％到90％，保留10％到30％的催化剂存在于反应油气内，与反应油气一起进入第2反应区)；从第1反应出口物流分离出的催化剂沉降进入下方的第一汽提段汽提后从待生剂输送管流出反应器；油气经过第1反应区出口物流输送管进入上方的第2反应区进行第二级催化裂化反应(具体实施时，第二级催化裂化反应可以进一步增加蒸汽比例，降低石油烃组分分压)；
第二催化剂从第二催化剂进入管进入设在第一汽提段和第二反应段之间第三反应段的第3反应区，轻烃和/或循环油进入第3反应区，在第二催化剂环境下进行独立的第一级催化裂化反应；然后第3反应区的催化剂和反应产物向上进入第2反应区继续进行第二级(再次)催化裂化反应；具体地，所述轻烃可为C4组分或终馏点低于210℃的石油烃；
在第3反应区内用蒸汽进行催化剂流化，使该区形成湍流流化床或循环流化床状态，这些蒸汽同时进入第2反应区；
第2反应区反应后的气体和催化剂输送到沉降器内旋风分离器进行气固分离，反应产物流出沉降器；
第2反应区反应后的催化剂从第2反应区催化剂回流管直接进入第一汽提段的第一催化剂汽提区与经催化剂分流器从第1反应区分离的催化剂一起进行汽提，汽提后从待生剂输送管流出反应器；或者在第2反应区外设置第二汽提段，汽提后的第二待生催化剂从催化剂流出管流出反应器。
根据本发明的具体实施方案，多种原料在第3反应区反应时，原料按沸点或终馏点分层进入第3反应区进行反应，沸点或终馏点低的组分在下层进入第3反应区，沸点或终馏点高的组分依次在上一层进入第3反应区；当轻烃和轻循环油组分同时在第3反应区反应时，轻烃在第3三反应区下方进入反应区，轻循环油在轻烃上方进入第3反应区。
上述的石油烃多级流化催化反应方法，进一步地，第1反应区采用提升管即输送床反应形式，第3反应区采用湍流流化床或循环流化床反应形式，第2反应区采用循环流化床或湍流流化床反应形式。以增产丙烯为目的时，所述湍流流化床气体表观流速小于1.2m/s，所述循环流化床气体表观流速1.2-4.0m/s。
上述的石油烃多级流化催化反应方法，进一步地，所述第三反应段内通过竖向隔板和/或竖向筒体竖向分成多个隔离的区即隔板区和/或筒体区(具体地，所述隔板和筒体可连接或支撑在第三反应段的壳体上)；进入第三反应段的催化剂通过各隔板区和/或筒体区底部联通通道分别进入各区；分割成的所述隔板区和/或筒体区全部或部分用作第3反应区；轻烃和/或循环油在上述隔板区和/或筒体区中的一个、多个部分或全部区内反应；当多种原料在第3反应区反应时，不同原料在不同的隔板区或筒体区按各自的条件反应；尤其是当C4组分独立进入第三反应段时，优选在独立的隔板区或筒体区内反应。
上述的石油烃多级流化催化反应方法，进一步地，进入第3反应区的反应物数量即轻烃和/或循环油数量按第2反应区反应后催化剂的含碳量即第2反应区的催化剂活性控制，使第2反应区反应后的催化剂含碳量重量比不大于0.8％。或者进入第3反应区的反应物数量即轻烃和/或循环油数量按第3反应区反应后催化剂的含碳量控制，使第3反应区反应后的催化剂含碳量重量比不大于0.4％。具体实施时，当第3反应区轻烃和循环油反应后不能满足第2反应区催化剂含碳量的要求时，在第3反应区设置催化剂取热管，通过催化剂取热实现第2反应区催化剂的控制要求；当第三反应段分割成为多个隔离的区时，催化剂取热管可以设在作为第3反应区的隔离的区内(即取热管同时设在隔离的反应区内，此反应区内同时取热和反应)，或者设在独立的不作为第3反应区的隔离的区内。流化床内催化剂取热是常见技术，不再详述。
上述的石油烃多级流化催化反应方法，进一步地，第2反应区采用比第1反应区 高或相同的反应温度；优选地，第2反应区反应温度比第1反应区反应温度高0℃到70℃，更优选高10℃到70℃；第2反应区反应温度由第二催化剂数量控制。
上述的石油烃多级流化催化反应方法，进一步地，在第2反应区和第二反应段外壳体之间设置第4反应区，该区用于循环油或回炼油反应。
本发明同时提供了一种石油烃多级流化催化反应器：
反应器自下而上包括第一反应段、催化剂分流器、第三反应段、第二反应段和沉降器，
第一反应段为提升管反应器，自下而上包括第一催化剂进入管、预提升段、原料油入口、第1反应区，用于原料油反应；在原料油反应区即第1反应区周围设置第一汽提段，所述催化剂分流器设置在所述第1反应区出口，并位于所述第一汽提段的第一催化剂汽提区内的顶部；
第一催化剂汽提区与第三反应段之间设置有隔板，隔板上设置有第1反应区出口物流输送管；第二反应段内设有第2反应区，第三反应段内设有第3反应区；所述第1反应区出口物流输送管出口位于第3反应区顶部第2反应区入口处(此第1反应区出口物流输送管位于第3反应区内部)；第三反应段设置有轻烃或循环油入口和第二催化剂进入管，用于轻烃和循环油反应；
第三反应段与第二反应段内的第2反应区外部之间设置有第二隔板，第2反应区与沉降器内气固分离器之间设置输送管。
上述的石油烃多级流化催化反应器，进一步地，第一反应段、催化剂分流器、第一汽提段、第三反应段、第二反应段和沉降器上下同轴布置。
上述的石油烃多级流化催化反应器，进一步地，第三反应段和第2反应区之间设置孔板，来自第1反应区的物流和来自第3反应区的催化剂和气体通过该孔板分布到第2反应区。
上述的石油烃多级流化催化反应器，进一步地，在第二反应段与第一汽提段之间设置有第2反应区催化剂回流管，从而使第2反应区反应后的催化剂直接回流到第一汽提段；或第二反应段与第一汽提段之间设置第2反应区催化剂回流管，并在第2反应区外(即第2反应区和第二反应段外壳体之间)设置第二汽提段，从而使第2反应区反应后的催化剂在第二汽提段内汽提后从第2反应区催化剂回流管进入第一汽提段，然后从待生剂输送管流出反应器。具体实施时，第2反应区外为第2反应区反应后的催化剂储存区，沉降器内的气固分离器分离出的催化剂先进入该储存区，然后进入第一汽提段汽提后流出反应器，或者在该储存区设置汽提构件，该储存区可以作为第二汽提段 使用，该第二汽提段设置催化剂流出管(口)，气固分离器分离出的催化剂直接在该储存区汽提后流出反应器，或者先在该第二汽提段汽提后再进入第一汽提段汽提。
上述的石油烃多级流化催化反应器，进一步地，需要时在第2反应区外围(在第2反应区和第二反应段外壳体之间)设置第4反应区，用于循环油或回炼油的独立反应。当第2反应区外的设置有第二汽提段时，所述第4反应区位于第二汽提段上方。具体实施时，在该第4反应区利用第2反应区催化剂的热量和剩余活性进行重循环油或回炼油催化转化，该反应区采用鼓泡流化床或湍流流化床条件；当C4独立回炼时，C4可以在第1反应区下方的预提升段反应。
上述的石油烃多级流化催化反应器，更进一步地，第三反应段壳体与第1反应区出口物流输送管之间的空间用竖向隔板和/或竖向筒体分成多个隔离的区即隔板区和/或筒体区(所述隔板和筒体可连接或支撑在第三反应段的壳体上)，竖向隔板或竖向筒体底端相互连通，从而实现催化剂各区间流动；在部分或全部隔板区和/或筒体区的底部入口设置轻烃和蒸汽入口，以实现不同原料在不同的隔板区或筒体区按各自的条件反应。具体实施时，全部隔板区和筒体区的底部入口均设置蒸汽入口，而仅在需要实现反应的隔离的区设置轻烃入口以引入轻烃原料。
本发明中：
1、当需要增加第一反应段反应剂油比时，可以在第3反应区或第二汽提段或第一汽提段的第一催化剂汽提区与第一反应段的预提升段之间设置催化剂回流管，使催化剂从第3反应区或第二汽提段或第一催化剂汽提区回流到预提升段；
2、反应器各部件采用上下一体布置，优选(通常情况是)上下同轴布置，第一反应段设在最下方，第三反应段设在第一反应段和第二反应段之间；第二汽提段设在第2反应区外围，沉降器在最上方，沉降器内设置两级气固分离器，第2反应区通过输送管与第一级气固分离器入口连接，第2反应区的反应物流直接通过输送管进入气固分离器，避免反应油气进入沉降器稀相空间，实现快速的气固分离和反应的空速终止；
3、本发明的反应器第一反应段为提升管，包括预提升段、第1反应区、第一汽提段、第1催化剂进入管、待生剂输送管；第二反应段包括第2反应区、(反应后的)催化剂流出管；第三反应段包括第3反应区、第1反应区出口物流输送管、第2催化剂进入管；
4、第三反应段可以设置独立的蒸汽入口，蒸汽从第3反应区进入第2反应区，补充第2反应区需要的蒸汽；
5、本发明所述轻烃为C4组分、催化轻汽油、汽油或石脑油组分，终馏点不高于 210℃；所述循环油包括轻循环油和重循环油组分；所述轻循环油为沸点180℃-255℃石油烃或轻柴油组分；所述重循环油为沸点250℃-365℃的石油烃或重柴油组分；回炼油为沸点340℃-520℃石油烃组分。
本发明的有益效果：
本发明实现不同原料的分区反应，同时实现反应中催化剂置换、反应温度的阶梯控制和对待生剂含碳量的控制反应条件；为不同原料提供相应的反应环境和条件，提高了反应器内催化剂效率，可以明显的降低干气和焦炭产率，提高丙烯产率，提高经济效益；由于第2反应区的反应物是来自第1反应区和第3反应区的气相，第2反应区内催化剂含碳量对反应的选择性至关重要；
本发明方法在一个反应器内实现反应原料、催化剂、反应温度的多级控制，不同原料或不同反应阶段采用相应的流态化反应器形式；设备简单，且不改变常规反应沉降器的布置和安装条件，大幅度降低投资和操作费用。
在一些具体的实施方案中，本发明的方法的反应过程为：
加热到200-350℃的催化流化原料油分几路进入雾化喷嘴，用原料油量3％-5％左右的蒸汽实现雾化后进入最下方的第一反应段提升管反应器，与下方经预提升气体输送来的第一催化剂(再生催化剂)混合，气化后在第1反应区进行1.0到2.0秒左右的催化裂化反应，反应温度为500℃左右，反应热量由再生催化剂提供，同时反应油气和蒸汽在第1反应区内输送催化剂到达第1反应区出口，用旋流分流器分流出部分催化剂，油气直接向上进入设在第三反应段内的第1反应区出口物流输送管；在第三反应段补充第二催化剂，该补充催化剂在第1反应区出口物流输送管出口与来自第1反应区的油气混合并向上流入第2反应区；在第2反应区内实现了催化剂补充，第1反应区产生的油气在新的催化剂条件下继续进行裂解反应，第2反应区内催化剂含碳量重量比控制在0.6％，不大于0.8％，实现了催化剂活性的提高，根据产品要求不同，第2反应区反应时间按1.0秒到6.0秒，同时实现了反应温度的提高，提高了转化效率。
附图说明
图1：本发明的石油烃多级流化催化反应方法实施方式一装置结构示意图；
图2：图1中A-A向视图，即第1反应区出口物流输送管布置方式示意图；
图3：图1中A-A向视图，即第1反应区出口物流输送管第二种布置方式示意图；
图4：本发明的石油烃多级流化催化反应方法实施方式二装置结构示意图，第2反应区入口设置孔板(分布板)；
图5：本发明的石油烃多级流化催化反应方法实施方式三装置结构示意图，设置第4反应区；
图6：本发明的石油烃多级流化催化反应方法实施方式四装置结构示意图，第2反应区反应后的催化剂返回第一汽提段；
图7：图6中第三反应段局部放大图；
图8：图7中B-B向视图，即分成多个隔离的区的第三反应段横截面示意图。
图中编号标记内容如下：
110第一反应段，112预提升段，113原料油入口(喷嘴)，114催化剂分流器，115第1反应区，116隔板，118第一催化剂进入管，121第三反应段(轻烃和循环油反应区)，122第1反应区出口物流输送管，123第一蒸汽分布器，124轻烃和/或循环油入口，125第3反应区，125Y隔板区，125T筒体区，126竖向隔板，127竖向筒体，128第二催化剂进入管，131第一汽提段，132第二蒸汽分布器，133第一催化剂汽提区，134待生剂输送管，135汽提内件，138催化剂回流管，141第二反应段，142第三蒸汽分布器，143第二汽提段，144(反应后的)催化剂流出管，145第2反应区，146输送管，147孔板，148第2反应区催化剂回流管，149第二隔板，151沉降器，152沉降器稀相区，153第一级气固分离器，154第二级气固分离器，155第4反应区(流化床或湍流流化床反应区)，156产品出口；
Y原料油，Y1轻烃原料，Y2轻循环油原料，Y3重循环油或回炼油，G预提升介质，P反应产物，S蒸汽，A1第一催化剂(即进入第一反应段的催化剂)、A2第二催化剂(即进入第2反应区的催化剂)，B1待生催化剂，B2第二待生催化剂(即第2反应区、第3反应区或/和第4反应区反应后的催化剂)；P1(进入第2反应区的)第1反应区反应后的物流；P3(进入第2反应区的)第3反应区反应后的物流；
FR流量记录仪表，FRC流量控制仪表，TIC温度控制仪表。
具体实施方式
以下以具体实施方式和实施例来说明本发明的技术方案，但本发明的保护范围不限于此。
具体实施过程如下：
实施方式一：
如图1-图3所示，
本发明的石油烃多级流化催化反应方法，在一个反应器内实现分区多级反应，包括 原料油一级反应、二级反应、轻烃和/或循环油裂解反应过程；催化裂化原料油一级和二级反应过程中实现催化剂置换和两级接力反应，也实现轻烃或/和循环油的两级反应；这些反应在一个反应器内不同区域实现；轻烃和/或循环油反应在独立的反应区进行；
反应器自下而上包括第一反应段110、催化剂分流器114、第三反应段121、第二反应段141和沉降器151；
第一反应段110为提升管反应器即采用提升管反应条件，自下而上包括第一催化剂进入管118、预提升段112、原料油入口113、第1反应区115，用于原料油Y反应；在原料油反应区即第1反应区115周围设置第一汽提段131，第一汽提段131内设有汽提内件135，催化剂分流器114设置在第1反应区115出口，并位于第一汽提段131的第一催化剂汽提区133内的顶部；
第一催化剂汽提区133与第三反应段121之间设置有隔板116，隔板116上设置有第1反应区出口物流输送管122；第1反应区出口物流输送管122出口位于第3反应区125顶部第2反应区145入口处；第三反应段121设置有轻烃和/或循环油入口124和第二催化剂进入管128，用于轻烃和循环油反应；
第2反应区145设在第二反应段141内部，第三反应段121与第二反应段141内的第2反应区145外部之间设置有第二隔板149，第2反应区145与沉降器151内气固分离器之间设置输送管146；
在第2反应区外的第2反应区反应后的催化剂储存区置汽提构件，该储存区作为第二汽提段143；第二汽提段143设置(反应后的)催化剂流出管144，气固分离器分离出的催化剂直接在该储存区经第三蒸汽分布器142引入的蒸汽S汽提后，第二待生催化剂B2经(反应后的)催化剂流出管144流出反应器；
具体实施时，第一反应段110、催化剂分流器114、第一汽提段131、第三反应段121、第3反应区125、第二反应段141和沉降器151上下同轴布置；催化剂分流器114为周向旋流形式；可以在第3反应区125内设置一个第1反应区出口物流输送管122，如图2所示，第1反应区出口物流输送管122设在第3反应区125中心，与第1反应区同轴，或者在第3反应区125内设置多个第1反应区出口物流输送管122，如图3所示，设置四个并列均布的第1反应区出口物流输送管122；
反应过程经历以下步骤：
第一催化剂A1从第一催化剂进入管118进入第一反应段110的预提升段112被预提升介质G流化输送到第1反应区115，原料油Y经过原料油入口113由蒸汽雾化后进入第一反应段110的第1反应区115内，与第一催化剂A1混合接触气化并进行第一 级催化裂化反应；完成第一级催化裂化反应后，从催化剂分流器114分离出部分催化剂，分离出的催化剂沉降进入下方的第一汽提段131的第一催化剂汽提区133，在第一催化剂汽提区133内经第二蒸汽分布器132引入的蒸汽S汽提后，待生催化剂B1从待生剂输送管134流出反应器；油气(包含在第1反应区反应后的物流P1中)经过第1反应区出口物流输送管122进入上方的第2反应区145进行第二级催化裂化反应；
第二催化剂A2从第二催化剂进入管128进入设在第一汽提段131和第二反应段141之间第三反应段121的第3反应区125，轻烃原料Y1进入第3反应区125，在第二催化剂A2环境下进行独立的第一级催化裂化反应；然后第3反应区的催化剂和反应产物即第3反应区反应后的物流P3与第1反应区的油气向上进入第2反应区145继续进行第二级催化裂化反应；
在第3反应区125内用经第一蒸汽分布器123引入的蒸汽S进行催化剂流化，使该区形成流化床或湍流流化床状态，这些蒸汽同时进入第2反应区145；
第2反应区反应后的气体和催化剂从输送管146输送到沉降器151内，在第一级气固分离器153和第二级气固分离器154进行气固分离，反应产物P经产品出口156(不经过沉降器稀相区152)流出沉降器；
第2反应区反应后的催化剂在第二汽提段143汽提，汽提后的第二待生催化剂B2从催化剂流出管144流出反应器。
本实施方式中，具体地，催化新鲜原料油Y被蒸汽雾化后，经原料油入口113进入第1反应区115，与经第一催化剂输送管118送来的再生剂即第一催化剂A1混合并气化，在第1反应区经过约1.2秒完成该区的催化裂化反应，原料油反应产物在第1反应区出口的催化剂分流器114分离出85％左右的催化剂，油气进入第1反应区出口物流输送管122；另一催化剂即第二催化剂A2从第2催化剂进入管128进入第3反应区125底部，轻烃原料Y1即C4和轻汽油进入第3反应区在催化剂A2环境下进行轻烃反应，第3反应区的反应气体和催化剂与来自第1反应区出口物流输送管122的油气一起进入第2反应区继续反应；第1反应区的待生催化剂B1从待生剂输送管134流出反应器；第2反应区反应后的第二待生催化剂B2在第二汽提段143汽提后从催化剂流出管144流出反应器；
第1反应区按提升管条件设计；第2反应区按循环流化床条件设计，气体表观流速2.0m/s左右；第3反应区按湍流流化床条件设计，气体表观流速0.6m/s-1.0m/s，催化剂空速4-6左右；
第1反应区的反应温度约500℃，第2反应区的反应温度约550℃左右，第3反应 区的反应温度约590℃-610℃；
第一催化剂A1为再生剂；第二催化剂A2为含碳量为0.1％-0.2％的催化剂；
从第3反应区补充蒸汽，该蒸汽先参与第3反应区的反应，然后进入第2反应区，降低油气分压。
图中，FR为流量记录仪表，FRC为流量控制仪表，TIC为温度控制仪表。
实施方式二：
如图4所示，石油烃多级流化催化反应方法，在第三反应段121和第2反应区145之间的第2反应区入口设置孔板147，来自第1反应区的物流和来自第3反应区的催化剂和气体通过该孔板147分布到第2反应区145；
在原料油入口113下部先引入C4组分，第一催化剂A1先与C4组分接触反应，之后再向上流动与原料油Y接触反应。
其他部分装置结构与实施方式一相同。
实施方式三：
如图5所示，石油烃多级流化催化反应方法，第2反应区145为提升管反应区(该提升管上部区域相当于图1中的输送管)，提升管出口直接与沉降器第一级气固分离器153连通，在提升管反应区外围设置第二汽提段和第4反应区155，第4反应区155设在第二汽提段143上方；轻循环油原料Y2和重循环油或回炼油Y3进入第4反应区155；
在第3反应区125与第一反应段的预提升段112之间设置催化剂回流管138，使部分催化剂从第3反应区回流到预提升段112，参与原料油Y的催化反应。
其他部分装置结构与实施方式一相同。
实施方式四：
如图6-图8所示，石油烃多级流化催化反应方法，在第二反应段141与第一汽提段131之间设置第2反应区催化剂回流管148，第2反应区反应后的催化剂先在第二汽提段143汽提后，从第2反应区催化剂回流管148进入第一汽提段131的第一催化剂汽提区133，与经催化剂分流器114从第1反应区115分离的催化剂一起进行汽提，汽提后从待生剂输送管134流出反应器；
如图7-图8所示，第三反应段121的壳体与第1反应区出口物流输送管122之间的空间用两个竖向隔板126和一个竖向筒体127分成三个隔离的区即两个隔板区125Y和一个筒体区125T，两个竖向隔板126连接在第三反应段121的壳体上，竖向筒体127 支撑在第三反应段121的壳体上，竖向隔板126和竖向筒体127底端相互连通，实现催化剂各区间流动；在筒体区125T的底部入口设置轻烃和蒸汽入口，筒体区125T用作第3反应区，轻烃在筒体区125T内反应；在两个隔板区125Y底部入口设置有蒸汽入口；具体实施时，第二催化剂A2从第2催化剂进入管128进入第3反应区125(即筒体区125T)底部，轻烃原料Y1进入筒体区125T在催化剂A2环境下进行轻烃反应；
其他部分装置结构与实施方式一相同。
实施例：
本实施例使用的装置见图1。
催化原料油为重油，加工量200万吨/年，原料油性质见表1；原料油预热280℃；第1反应区反应时间1.2秒，反应温度510℃，剂油比6.5，再生剂A1温度680℃；雾化蒸汽量为原料油量的5％；
第2反应区反应温度550℃，反应时间3.5秒，催化剂重时空速8，剂油比7.0，待生剂含碳0.6％；补充蒸汽量为原料油量的12％；
第1反应区反应后的催化剂分离率为85％，15％保留在油气中，催化剂A2为半再生剂，温度660℃，含碳0.15％；
第3反应区原料为自产轻汽油，温度40℃，轻汽油量30万吨/年，反应温度600℃；
第1反应区直径1.2米，反应区高度14米；
第2反应区直径3.8米，高度9米；
第3反应区直径5米，直径4米，第1反应区出口物流输送管122直径1.4米；
第一汽提段131直径4.8米；第二汽提段143直径5米；
沉降器部分按常规技术设计，本领域普通技术人员掌握。
反应条件及产物分布如表2。
对比例：
使用已有催化裂化工艺，即采用一个提升管与流化床反应串联的反应技术；
反应条件为：第1反应区反应温度580℃，第2反应区即流化床反应区反应温度560℃，流化床反应区催化剂空速4，流化床反应区气体表观流速1.0m/s。
反应条件及产物分布如表2所示。
本发明所述C4即四个碳原子的烃，为公知技术。
表1催化原料油性质
项目 数据
密度g/cm 3(20℃) 0.9035
残碳，w％ 0.62
氢含量，w％ 12.56
硫含量，w％ 0.31
氮含量，w％ 0.16
馏程，℃ 256～545
表2实施例与对比例反应条件及产物分布对比
项目 实施例 对比例
第1反应区出口温度，℃ 510 580
第2反应区反应温度，℃ 550 560
第3反应区反应温度，℃ 600  
再生温度，℃ 680 680
蒸汽比例，％ 17 24
原料油温度，℃ 280 340
半再生剂温度，℃ 660  
半再生剂含碳，％ 0.15  
产物分布％    
干气(H 2～C2) 7.1 9.2
液化气(C3～C4) 36.8 35.33
汽油 28.2 28.81
柴油 15.1 12.19
重油 3.6 4.28
焦炭 8.8 9.73
损失 0.4 0.46
丙烯 16.13 15.49
丙烯/干气 2.272465 1.683696
从表2的对比结果可以看出，本发明与常规催化工艺相比：低价值产品如干气和焦炭产率明显下降，高价值产品如烯烃、液化气和汽油产率明显提高，其中干气产率下降了2.1个百分点，焦炭产率降低0.93个百分点，丙烯产率增加了0.64个百分点，丙烯对干气的选择性增加了0.589。可见，本发明大幅的增加了丙烯选择性，干气和焦炭产率大幅下降。

Claims (16)

  1. 一种石油烃多级流化催化反应方法，其特征在于：
    该反应方法在一个反应器内实现分区多级反应，包括原料油一级反应、二级反应、轻烃和/或循环油裂解反应过程；该方法中催化裂化原料油一级和二级反应过程中实现催化剂置换和两级接力反应，也实现轻烃或/和循环油的两级反应；这些反应在一个反应器内不同区域实现；轻烃和/或循环油第一级反应在独立的反应区进行；
    催化裂化原料油(Y)先进入第1反应区(115)在催化剂环境下进行第一级催化裂化反应，随即在第1反应区出口进行催化剂分离，油气经设在第三反应段(121)内部的第1反应区出口物流输送管(122)向上进入第2反应区(145)继续深度裂解的第二级催化裂化反应，第一催化剂(A1)在原料油入口(113)下方进入第一反应段(110)，在预提升介质(G)作用下进入第1反应区(115)；第2反应区的物流进入沉降器(151)内进行气固分离；
    第1反应区(115)采用提升管反应条件，其出口设置催化剂分流器(114)；
    第二催化剂(A2)在第三反应段(121)进入第3反应区(125)，轻烃和/或循环油进入第3反应区(125)在第二催化剂(A2)环境下，进行独立的第一级催化裂化反应；然后第3反应区的催化剂和反应产物向上进入第2反应区(145)继续进行第二级催化裂化反应。
  2. 如权利要求1所述的石油烃多级流化催化反应方法，其特征在于：
    反应过程在一个上下分区的石油烃多级流化催化反应器内完成，经历以下步骤：
    第一催化剂(A1)从第一催化剂进入管(118)进入第一反应段(110)的预提升段(112)，被预提升介质(G)流化输送到第1反应区(115)，原料油(Y)经过原料油入口(113)由蒸汽雾化后进入第一反应段(110)的第1反应区(115)内，与第一催化剂(A1)混合接触气化并进行第一级催化裂化反应；完成第一级催化裂化反应后，直接从设在第1反应区出口的催化剂分流器(114)分离出部分催化剂，分离出的催化剂沉降进入下方的第一汽提段(131)汽提后从待生剂输送管(134)流出反应器；油气经过第1反应区出口物流输送管(122)进入上方的第2反应区(145)进行第二级催化裂化反应；
    第二催化剂(A2)从第二催化剂进入管(128)进入设在第一汽提段(131)和第二反应段(141)之间第三反应段(121)的第3反应区(125)，轻烃和/或循环油进入第3反应区(125)，在第二催化剂(A2)环境下进行独立的第一级催化裂化反应；然 后第3反应区的催化剂和反应产物(P3)向上进入第2反应区(145)继续进行第二级催化裂化反应；
    在第3反应区(125)内用蒸汽进行催化剂流化，使该区形成流化床或湍流流化床状态，这些蒸汽同时进入第2反应区(145)；
    第2反应区反应后的气体和催化剂输送到沉降器(151)内进行气固分离，反应产物(P)流出沉降器；
    第2反应区反应后的催化剂从第2反应区催化剂回流管(148)直接进入第一汽提段(131)的第一催化剂汽提区(133)与经催化剂分流器(114)从第1反应区(115)分离的催化剂一起进行汽提，汽提后从待生剂输送管(134)流出反应器；或者在第2反应区外设置第二汽提段(143)，汽提后的第二待生催化剂(B2)从催化剂流出管144流出反应器。
  3. 如权利要求1或2所述的石油烃多级流化催化反应方法，其特征在于，多种原料在第3反应区(125)反应时，原料按沸点或终馏点分层进入第3反应区(125)进行反应，沸点或终馏点低的组分在下层进入第3反应区(125)，沸点或终馏点高的组分依次在上一层进入第3反应区(125)；当轻烃和轻循环油组分同时在第3反应区(125)反应时，轻烃在第3反应区下方进入反应区，轻循环油在轻烃上方进入第3反应区(125)。
  4. 如权利要求1或2或3所述的石油烃多级流化催化反应方法，其特征在于，第1反应区(115)采用提升管即输送床反应形式，第3反应区(125)采用湍流流化床或循环流化床反应形式，第2反应区(145)采用循环流化床或湍流流化床反应形式；可选择性地，所述湍流流化床气体表观流速小于1.2m/s，所述循环流化床气体表观流速1.2-4.0m/s。
  5. 如权利要求1所述的石油烃多级流化催化反应方法，其特征在于，所述第三反应段(121)内通过竖向隔板(126)和/或竖向筒体(127)竖向分成多个隔离的区即隔板区(125Y)和/或筒体区(125T)；进入第三反应段(121)的催化剂通过各隔板区(125Y)和/或筒体区(125T)底部联通通道分别进入各区；分割成的所述隔板区(125Y)和/或筒体区(125T)全部或部分用作第3反应区；轻烃和/或循环油在上述隔板区(125Y)和/或筒体区(125T)中的一个、多个部分或全部区内反应；当多种原料在第3反应区反应时，不同原料在不同的隔板区(125Y)或筒体区(125T)按各自的条件反应。
  6. 如权利要求1～5任一项所述的石油烃多级流化催化反应方法，其特征在于，进 入第3反应区(125)的反应物数量即轻烃和/或循环油数量按第2反应区反应后催化剂的含碳量即第2反应区的催化剂活性控制，使第2反应区反应后的催化剂含碳量重量比不大于0.8％。
  7. 如权利要求1～5任一项所述的石油烃多级流化催化反应方法，其特征在于，进入第3反应区(125)的反应物数量即轻烃和/或循环油数量按第3反应区反应后催化剂的含碳量控制，使第3反应区反应后的催化剂含碳量重量比不大于0.4％。
  8. 如权利要求1～5任一项所述的石油烃多级流化催化反应方法，其特征在于，第2反应区(145)采用比第1反应区(115)高或相同的反应温度；第2反应区反应温度比第1反应区反应温度高0℃到70℃；第2反应区(145)反应温度由第二催化剂(A2)数量控制。
  9. 如权利要求1～5任一项所述的石油烃多级流化催化反应方法，其特征在于，在第2反应区(145)和第二反应段(141)外壳体之间设置第4反应区(155)，该区用于循环油或回炼油反应。
  10. 一种石油烃多级流化催化反应器，其特征在于，反应器自下而上包括第一反应段(110)、催化剂分流器(114)、第三反应段(121)、第二反应段(141)和沉降器(151)；
    第一反应段(110)为提升管反应器，自下而上包括第一催化剂进入管(118)、预提升段(112)、原料油入口(113)、第1反应区(115)，用于原料油(Y)反应；在原料油反应区即第1反应区(115)周围设置第一汽提段(131)，所述催化剂分流器(114)设置在所述第1反应区(115)出口，并位于所述第一汽提段(131)的第一催化剂汽提区(133)内的顶部；
    第一催化剂汽提区(133)与第三反应段(121)之间设置有隔板(116)，隔板(116)上设置有第1反应区出口物流输送管(122)；第二反应段(141)内设有第2反应区(145)，第三反应段(121)内设有第3反应区(125)；所述第1反应区出口物流输送管(122)出口位于第3反应区(125)顶部第2反应区(145)入口处；第三反应段(121)设置有轻烃和/或循环油入口(124)和第二催化剂进入管(128)，用于轻烃和循环油反应；
    第三反应段(121)与第二反应段(141)内的第2反应区(145)外部之间设置有第二隔板(149)，第2反应区(145)与沉降器(151)内气固分离器之间设置输送管(146)。
  11. 如权利要求10所述的石油烃多级流化催化反应器，其特征在于，第一反应段 (110)、催化剂分流器(114)、第一汽提段(131)、第三反应段(121)、第二反应段(141)和沉降器(151)上下同轴布置。
  12. 如权利要求10所述的石油烃多级流化催化反应器，其特征在于，第三反应段(121)和第2反应区(145)之间设置孔板(147)，来自第1反应区的物流和来自第3反应区的催化剂和气体通过该孔板(147)分布到第2反应区(145)。
  13. 如权利要求10所述的石油烃多级流化催化反应器，其特征在于，在第二反应段(141)与第一汽提段(131)之间设置有第2反应区催化剂回流管(148)；或第二反应段(141)与第一汽提段(131)之间设置第2反应区催化剂回流管(148)，并在第2反应区(145)外设置第二汽提段(143)。
  14. 如权利要求10所述的石油烃多级流化催化反应器，其特征在于，在第2反应区(145)和第二反应段(141)外壳体之间设置第4反应区(155)，用于循环油或回炼油反应。
  15. 如权利要求10所述的石油烃多级流化催化反应器，其特征在于，第三反应段(121)的壳体与第1反应区出口物流输送管(122)之间的空间用竖向隔板(126)和/或竖向筒体(127)分成多个隔离的区即隔板区(125Y)和/或筒体区(125T)，竖向隔板(126)或竖向筒体(127)底端相互连通；在部分或全部隔板区(125Y)和/或筒体区(125T)的底部入口设置轻烃和蒸汽入口。
  16. 如权利要求15所述的石油烃多级流化催化反应器，其特征在于，所述隔板、筒体连接或支撑在第三反应段的壳体上。
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Cited By (1)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN113355935A (zh) * 2021-06-22 2021-09-07 中国联合装备集团安阳机械有限公司 一种造纸行业用外置式装锅器、蒸煮锅及其使用方法

Families Citing this family (4)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN114195612A (zh) * 2020-09-02 2022-03-18 青岛京润石化设计研究院有限公司 一种石油烃催化转化制丙烯和乙烯的方法及其装置
CN113244861B (zh) * 2021-06-03 2021-09-24 山东胜星新能源有限公司 一种石油烃多级流化催化反应器
CN116196850A (zh) * 2021-12-01 2023-06-02 中国石油天然气股份有限公司 一种石油烃催化转化制丙烯的装置和方法
CN116196848A (zh) * 2021-12-01 2023-06-02 中国石油天然气股份有限公司 一种原料油和轻烃催化转化的装置和方法

Citations (6)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN1760342A (zh) * 2005-10-12 2006-04-19 洛阳石化设备研究所 一种多产丙烯的催化裂化方法和装置
CN1912067A (zh) * 2005-08-08 2007-02-14 洛阳石化设备研究所 一种催化剂接力使用的催化裂化转化方法及其装置
CN1912068A (zh) * 2005-08-08 2007-02-14 洛阳石化设备研究所 一种重油两段催化裂化转化和汽油改质偶合的方法及装置
CN102465006A (zh) * 2010-11-11 2012-05-23 石宝珍 一种催化裂化方法及装置
CN102485841A (zh) * 2010-12-01 2012-06-06 石宝珍 一种催化裂化方法及装置
WO2013054173A1 (en) * 2011-10-12 2013-04-18 Indian Oil Corporation Ltd. A process for production of c3 olefin in a fluid catalytic cracking unit

Family Cites Families (4)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN102746880A (zh) * 2011-04-20 2012-10-24 中国石油化工股份有限公司 轻烃和重油耦合催化裂化制汽油、柴油和乙烯、丙烯的方法
CN102399579B (zh) * 2011-10-14 2014-03-05 石宝珍 一种催化裂化方法
EP3284804B1 (en) * 2016-08-19 2020-04-22 INDIAN OIL CORPORATION Ltd. Production of propylene in a fluid catalytic cracking unit
CN107286972B (zh) * 2017-07-21 2019-11-22 石宝珍 一种多产丙烯的催化转化方法

Patent Citations (6)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN1912067A (zh) * 2005-08-08 2007-02-14 洛阳石化设备研究所 一种催化剂接力使用的催化裂化转化方法及其装置
CN1912068A (zh) * 2005-08-08 2007-02-14 洛阳石化设备研究所 一种重油两段催化裂化转化和汽油改质偶合的方法及装置
CN1760342A (zh) * 2005-10-12 2006-04-19 洛阳石化设备研究所 一种多产丙烯的催化裂化方法和装置
CN102465006A (zh) * 2010-11-11 2012-05-23 石宝珍 一种催化裂化方法及装置
CN102485841A (zh) * 2010-12-01 2012-06-06 石宝珍 一种催化裂化方法及装置
WO2013054173A1 (en) * 2011-10-12 2013-04-18 Indian Oil Corporation Ltd. A process for production of c3 olefin in a fluid catalytic cracking unit

Cited By (1)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN113355935A (zh) * 2021-06-22 2021-09-07 中国联合装备集团安阳机械有限公司 一种造纸行业用外置式装锅器、蒸煮锅及其使用方法

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