WO2025060210A1 - 催化裂解的装置和方法 - Google Patents

催化裂解的装置和方法 Download PDF

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Publication number
WO2025060210A1
WO2025060210A1 PCT/CN2023/130672 CN2023130672W WO2025060210A1 WO 2025060210 A1 WO2025060210 A1 WO 2025060210A1 CN 2023130672 W CN2023130672 W CN 2023130672W WO 2025060210 A1 WO2025060210 A1 WO 2025060210A1
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Prior art keywords
cracking
reaction
reaction zone
section
catalyst
Prior art date
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Ceased
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PCT/CN2023/130672
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English (en)
French (fr)
Inventor
张兆前
王剑
张忠东
柳召永
王骞
孙志国
迟克彬
孟凡芳
余倩倩
王瑞浦
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Petrochina Co Ltd
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Petrochina Co Ltd
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Priority to EP23952856.5A priority Critical patent/EP4632045A4/en
Priority to KR1020257031047A priority patent/KR20250151473A/ko
Publication of WO2025060210A1 publication Critical patent/WO2025060210A1/zh
Anticipated expiration legal-status Critical
Ceased legal-status Critical Current

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Classifications

    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J8/00—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes
    • B01J8/18—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles
    • B01J8/24—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
    • B01J8/26—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
    • 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
    • B01J19/00—Chemical, physical or physico-chemical processes in general; Their relevant apparatus
    • B01J19/24—Stationary reactors without moving elements inside
    • B01J19/2405—Stationary reactors without moving elements inside provoking a turbulent flow of the reactants, such as in cyclones, or having a high Reynolds-number
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J8/00—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes
    • B01J8/18—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles
    • B01J8/24—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
    • B01J8/44—Fluidisation grids
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G11/00—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
    • C10G11/14—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts
    • C10G11/18—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts according to the "fluidised-bed" technique
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • 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
    • C10G51/06—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
    • C10G55/00—Treatment of hydrocarbon oils, in the absence of hydrogen, by at least one refining process and at least one cracking process

Definitions

  • the present invention relates to the field of catalytic technology, and in particular to a device and method for catalytic cracking.
  • variable diameter risers arched (convex) distribution plates, special-shaped (basin-shaped) distribution plates, concave distribution plates and mushroom-head distributors have been introduced, which can make the high-speed oil and gas and catalyst coming out of the riser form a fully uniform steady-state fluidization, and some catalysts form a stable fluidized bed; however, the expanded diameter section of the riser and the upper riser are generally connected by a reducing circle. It is connected to a trapezoidal outer cylinder, and no internal components are installed in this section to adjust the fluidization of gas and solid.
  • US5043522A and US5846403A disclose injecting catalytic cracking gasoline into the upstream of the feedstock oil nozzle and using a high-temperature, high-activity regeneration catalyst to catalytically convert it;
  • CN1160746A also discloses a method of injecting low-quality gasoline such as straight-run gasoline and coking gasoline into the lower part of the riser reaction to give priority to the reaction of the gasoline.
  • the first method is to use a separate gasoline or light hydrocarbon riser for refining.
  • CN1069054A and US3784463A both use a double riser reactor to carry out the reaction, injecting low-quality gasoline including catalytic cracking crude gasoline into the gasoline riser reactor, and using high temperature and large agent-to-oil ratio reaction conditions to achieve catalytic modification of low-quality gasoline to increase the liquefied gas yield and gasoline octane number; the molecules generated after the light hydrocarbon feed contacts the catalyst are relatively small, and the volume expansion is obvious, which affects the heavy oil feed at the top of the catalytic cracking riser reactor and the contact between the heavy oil and the catalyst.
  • the first method is to keep the heavy oil reaction from having a negative impact and the proportion of light hydrocarbon recycling is low;
  • the second method is to add a separate light hydrocarbon riser on the basis of the traditional single riser reactor, so that the light hydrocarbons can undergo a secondary cracking reaction separately and the separate optimized reactions of different raw materials can be achieved in the separators.
  • the traditional riser has poor control over the reaction depth. For example, if the riser outlet temperature is increased to achieve a higher conversion depth, over-cracking reaction is likely to occur; if the reaction temperature at the riser outlet is low, the ideal conversion depth cannot be achieved.
  • the purpose of the present invention is to provide a catalytic cracking device and method, which can strengthen the gas-solid mixing and contact in the heavy oil cracking process and the light hydrocarbon cracking process, improve the catalytic reaction efficiency, realize the close coupling of the heavy oil reactor and the light hydrocarbon reactor, and improve the flexibility of the operation of regulating the product properties and product distribution. sex.
  • the inventors of the present invention unexpectedly discovered in the study of catalytic cracking that after a heavy oil reactor using an expanded diameter riser is coupled in parallel with a light hydrocarbon cracking reactor, the catalyst after cracking of light hydrocarbons is recycled from the cracking catalyst outlet to the expanded diameter section of the heavy oil reactor, which not only allows the heavy oil cracking process and the light hydrocarbon cracking process to be carried out independently, thereby increasing the recycling ratio of light hydrocarbons, but also helps to strengthen the catalytic reaction of the heavy oil; by arranging an internal conveying pipe inside the expanded diameter section of the heavy oil reactor, the gas-solid fluidization state of the expanded diameter section can be changed and optimized, and the adjustment range of the gas linear velocity and the catalyst bed density in the expanded diameter section can be widened; and the turbulent reaction section is set in the cracking settler of the light hydrocarbon cracking reactor, which is also conducive to the fully enhanced contact between the light hydrocarbons and the cracking catalyst in the reaction section, and can achieve efficient contact reaction between the light hydrocarbons and the catalyst at a
  • the first aspect of the present invention provides a catalytic cracking device, which includes a multi-stage reaction unit and a cracking reaction unit, the multi-stage reaction unit includes a multi-variable diameter riser and a multi-stage reaction settler, and the cracking reaction unit includes a cracking riser and a cracking settler;
  • the multi-variable diameter riser includes a first reaction zone, a second reaction zone and a third reaction zone connected in series in sequence along its length direction, the outlet of the third reaction zone is connected to the multi-stage reaction settler, the inner diameter of the second reaction zone is larger than the inner diameters of the first reaction zone and the third reaction zone, and the connection between the third reaction zone and the second reaction zone is provided with an inner conveying pipe extending into the interior of the second reaction zone;
  • the cracking settler includes a cracking stripping section, a turbulent reaction section and a cracking sedimentation section arranged in sequence along its length direction, the outlet of the cracking riser is located in the turbulent
  • the ratio of the inner diameter of the inner transport pipe to that of the third reaction zone is 0.9-1.2:1, and the ratio of the length of the inner transport pipe to that of the second reaction zone is 0.15-0.8:1.
  • the ratio of the inner diameters of the first reaction zone, the second reaction zone and the third reaction zone is 1:2-5:0.7-1.5, and the ratio of the lengths is 1:0.3-2:0.5-2; the ratio of the inner diameters of the turbulent reaction section, the cracking stripping section and the cracking sedimentation section is 1:0.4-0.8:2.5-5, and the ratio of the lengths is 1:1.5-3:1.5-4; the ratio of the lengths of the multi-variable riser and the cracking riser is 1:0.5-1.2.
  • At least one inner delivery hole is provided on the side wall of the inner delivery pipe, and the distance between the inner delivery hole and the entrance of the third reaction zone is 0.5-2 m.
  • the ratio of the total area of at least one of the inner delivery holes to the cross-sectional area of the inner delivery tube is 0.05-0.3:1.
  • a distribution internal component connected to the first reaction zone is provided in the second reaction zone, and a plurality of distribution holes are provided on the distribution internal component.
  • the distribution internal component includes a distribution cylinder connected to the first reaction zone and an end distribution plate connected to the distribution cylinder, the distribution cylinder and the end distribution plate are both provided with the distribution holes, and the end distribution plate is located at one end of the distribution cylinder away from the first reaction zone.
  • the ratio of the sum of the areas of the distribution holes on the distribution cylinder to the cross-sectional area of the first reaction zone is 0.6-4:1; the ratio of the sum of the areas of the distribution holes on the end distribution plate to the vertical projection area of the end distribution plate is 0.1-0.3:1.
  • a side channel extending toward the outside of the distribution tube is provided at the distribution holes on the distribution tube, and an opening section of the side channel away from the distribution tube is an inclined surface toward the first reaction zone.
  • the ratio of the area of the end distribution plate to the cross-sectional area of the second reaction zone is 0.7-0.85:1.
  • the first reaction zone, the second reaction zone, the third reaction zone and the inner transport pipe are coaxially arranged, and the cracking catalyst outlet is located at the bottom of the cracking stripping section and is connected to the bottom of the second reaction zone.
  • the multi-variable diameter riser also includes a multi-stage reaction pre-lifting section connected to the inlet of the first reaction zone
  • the cracking riser includes a cracking pre-lifting section and a cracking reaction section connected in series along its length direction, the outlet of the cracking reaction section is located at one end of the turbulent reaction section close to the cracking stripping section, the transition section between the multi-stage reaction pre-lifting section and the first reaction zone is connected to a multi-stage reaction feed unit, and the transition section between the cracking pre-lifting section and the cracking reaction section is connected to a cracking feed unit.
  • a cracking distributor is provided at the outlet of the cracking reaction section, and the gas phase outlet of the cracking sedimentation section is connected to the multi-stage reaction settler.
  • the device also includes a regeneration unit, and the multistage reaction settler is provided with a multistage reaction catalyst outlet connected to the feed inlet of the regeneration unit, and the regeneration catalyst outlet of the regeneration unit is connected to the multi-variable riser and/or the cracking riser.
  • a second aspect of the present invention provides a method for catalytic cracking, using the above-mentioned device, the method comprising the following steps:
  • a multi-stage reaction mixture formed by mixing heavy raw materials and a multi-stage reaction catalyst is input from the inlet of the first reaction zone, enters the first reaction zone and the second reaction zone in sequence, and then enters the third reaction zone from the inner delivery pipe to perform a multi-stage catalytic reaction to obtain a multi-stage reactant.
  • the multi-stage reactant is subjected to multi-stage reaction separation in the multi-stage reaction settler to obtain a multi-stage reaction oil and gas product and a multi-stage reaction catalyst to be produced;
  • a cracking mixture formed by mixing a light raw material and a cracking catalyst is input from the inlet of the cracking riser, undergoes a cracking catalytic reaction through the cracking riser to obtain a cracking reactant, and after the cracking reactant enters the turbulent reaction section for a turbulent reaction, undergoes cracking separation through the cracking settling section, and undergoes cracking stripping through the cracking stripping section to obtain a cracking oil and gas product and a cracking catalyst to be produced;
  • the catalyst to be produced by cracking is introduced into the second reaction zone from the cracking catalyst outlet and mixed with the multi-stage reaction mixture.
  • the inlet temperature of the inner transport pipe is 510-540°C
  • the outlet temperature of the cracking riser is 540-620°C.
  • the inlet temperature of the inner transport pipe is 550-640°C
  • the outlet temperature of the cracking riser is 560-660°C.
  • the residence time of the multi-stage reaction mixture in the first reaction zone is 0.8-2s, the residence time in the second reaction zone is 5-15s, and the residence time in the third reaction zone is 0.5-1.5s; the residence time of the cracking mixture in the cracking riser is 0.8-3.5s, and the residence time in the turbulent reaction section is 3-13s.
  • the catalytic cracking device provided by the present invention can independently carry out the heavy oil cracking process and the light hydrocarbon cracking process through the multi-stage reaction unit and the cracking reaction unit arranged in parallel, thereby increasing the recycling ratio of light hydrocarbons, and recycling the catalyst after the light hydrocarbon cracking from the cracking catalyst outlet to the second reaction zone, which is helpful to increase the flow rate and activity of the catalyst in the second reaction zone, strengthen the catalytic reaction of the heavy oil, and realize the close coupling of the heavy oil cracking process and the light hydrocarbon cracking process; by arranging the inner conveying pipe inside the multi-variable diameter riser of the multi-stage reaction unit, the gas-solid fluidization state of the second reaction zone can be changed and optimized, and the adjustment range of the gas linear velocity and the catalyst bed density in the second reaction zone can be widened, and a turbulent reaction section is arranged in the cracking settler of the cracking reaction unit.
  • the cracking catalyst density of the turbulent reaction section is relatively large, which is conducive to the fully enhanced contact between the light hydrocarbons and the cracking catalyst in the reaction section, and can realize the efficient contact reaction of the light hydrocarbons and the catalyst at a relatively low temperature and low space velocity, and avoid the excessive cracking reaction of the light hydrocarbons caused by the reaction temperature being too high; in the device provided by the present invention
  • the coupling of the two reaction units can simultaneously perform independent cracking/cracking operation control on heavy oil and light hydrocarbons, with more operating methods, realizing various flexible operations to control product properties and product distribution.
  • FIG1 is a schematic structural diagram of a specific embodiment of a catalytic cracking device in the present invention.
  • FIG. 2 is a schematic structural diagram of a specific implementation of the internal delivery pipe and the distribution internal component in the present invention.
  • connection and “connected” should be understood in a broad sense, for example, it can be a fixed connection, a detachable connection, or an integral connection; it can be a direct connection, or an indirect connection through an intermediate medium, it can be abutment, it can be the internal connection of two elements or the interaction relationship between two elements.
  • connection and “connected” should be understood in a broad sense, for example, it can be a fixed connection, a detachable connection, or an integral connection; it can be a direct connection, or an indirect connection through an intermediate medium, it can be abutment, it can be the internal connection of two elements or the interaction relationship between two elements.
  • the catalytic cracking device comprises a multistage reaction unit 1 and a cracking reaction unit 2, wherein the multistage reaction unit 1 comprises a multi-variable riser 11 and a multistage reaction settler 12, and the cracking reaction unit 2 comprises a cracking riser 21 and a cracking settler 23; the multi-variable riser 11 comprises a first reaction zone 11a, a second reaction zone 11b and a third reaction zone 11c connected in series in a length direction thereof, the outlet of the third reaction zone 11c is connected to the multistage reaction settler 12, and the outlet of the second reaction zone 11a is connected to the multistage reaction settler 12, and the outlet of the third reaction zone 11c is connected to the multistage reaction settler 12.
  • the inner diameter of 11b is larger than the inner diameters of the first reaction zone 11a and the third reaction zone 11c, and an inner transport pipe 13 extending into the second reaction zone 11b is arranged at the connection between the third reaction zone 11c and the second reaction zone 11b;
  • the cracking settler 23 includes a cracking stripping section 23a, a turbulent reaction section 23b and a cracking sedimentation section 23c arranged in sequence along its length direction, the outlet of the cracking riser 21 is located in the turbulent reaction section 23b, and the cracking stripping section 23a is provided with a cracking catalyst outlet 25 connected to the second reaction zone 11b.
  • the catalytic cracking device provided by the present invention can be applied to catalytic cracking or catalytic cracking processes in various chemical production fields, and is particularly suitable for catalytic cracking/catalytic cracking of heavy oil and light hydrocarbons, for example, producing ultra-low olefin gasoline and high-yield cracking gas; its reaction unit 1 is used as a heavy raw material (heavy oil) for catalytic cracking/catalytic cracking, and the cracking reaction unit 2 is used as a light raw material (light hydrocarbon) for catalytic cracking/catalytic cracking.
  • the catalytic cracking device in the present application is based on a parallel arrangement of a multi-stage reaction unit 1 and a cracking reaction unit 2, and ultra-low olefin gasoline and high-yield cracking can be achieved by replacing the catalyst and adjusting the operating conditions.
  • the switching of the two operation modes of gas can also increase the adjustment depth and variation range of the two modes, and realize a processing volume with a larger proportion of light hydrocarbons relative to heavy oil; wherein, the light hydrocarbons can come from outside the multistage reaction unit 1 in the device or even outside the use workshop or factory.
  • the catalyst after the cracking of the light hydrocarbons is recycled to the second reaction zone 11b from the cracking catalyst outlet 25, which can make full use of the high-temperature spent catalyst after the light hydrocarbon reaction with high cracking activity, and replenish it to the second reaction zone 11b of the multistage reaction unit 1, effectively improving the catalyst concentration and catalyst activity inside it, and realizing the close coupling of the multistage reaction unit 1 and the cracking reaction unit 2.
  • the above defects can be effectively overcome, and the following effects are produced: First, when the superficial gas velocity in the second reaction zone 11b is reduced to form a turbulent bed, when the upper interface of the turbulent bed submerges the bottom port of the internal delivery pipe 13, the catalyst on the top of the turbulent bed can be sucked away, and the height of the turbulent bed in the second reaction zone 11b can be controlled. Secondly, the upper part of the expanded diameter section at the upper part of the second reaction zone 11b, which is coaxial with the inner conveying pipe 13 and has the same height, can be used as the dilute phase section of the turbulent bed.
  • the gas carries a large amount of particulate solids and scatters them to the dilute phase section above the turbulent bed.
  • a free space for the solid particles to settle is formed, and the catalyst particles entrained by the turbulent bed will fall into the lower part of the second reaction zone 11b again, thereby enhancing the back mixing of the catalyst.
  • a buffer space can be provided for the gas pulsation in the multi-variable riser 11. The above two effects can, to a certain extent, avoid the rapid increase and sharp fluctuation of a large amount of catalyst in the multi-variable riser 11 and the gas pressure, thereby widening the gas linear velocity range for optimal operation in the second reaction zone 11b.
  • the gas-solid fluidization in the second reaction zone 11b is a turbulent bed.
  • the solid phase is violently turbulent and is in a continuous phase.
  • the solid-gas ratio per unit space is large, the bubble diameter of the solid particles in the turbulent bed is small, and the gas-solid contact efficiency is enhanced.
  • the gas-solid contact efficiency is higher than that of various forms of conveying beds, and it is suitable for some fields that require further enhancement of the gas-solid contact efficiency.
  • the catalyst thrown up by the high-speed gas below can be made to land in the second reaction zone 11b at the bottom, thereby increasing the catalyst density at the bottom of the second reaction zone 11b and enhancing gas-solid mixing and contact.
  • the inner transport pipe 13 provided by the present invention is The flow state of the gas and the catalyst in the second reaction zone 11b is changed.
  • the gas superficial velocity range in the second reaction zone 11b is wider.
  • the superficial gas velocity of the conventional MIP-CGP expansion section is 1.4-2.8 m/s, and the industrial operation device is generally preferably 1.5-2.0 m/s; in the multi-stage reaction unit 1 provided by the present invention, the superficial gas velocity in the second reaction zone 11b can form a dense phase section at the bottom of the second reaction zone 11b at 0.7-3.5 m/s, and the preferred operating range of the superficial gas velocity is 0.9-2.5 m/s; correspondingly, the ratio of the inner diameter of the second reaction zone 11b to the first reaction zone 11a in the present invention is larger.
  • the ratio of the inner diameter of the second reaction zone 11b to the first reaction zone 11a is 2-5:1, and more preferably 1.7-3.5:1.
  • the multi-stage reaction unit 1 and the cracking reaction unit 2 are usually arranged vertically, the length direction of the multi-variable riser 11 is the vertical direction, the first reaction zone 11a, the second reaction zone 11b and the third reaction zone 11c are connected in sequence from bottom to top, the length direction of the cracking settler 23 is also the vertical direction, the cracking stripping section 23a, the turbulent reaction section 23b and the cracking sedimentation section 23c are connected in sequence from bottom to top; the multi-stage reaction settler 12 and the multi-variable riser 11 can be coaxially arranged or arranged in parallel.
  • the multi-stage reaction settler 12 generally includes a multi-stage reaction stripping section 12a and a multi-stage reaction settling section 12b arranged from bottom to top, and the multi-stage reaction settling section 12b is provided with a single-stage or multi-stage reaction cyclone separator 12c, so that the material of the multi-variable diameter riser 11 can enter the multi-stage reaction settling section 12b and be separated by the multi-stage reaction cyclone separator 12c, so that the separated catalyst to be regenerated enters the multi-stage reaction stripping section 12a and is stripped by the multi-stage reaction stripping steam F and then discharged, and can enter the catalyst regenerator for air charring regeneration treatment and then return to the multi-stage reaction unit 1 for recycling.
  • the basic implementation of the catalytic cracking device provided above, when applied to catalytic cracking, comprises the following steps:
  • the heavy raw material is atomized by steam and mixed with a multi-stage reaction catalyst at the bottom of the first reaction zone 11a to form a multi-stage reaction mixture.
  • the multi-stage reaction mixture is input from the inlet of the first reaction zone 11a, and enters the first reaction zone 11a and the second reaction zone 11b in sequence.
  • the expansion of the second reaction zone 11b can prolong the residence time of the multi-stage reaction mixture and fully mix and contact, fully complete the secondary reaction of the heavy oil and its intermediate fraction, and then enter the third reaction zone 11c from the inner delivery pipe 13 to perform a multi-stage catalytic reaction to obtain a multi-stage reactant.
  • the mixture quickly enters the multi-stage reaction settler 12 for multi-stage reaction separation to obtain a multi-stage reaction oil and gas product and a multi-stage reaction catalyst to be produced.
  • the multi-stage reaction catalyst to be produced with carbon after the reaction is discharged with or without steam stripping;
  • the cracking mixture formed by mixing the light raw material and the cracking catalyst is input from the inlet of the cracking riser 21, and undergoes a cracking catalytic reaction through the cracking riser 21 to obtain a cracking reactant.
  • the cracking reactant enters the turbulent reaction section 23b for a turbulent reaction, and then undergoes cracking separation through the cracking settling section 23c, and undergoes cracking stripping through the cracking stripping section 23a to obtain a cracking oil and gas product and a cracking catalyst to be produced.
  • the catalyst to be produced by cracking is introduced into the second reaction zone 11b from the cracking catalyst outlet 25 and mixed with the multi-stage reaction mixture to increase the activity and density of the multi-stage reaction catalyst in the second reaction zone 11b.
  • the heavy raw material is petroleum hydrocarbon and/or other mineral oil
  • the petroleum hydrocarbon is selected from at least one of vacuum gas oil (VGO), atmospheric gas oil (AGO), coker gas oil (CGO), deasphalted oil (DAO), vacuum residue oil (VR), atmospheric residue oil (AR) and hydrogenated heavy oil
  • other mineral oils are selected from at least one of coal liquefaction oil, oil sand oil and shale oil.
  • the heavy raw material is selected from at least one of vacuum gas oil, atmospheric gas oil, coker gas oil, deasphalted oil, vacuum residue oil, atmospheric residue oil and hydrogenated heavy oil; wherein VGO, AGO, CGO, DAO, VR, AR are unhydrogenated whole fractions or partial fractions, or are hydrogenated whole fractions or partial fractions.
  • the multi-stage reaction catalyst and the cracking catalyst in the method can be all types of catalytic cracking or catalytic cracking catalysts, and the active components thereof are selected from one or more catalysts selected from Y or HY type zeolites containing or not containing rare earth, ultrastable Y type zeolites containing or not containing rare earth, ZSM-5 series zeolites or high silicon zeolites with a five-membered ring structure prepared by other methods, and amorphous silicon-aluminum catalysts.
  • the multi-stage reaction catalyst and the cracking catalyst can be the same catalyst or different catalysts.
  • the catalyst in the cracking reaction unit 2 falls into the turbulent reaction section 23b, so that the catalyst density of the turbulent reaction section 23b and the cracking stripping section 23a is relatively large.
  • the catalyst that falls into the cracking stripping section 23a can be output from the cracking catalyst outlet 25 after or without steam stripping.
  • the cracking catalyst outlet 25 is connected to any position from the end of the first reaction zone 11a to the middle of the second reaction zone 11b through a pipeline to input the cracking catalyst to be generated into the second reaction zone 11b of the multi-diameter riser 11.
  • a pipeline can be added to connect the cracking catalyst outlet 25 to the catalyst regenerator for recovery.
  • the cracking catalyst outlet 25 is located at the bottom of the cracking stripping section 23a and is connected to the bottom of the second reaction zone 11b (one end close to the first reaction zone 11a) through a pipeline.
  • a sliding valve can be provided on the pipeline to adjust the amount of cracking catalyst to be input into the second reaction zone 11b by controlling the opening of the sliding valve, and the catalyst stock in the cracking settler 23 can also be adjusted.
  • the cracking catalyst Since the cracking catalyst has a low coke content and a high temperature after the reaction with light hydrocarbons, it still has a high catalytic activity.
  • the flow rate of the catalyst and its catalytic activity are increased, which is beneficial to strengthening the catalytic reaction in the multi-variable riser 11; the oil and gas after the reaction of the light hydrocarbons separated by the cracking settler 23 do not enter the various reaction zones of the multi-variable riser 11 for the heavy oil reaction, and do not affect the oil and gas flow rate and flow velocity of the second reaction zone 11b, and thus do not affect the gas-solid fluidization of the second reaction zone 11b, and at the same time shorten the residence time of the light hydrocarbons in the high-temperature reaction zone in the cracking reaction unit 2, thereby reducing the degree of thermal cracking reaction.
  • the length of the inner delivery pipe 13 is less than the length of the second reaction zone 11b.
  • the ratio of the length of the inner delivery pipe 13 to the second reaction zone 11b is 0.15-0.8:1, more preferably 0.2-0.6:1, which can better make the gas-solid fluidization state and turbulence effect in the second reaction zone 11b better.
  • the second reaction zone 11b can be cylindrical, or a truncated cone with a gradually expanding diameter from bottom to top, or a cylindrical shape with two or three stages of expansion;
  • the height of the second reaction zone 11b refers to the sum of the heights of its own region, the transition section between it and the first reaction zone 11a, and the transition section between it and the third reaction zone 11c.
  • the distance between the bottom port of the inner transport pipe 13 and the bottom of the second reaction zone 11b is 2-8m, preferably 3-6m; the length of the inner transport pipe 13 is 2-7m, preferably 3-5m; at this time, the height of the turbulent bed in the second reaction zone 11b can be better controlled.
  • the side wall effect is more obvious; in the lower part of the expansion section, the catalyst particle velocity is relatively high in the region of r/R ⁇ 0.7, but the axial velocity of the catalyst particles close to the side wall is zero or even negative; in the upper part of the expansion section, the catalyst particle velocity is relatively high in the region of r/R ⁇ 0.6, but the catalyst particle velocity close to the side wall is maintained at around 0, that is, a large part of the catalyst is suspended in this section, and this part of the catalyst may stay in the expansion section for too long, affecting the best process effect.
  • At least one inner delivery hole 13a is provided on the side wall of the inner delivery pipe 13, and the distance between the inner delivery hole 13a and the entrance of the third reaction zone 11c is 0.5-2m.
  • the setting of the inner delivery hole 13a enables the gas in the upper part of the second reaction zone 11b of the expanded diameter section and the coaxial part of the inner delivery pipe 13 to be in a flowing state, so as to prevent the appearance of a dead zone without gas flow in the upper part of the second reaction zone 11b, so as to further improve the turbulence effect and catalytic reaction efficiency in the second reaction zone 11b.
  • a plurality of inner delivery holes 13a are provided, and they can be evenly distributed circumferentially on the inner delivery tube 13. Further preferably, the ratio of the total area of at least one inner delivery hole 13a to the cross-sectional area of the inner delivery tube 13 is 0.05-0.3:1.
  • the cross-sectional area of the inner delivery tube 13 refers to the area of the cross section perpendicular to the length direction of the inner delivery tube 13.
  • the transition section between the first reaction zone 11a and the second reaction zone 11b, and the transition section between the second reaction zone 11b and the third reaction zone 11c are dome-shaped or truncated cone-shaped, and are coaxially arranged with the entire multi-stage reaction unit 1.
  • the first reaction zone 11a, the second reaction zone 11b, the third reaction zone 11c and The inner delivery pipe 13 is coaxially arranged.
  • the ratio of the inner diameters of the first reaction zone 11a, the second reaction zone 11b and the third reaction zone 11c is 1:2-5:0.7-1.5, and the ratio of the lengths is 1:0.3-2:0.5-2. In this case, it is more conducive to prolonging the residence time of the oil agent in the second reaction zone 11b and fully mixing and contacting, so as to fully complete the secondary reaction of the heavy oil and its intermediate fractions.
  • the cracking settler 23 is a two-stage expansion structure, the outer diameters of the cylinders of the cracking stripping section 23a, the turbulent reaction section 23b and the cracking settling section 23c are increased successively, and the cracking riser 21 is located in the turbulent reaction section 23b or in the transition section between the turbulent reaction section 23b and the cracking settling section 23c.
  • the ratio of the inner diameters of the turbulent reaction section 23b, the cracking stripping section 23a and the cracking settling section 23c is 1:0.4-0.8:2.5-5, and the ratio of the lengths is 1:1.5-3:1.5-4. At this time, it is more conducive to the efficient and sufficient reaction of light hydrocarbons and catalysts in the turbulent reaction section 23b, and the efficient contact reaction of light hydrocarbons and catalysts at a lower temperature and low space velocity is achieved.
  • the cracking riser 21 and the cracking settler 23 of the cracking reaction unit 2 are preferably coaxially arranged.
  • the ratio of the length of the multi-diameter riser 11 to the cracking riser 21 is 1:0.5-1.2, more preferably 1:0.6-1.
  • the ratio of the inner diameter of the first reaction zone 11a in the multi-diameter riser 11 to the inner diameter of the cracking riser 21 is 1:0.2-1.
  • the coupling effect between the multi-stage reaction unit 1 and the cracking reaction unit 2 can be better achieved.
  • the height of the multi-diameter riser 11 in a medium-sized device in a laboratory is about 20 meters; the height of the variable diameter riser 11 in industrial production is about 40 meters.
  • a distribution internal member 14 connected to the first reaction zone 11a is provided in the second reaction zone 11b, and a plurality of distribution holes 14a are provided on the distribution internal member 14, that is, the distribution internal member 14 is an open-hole distribution member.
  • the main body of the distribution internal member 14 is located in the transition section between the first reaction zone 11a and the second reaction zone 11b, and its function is to rapidly reduce the high-speed gas-solid particles in the first reaction zone 11a to the gas-solid fluidized state required by the second reaction zone 11b.
  • the distribution internal component 14 can be an open distribution plate with an opening rate of 8-35%; or a cap-type distributor with a mushroom-shaped, conical or umbrella-shaped top and outputting oil and gas from a side channel can be used.
  • the cap-type distributor can be a conventional umbrella-shaped distributor, a mushroom-shaped distributor or a conical cap-shaped distributor.
  • the distribution internal component 14 includes a distribution cylinder 14b connected to the first reaction zone 11a and an end distribution plate 14c connected to the distribution cylinder 14b.
  • the distribution cylinder 14b and the end distribution plate 14c are both provided with a distribution
  • the end distribution plate 14c is located at the end of the distribution tube 14b away from the first reaction zone 11a.
  • the end distribution plate 14c can be in an umbrella-shaped or dome-shaped form, for example, an arched (convex) distribution plate, a special-shaped (basin-shaped) distribution plate, or a concave distribution plate.
  • the ratio of the sum of the areas of the multiple distribution holes 14a opened on the end distribution plate 14c (i.e., the total hole area) to the vertical projection area of the end distribution plate 14c is 0.1-0.3:1, which can make the gas above the end distribution plate 14c flow smoothly and prevent the accumulation of catalyst on the end distribution plate 14c.
  • the vertical projection area of the end distribution plate 14c refers to the area occupied by the shadow of the outer contour of the end distribution plate 14c on the vertical plane after the horizontal parallel light irradiates the end distribution plate 14c.
  • the distribution holes 14a on the distribution cylinder 14b may be rectangular or elliptical.
  • the ratio of the sum of the areas of the distribution holes 14a on the distribution cylinder 14b to the cross-sectional area of the first reaction zone 11a is 0.6-4:1, which can better improve the deceleration and distribution effect of the high-speed gas-solid particles in the first reaction zone 11a.
  • the cross-sectional area of the first reaction zone 11a refers to the area of the cross section perpendicular to the length direction of the first reaction zone 11a.
  • a side channel 14d extending to the outside of the distribution tube 14b is provided at the distribution hole 14a on the distribution tube 14b, and the opening section of the side channel 14d away from the distribution tube 14b is an inclined surface facing the first reaction zone 11a.
  • the cross section of the side channel 14d corresponds to the distribution hole 14a on the distribution tube 14b, and can be rectangular or elliptical.
  • the opening section of the side channel 14d away from the distribution tube 14b is generally not perpendicular to the axis of the side channel 14d, and the opening section is an inclined surface facing the first reaction zone 11a, that is, the extension line of the section is at a certain oblique upward angle with the axis of the multi-stage reaction unit 1; the upper part of the opening section is a part of the side channel 14d, so as to avoid the outlet of the side channel 14d from being blocked by the catalyst as much as possible, and ensure the patency of the outlet of the side channel 14d.
  • the ratio of the area of the end distribution plate 14c to the cross-sectional area of the second reaction zone 11b is 0.7-0.85:1, so that the gas velocity at the edge of the end distribution plate 14c is above 2.5m/s, preferably above 3.5m/s, to ensure that no catalyst accumulation occurs at the bottom of the second reaction zone 11b or even in the transition section to cause an operation accident of the device.
  • the area of the end distribution plate 14c is the projected area of the outer contour of the end distribution plate 14c
  • the cross-sectional area of the second reaction zone 11b is the area of the cross section perpendicular to the length direction of the second reaction zone 11b.
  • a wear-resistant lining such as a ceramic wear-resistant lining, is installed or coated in the inner conveying hole 13a and/or the distribution hole 14a.
  • the multi-variable riser 11 also includes a multi-stage reaction pre-lifting section 11d connected to the inlet of the first reaction zone 11a
  • the cracking riser 21 includes a cracking pre-lifting section 21b and a cracking reaction section 21a connected in series along its length direction, and the outlet of the cracking reaction section 21a is located at one end of the turbulent reaction section 23b close to the cracking stripping section 23a, the transition section between the multi-stage reaction pre-lifting section 11d and the first reaction zone 11a is connected to the multi-stage reaction feed unit 7, and the transition section between the cracking pre-lifting section 21b and the cracking reaction section 21a is connected to the cracking feed unit 8.
  • Multistage reaction pre-lifting gas A is injected through the multistage reaction pre-lifting section 11d, mixed with the multistage reaction catalyst and then ascended, and the heavy raw material atomized by water vapor enters the multi-variable diameter riser 11 from the multistage reaction feeding unit 7, so that the heavy raw material and the multistage reaction catalyst contact and react in the first reaction zone 11a; cracking pre-lifting gas B is injected through the cracking pre-lifting section 21b, mixed with the cracking catalyst and then ascended, and the light raw material atomized by water vapor enters the cracking riser 21 from the cracking feeding unit 8, so that the light raw material and the cracking catalyst contact and react in the cracking reaction section 21a.
  • a cracking distributor 22 is provided at the outlet of the cracking reaction section 21a.
  • the gas phase outlet of the cracking and settling section 23c can be merged into the main oil and gas pipeline, or can be communicated with the multi-stage reaction settler 12.
  • a gas collecting chamber is provided at the top of the multi-stage reaction settler 12, and the gas phase outlet of the cracking and settling section 23c is communicated with the gas collecting chamber through the cracking gas phase delivery pipe 24, so that the reacted oil and gas E (including the reacted oil and gas of the heavy raw material and the reacted oil and gas of the light raw material) are discharged from the gas collecting chamber.
  • the bottom of the cracking and stripping section 23a can be provided with a stripping steam unit 10 to pass steam to strip the catalyst in the cracking and stripping section 23a.
  • the multi-stage reaction unit 1 and the cracking reaction unit can be respectively equipped with corresponding catalyst regenerators.
  • the two can also share a catalyst regenerator.
  • the device also includes a regeneration unit 3, and the multi-stage reaction settler 12 is provided with a multi-stage reaction catalyst outlet 15 connected to the feed inlet of the regeneration unit 3, and the regeneration catalyst outlet 31 of the regeneration unit 3 is connected to the multi-variable riser 11 and/or the cracking riser 21.
  • the cracking catalyst outlet 25 is not only connected to the second reaction zone 11b, but also connected to the feed inlet of the regeneration unit 3, so that part of the catalyst to be regenerated of the cracking reaction unit 2 is directly input into the regeneration unit 3.
  • a regeneration cyclone separator 32 is arranged in the regeneration unit 3, and a regeneration air inlet 33 is arranged at the bottom and a regeneration flue gas outlet 34 is arranged at the top.
  • the catalyst to be regenerated enters the regeneration unit 3 through the regeneration pipeline 5, contacts with the regeneration air C entering through the regeneration air inlet 33 to undergo an oxidation reaction, and the regeneration flue gas D is output from the regeneration flue gas outlet 34 to enter the flue gas turbine for further energy recovery and treatment;
  • the high-temperature regenerated catalyst generated by burning the coke enters the catalyst heat exchanger 4 through the regeneration pipeline 6, and the low-temperature water vapor G enters the catalyst heat exchanger 4 from the heat extraction steam inlet to extract heat from the regenerated catalyst to become high-temperature water vapor H, which is output from the heat extraction steam outlet.
  • the cooled regenerated catalyst flows back to the bottom of the first reaction zone 11a and the bottom of the cracking riser 21 through the regeneration pipeline 6 for recycling.
  • part or all of the multi-stage reaction catalyst and part or all of the cracking catalyst adopt the regenerated catalyst outputted from the regenerated catalyst outlet 31 .
  • the second aspect of the present invention provides a method for applying the device to catalytic cracking, the method comprising the following steps:
  • the multi-stage reaction mixture formed by mixing the heavy raw material and the multi-stage reaction catalyst is input from the inlet of the first reaction zone 11a, and enters the first reaction zone 11a and the second reaction zone 11b in sequence, and then enters the third reaction zone 11c from the inner delivery pipe 13 to perform a multi-stage catalytic reaction to obtain a multi-stage reactant.
  • the multi-stage reactant is subjected to a multi-stage reaction separation in the multi-stage reaction settler 12 to obtain a multi-stage reaction oil and gas product and a multi-stage reaction catalyst to be produced;
  • the cracking mixture formed by mixing the light raw material and the cracking catalyst is input from the inlet of the cracking riser 21, and undergoes a cracking catalytic reaction through the cracking riser 21 to obtain a cracking reactant.
  • the cracking reactant enters the turbulent reaction section 23b for a turbulent reaction, and then undergoes cracking separation through the cracking settling section 23c, and undergoes cracking stripping through the cracking stripping section 23a to obtain a cracking oil and gas product and a cracking catalyst to be produced.
  • the catalyst to be produced by cracking is introduced into the second reaction zone 11b from the cracking catalyst outlet 25 and mixed with the multi-stage reaction mixture.
  • the multi-stage reaction unit 1 adopts a catalyst with high gasoline production and olefin reduction function, controls the catalyst circulation amount, the quenching agent injection amount of the quenching oil and gas unit 9, does not supplement or supplements an appropriate amount of regeneration agent or a cooled regeneration agent (the temperature is generally 560-660° C.) or a carbon-containing catalyst to be regenerated with a carbon content of ⁇ 0.5% at the bottom or middle of the second reaction zone 11b, and controls the temperature in the second reaction zone 11b to be 510-540° C., with an optimization range of 515-530° C., and controls the inlet temperature of the inner delivery pipe 13 to be 510-540° C., with an optimization range of 515-530° C., so as to realize the production mode of ultra-low olefin gasoline;
  • the cracking reaction unit 2 cracks various olefin-rich catalytic light gasoline, coker gasoline, etc., which can further reduce the amount of high-olefin gasoline throughout the plant and reduce the olefin content in the gasoline pool of the whole plant.
  • the outlet temperature of the cracking riser 21 is controlled at 540-620°C, preferably 560-600°C.
  • the multi-stage reaction unit 1 uses a catalytic cracking catalyst that produces high-yield cracking gas, and the bottom or middle of the second reaction zone 11b is not supplemented or supplemented with an appropriate amount of regeneration agent or a cooled regeneration agent (temperature is generally 580-680°C) or a carbon-containing catalyst to be regenerated with a carbon content of ⁇ 0.5% (temperature is generally 580-680°C), and the temperature in the second reaction zone 11b is controlled at 560-660°C, with an optimization range of 560-600°C, and the inlet temperature of the inner delivery pipe 13 is controlled at 550-640°C, with an optimization range of 560-600°C, so that the high-yield cracking gas operation mode can be achieved;
  • the cracking reaction unit 2 recycles light hydrocarbons rich in C4-C8 olefins, or gasoline or diesel, and the outlet temperature of the cracking riser 21 is controlled at 560-660°C, preferably 580-640°C; in this way, the cracking gas output of the whole plant can be further increased, that is, the output of raw materials required for light hydrocarbon processing units such as MTBE, light gasoline etherification, and polypropylene can be increased, and the refining product structure of the whole plant can be adjusted.
  • the temperature of the regenerated catalyst before the heavy feedstock or light feedstock contacts the catalyst is 660-760°C.
  • the apparent residence time of oil and gas in the multi-stage reaction mixture in the first reaction zone 11a is 0.8-2s
  • the apparent residence time of oil and gas in the second reaction zone 11b is 5-15s
  • the apparent residence time of oil and gas in the third reaction zone 11c is 0.5-1.5s
  • the residence time of the cracking mixture in the cracking riser 21 is 0.8-3.5s, more preferably 1-2.5s
  • the residence time in the turbulent reaction section 23b is 3-13s.
  • the catalytic cracking device includes a multi-stage reaction unit 1, a cracking reaction unit 2 and a regeneration unit 3, the multi-stage reaction unit 1 includes a multi-variable diameter riser 11 and a multi-stage reaction settler 12; the multi-variable diameter riser 11 includes a multi-stage reaction pre-lifting section 11d, a first reaction zone 11a, a second reaction zone 11b and a third reaction zone 11c, which are connected in series and coaxially along the length direction thereof, and the outlet of the third reaction zone 11c is connected to the multi-stage reaction settler 12 ...
  • the inner diameter ratio of the first reaction zone 11a, the second reaction zone 11b and the third reaction zone 11c is 1:2-5:0.7-1.5, and the length ratio is 1:0.3-2:0.5-2
  • the connection between the third reaction zone 11c and the second reaction zone 11b is provided with an inner delivery pipe 13 extending into the second reaction zone 11b
  • the inner diameter ratio of the inner delivery pipe 13 to the third reaction zone 11c is 0.9-1.2:1
  • the length ratio of the inner delivery pipe 13 to the second reaction zone 11b is 0.15-0.8:1
  • the inner delivery pipe 13 is provided at the connection between the third reaction zone 11c and the second reaction zone 11b.
  • a plurality of inner delivery holes 13a are arranged on the side wall, the distance between the inner delivery holes 13a and the entrance of the third reaction zone 11c is 0.5-2m, the ratio of the total area of the plurality of inner delivery holes 13a to the cross-sectional area of the inner delivery pipe 13 is 0.05-0.3:1, a distribution internal component 14 connected to the first reaction zone 11a is arranged in the second reaction zone 11b, the distribution internal component 14 comprises a distribution cylinder 14b connected to the first reaction zone 11a and an end distribution plate 14c connected to the distribution cylinder 14b, a plurality of distribution holes 14a are arranged on the distribution cylinder 14b and the end distribution plate 14c, the end distribution plate 14c is located at one end of the distribution cylinder 14b away from the first reaction zone 11a, the ratio of the area of the end distribution plate 14c to the cross-sectional area of the second reaction zone 11b is 0.7-0.85:1, the surface area of the plurality of distribution holes 14a opened on the end distribution plate 14c is The ratio of the sum of
  • the cracking reaction unit 2 includes a cracking riser 21 and a cracking settler 23 which are coaxially arranged.
  • the cracking riser 21 includes a cracking pre-lifting section 21b and a cracking reaction section 21a which are sequentially connected in series along its length direction.
  • the transition section between the cracking pre-lifting section 21b and the cracking reaction section 21a is connected to the cracking feeding unit 8.
  • the cracking settler 23 includes a cracking stripping section 23a, a turbulent reaction section 23b and a cracking settling section 23c which are sequentially arranged along its length direction.
  • the outlet of the cracking reaction section 21a is located at one end of the turbulent reaction section 23b close to the cracking stripping section 23a.
  • a cracking settling section 23c is provided at the outlet of the cracking reaction section 21a.
  • the distributor 22, the bottom of the cracking stripping section 23a is provided with a cracking catalyst outlet 25 connected to the bottom of the second reaction zone 11b, the inner diameter ratio of the turbulent reaction section 23b, the cracking stripping section 23a and the cracking settling section 23c is 1:0.4-0.8:2.5-5, and the length ratio is 1:1.5-3:1.5-4, a single-stage or two-stage cracking cyclone separator 23d is arranged in the cracking settling section 23c, the gas phase outlet of the cracking settling section 23c is connected to the gas collecting chamber at the top of the multi-stage reaction settler 12 through the cracking gas phase conveying pipe 24, and the bottom of the cracking stripping section 23a is provided with a stripping steam unit 10;
  • the regeneration unit 3 is provided with a regeneration cyclone separator 32, and a regeneration catalyst outlet 31 is provided on the side wall, a regeneration air inlet 33 is provided at the bottom, and a regeneration flue gas outlet 34 is provided at the top.
  • the multi-stage reaction catalyst outlet 15 and the cracking catalyst outlet 25 are connected to the feed port of the regeneration unit 3 through the waiting pipeline 5.
  • the outlet 31 is connected to the multi-stage reaction pre-elevation section 11d and the cracking pre-elevation section 21b through a regeneration pipeline 6, and a catalyst heat collector 4 is provided on the regeneration pipeline 6;
  • the length ratio of the multi-variable riser 11 to the cracking riser 21 is 1:0.5-1.2, and wear-resistant linings are installed or coated in the inner conveying holes 13a and the distribution holes 14a.
  • heavy raw materials (heavy oil), light hydrocarbons I, and light hydrocarbons II were all provided by a refinery of China Petrochemical Corporation, catalysts I and II were all provided by China Petrochemical Catalyst Co., Ltd., and product distribution and gasoline composition were detected by chromatography; unless otherwise specified, the remaining raw materials were conventional commercial products, and the parameters involved were all based on existing corresponding methods.
  • a relatively preferred specific embodiment of the catalytic cracking device provided above is used to process the heavy feedstock oil in Table 1 using the catalyst I in Table 2 to produce ultra-low olefin gasoline.
  • the properties of the feedstock in Example 1 are shown in Table 1.
  • the recycled light gasoline uses light hydrocarbon A as the light hydrocarbon feedstock, and the recycled light gasoline accounts for 30% of the heavy feedstock, half of which is the light gasoline produced by the device itself, and the other half comes from other devices outside the device.
  • the total length of the multi-variable riser 11 is about 18 meters, the diameter of the multi-stage reaction pre-lifting section 11d is 0.25 meters, and the height is 1.5 meters.
  • the diameter of the first reaction zone 11a is 0.30 meters, and the height is 4 meters; the diameter of the second reaction zone 11b is 1.2 meters, and the height is 7.6 meters.
  • the diameter of the inner delivery pipe 13 is 0.30 meters, and the height is 2.5 meters; the diameter of the third reaction zone 11c is 0.30 meters, and the height is 4 meters.
  • the longitudinal section of the junction of the first reaction zone 11a and the second reaction zone 11b is an isosceles trapezoid, and the angle between the side line of the isosceles trapezoid and the axis is 45°, and the height of the isosceles trapezoid is 0.45 meters; the longitudinal section of the junction of the second reaction zone 11b and the third reaction zone 11c is 0.30 meters, and the height of the isosceles trapezoid is 0.45 meters.
  • the surface is an isosceles trapezoid, the angle between the side line of the isosceles trapezoid and the axis is 45°, and the height of the isosceles trapezoid is 0.45 meters;
  • the distance between the inner delivery hole 13a and the entrance of the third reaction zone 11c is 1 meter, the ratio of the total area of the multiple inner delivery holes 13a to the cross-sectional area of the inner delivery pipe 13 is 0.2:1
  • the distribution internal component 14 forms a cap-shaped distribution internal component
  • the ratio of the area of the end distribution plate 14c to the cross-sectional area of the second reaction zone 11b is 0.8:1
  • the ratio of the sum of the areas of the multiple distribution holes 14a opened on the end distribution plate 14c to the vertical projection area of the end distribution plate 14c is 0.2:1
  • the ratio of the sum of the areas of the distribution holes 14a on the distribution tube 14b to the cross-sectional area of the first reaction zone 11a is 2:1;
  • the total height of the cracking riser 21 is 13 meters, of which the height of the cracking pre-lifting section 21b is 2 meters and the diameter is 0.12 meters, the height of the cracking reaction section 21a is 11 meters and the diameter is 0.15 meters, the height of the cracking stripping section 23a is 3 meters and the diameter is 0.35 meters, the height of the turbulent reaction section 23b is 1.5 meters and the diameter is 0.5 meters, the height of the cracking sedimentation section 23c is 2.5 meters and the diameter is 1.3 meters, and the diameter of the cracking gas phase conveying pipe 24 is 0.2 meters.
  • the cracking pre-lift gas B is injected from the bottom of the cracking pre-lift section 21b and mixed with the regenerated catalyst (catalyst I) directly from the catalyst heat collector 4 or the regeneration unit 3 and moves upward.
  • the light raw material (light hydrocarbon I) atomized by water vapor or directly mechanically atomized enters the cracking riser 21 through the cracking feed unit 8.
  • the oil and gas are output from the top of the multi-stage reaction cyclone separator 12c, and then enter the oil and gas fractionation tower at the bottom of the multi-stage reaction settler 12 through the large oil transfer line, etc.; after the reaction of the heavy raw material, the catalyst containing coke enters the multi-stage reaction cyclone separator 12c from the delivery pipe at the bottom
  • the catalyst to be regenerated enters the multi-stage reaction stripping section 12a at the bottom of the multi-stage reaction settler 12, and the catalyst to be regenerated after being stripped by the multi-stage reaction stripping steam F enters the regeneration unit 3 through the regeneration pipeline 5, and contacts with the regeneration air C entering from the bottom of the regeneration unit 3 to undergo an oxidation reaction.
  • the regenerated flue gas D is output from the top of the regeneration unit 3 and enters the flue gas turbine for further energy recovery and treatment; the high-temperature regenerated catalyst generated by burning the coke enters the catalyst heat exchanger 4 through the regeneration pipeline 6, and converts the low-temperature water vapor G into high-temperature water vapor H after heat extraction.
  • the regenerated catalyst after cooling flows back to the bottom of the first reaction zone 11a and the bottom of the cracking riser 21 through the regeneration pipeline 6 exiting the catalyst heat exchanger 4 for recycling.
  • the catalytic cracking device includes the multistage reaction unit 1 used in Example 1 and a conventional cracking unit, the cracking unit uses a conventional riser, does not contain a cracking settler, and the riser is directly connected to a cyclone separator, and the heavy feedstock oil in Table 1 is processed using the catalyst I in Table 2 to produce ultra-low olefin gasoline.
  • the parameters of the multistage reaction unit 1 are the same as those in Example 1, and the outlet temperature of the riser of the cracking unit is the same as the outlet temperature of the cracking riser 21 in Example 1.
  • the catalytic cracking device includes the multistage reaction unit 1 described in Example 1, without the cracking reaction unit 2, without refining light hydrocarbons, and using the catalyst I in Table 2 to process the heavy crude oil in Table 1 to produce ultra-low olefin gasoline; the various parameters of the multistage reaction unit 1 are the same as those in Example 1.
  • the catalytic cracking device includes the multi-stage reaction unit 1 used in Example 1, without the cracking reaction unit 2, without refining light hydrocarbons, and the diameter of the second reaction zone 11b is 0.9 meters, the second reaction zone 11b is not provided with an internal delivery pipe 13, and the remaining parameters of the multi-stage reaction unit 1 are the same as those of Comparative Example 2.
  • the heavy feedstock oil in Table 1 is processed using the catalyst I to produce a low-olefin gasoline operation mode.
  • the catalytic cracking device includes the multi-stage reaction unit 1 used in Example 1 and a conventional cracking unit, the cracking unit uses a conventional riser, does not contain a cracking settler, and the riser is directly connected to a cyclone separator, and uses the catalyst II in Table 2 to process the heavy feedstock oil in Table 1 to produce more cracked gas.
  • the various parameters of the multi-stage reaction unit 1 are the same as those in Example 2, and the outlet temperature of the riser of the cracking unit is the same as the outlet temperature of the cracking riser 21 in Example 1.
  • the catalytic cracking device includes the multi-stage reaction unit 1 used in Example 1, without the cracking reaction unit 2, without recycling light hydrocarbons, and the diameter of the second reaction zone 11b is 0.9 meters, the internal transport pipe 13 is not set in the second reaction zone 11b, and the remaining parameters of the multi-stage reaction unit 1 are the same as those in Comparative Example 5.
  • Example 1, Comparative Example 1, Comparative Example 2 and Comparative Example 3 are all in the mode of producing ultra-low olefin gasoline.
  • Example 1 and Comparative Example 1 recycle 30% of light hydrocarbons I, half of which come from olefin-rich light gasoline outside the device; Comparative Example 2 and Comparative Example 3 do not recycle gasoline;
  • the gasoline olefins in Example 1 are reduced by 2.6, 3.6 and 6.4 percentage points, respectively, and the propylene yield is increased by 0.5, 0.7 and 1.1 percentage points, respectively;
  • This shows that the device and method provided by the present invention can significantly reduce the olefin content of gasoline and moderately increase the propylene yield.
  • Example 2 Comparative Example 4, Comparative Example 5 and Comparative Example 6 are all operating modes of producing more cracked gas.
  • Example 2 and Comparative Example 4 recycle 30% of light hydrocarbons II, half of which come from olefin-rich light gasoline outside the device; Comparative Example 5 and Comparative Example 6 do not recycle gasoline; Compared with Comparative Example 4, Comparative Example 5 and Comparative Example 6, the gasoline olefins in Example 2 are reduced by 3.3, 11.2 and 13.8 percentage points respectively, and the propylene yield is increased by 1.3, 2.6 and 4.5 percentage points respectively; This shows that the device and method provided by the invention can significantly reduce the olefin content of gasoline and increase the propylene yield.
  • the device and method provided by the present invention can promote the reduction of gasoline olefins in the gasoline pool of the entire refinery, and the cracking gas yield is significantly improved. However, for the entire refinery, the yield or output of cracking gas is significantly improved.

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Abstract

涉及催化领域，提供了一种催化裂解的装置和方法。该装置包括多级反应单元（1）和裂解反应单元（2），多级反应单元包括多变径提升管（11）和多级反应沉降器（12），裂解反应单元（2）包括裂解提升管（21）和裂解沉降器（23）；多变径提升管（11）包括第一反应区（11a）、第二反应区（11b）和第三反应区（11c），第二反应区（11b）的内径大于第一反应区（11a）和第三反应区（11c）的内径，第三反应区（11c）与第二反应区（11b）的连接处设置有向第二反应区（11b）内部延伸的内输送管（13）；裂解沉降器（23）的裂解汽提段（23a）设置有与第二反应区（11b）连通的裂解催化剂出口（25）。该装置能够提高重油与轻烃催化裂解的反应效率，实现重油和轻烃的紧密耦合。

Description

催化裂解的装置和方法
相关申请的交叉引用
本申请要求2023年09月20日提交的中国专利申请202311220149.6的权益，该申请的内容通过引用被合并于本文。
技术领域
本发明涉及催化技术领域，具体地，涉及一种催化裂解的装置和方法。
背景技术
早期催化裂化反应以硅酸铝小球为催化剂，反应器采用密相床层反应器；随着高活性和高选择性的沸石催化剂的出现，石油烃与催化剂的反应速率大大加快，反应器采用快速输送床提升管，有利于提高的液体产品收率和选择性。提升管反应器较密相床反应器在结构和操作方式均有很大的进步，主要表现在进料段油气和催化剂的混合、出口产物的迅速分离、减小提升管截面上温度梯度和减少返回等方面。由于高活性的沸石催化剂大大提高了反应强度，自上世纪60-70年代以来，多家公司开发了各自快速输送床反应器，即提升管反应器，由于具有较好弹性和灵活性，强调重油的一次裂化，轻质油收率和选择性均较为理想，此类型的反应器至今仍有较强的生命力。自上世纪末，国III汽油标准开始限制汽油烯烃含量，CN1237477A、CN1245202A、CN1232069A等提出了变径提升管，在提升管中部增加一段扩径，从而形成流化床反应段，进而可强化已生成汽油的二次反应，研究发现此类反应器可在重油一次裂化的同时兼顾汽油二次转化。
初始的变径提升管类技术没有考虑提升管段与流化床段之间需要内构件以调整气固流化状态，由于工业装置的尺寸是米级，若没有内构件，则由于下部提升管进入扩径段的高速流体射流等因素，很难在扩径段下部形成气固流体的稳定流化进而形成稳定的固体颗粒密相区。虽然在变径提升管的工业实施过程中，已先后推出了拱形(凸式)分布板、异型(盆式)分布板、凹式分布板以及蘑菇头式分布器，可使提升管出来的高速油气及催化剂形成充分均匀的稳态流化，部分催化剂形成稳定的流化床；但是，提升管的扩径段与上部提升管一般通过缩径圆 台形外筒相连，没有在此段安装内构件调整气固的流化。
为了提高汽油品质、降低汽油烯烃含量和提高裂化气产率，国内外公开了大量用于轻烃回炼的技术措施，典型的有两种方式：第一种是将汽油或其他轻烃回炼至原重油提升管，一般是注入原料油喷嘴的上游，例如US5043522A和US5846403A公开了将催化裂化汽油注入原料油进料喷嘴上游，利用高温、高活性的再生催化剂，对其进行催化转化；CN1160746A也披露了一种将直馏汽油、焦化汽油等低品质汽油注入提升管反应下部，使其优先与再生催化剂接触的方法；第二种是采用单独的汽油或轻烃提升管回炼，例如CN1069054A和US3784463A均采用带有双提升管反应器的进行反应，将包括催化裂化粗汽油在内的低品质汽油注入汽油提升管反应器，利用高温、大剂油比的反应条件实现对低品质汽油的催化改质，以提高液化气产率和汽油辛烷值；轻烃进料与催化剂接触后所生成的分子相对较小，体积膨胀明显，对催化裂化提升管反应器上部重油进料及重油与催化剂的接触产生影响。其中，第一种方式为保持重油反应不产生负面影响，轻烃回炼的比例较低；第二种方式采用在传统的单提升管反应器基础上，另外增设单独的轻烃提升管，可使轻烃单独进行二次裂解反应，实现不同原料的分器单独优化反应，但传统提升管对反应深度控制不理想，如通过提高提升管出口温度以达到较高转化深度，则易发生过裂化反应；如提升管出口反应温度低则达不到理想的转化深度。
此外，“催化裂解DCC双提升管工艺应用浅析”报道了催化裂解DCC双提升管工艺(即DCC plus)的装置特点，重油反应器包括提升管和湍流床，轻烃提升管末端的油气及催化剂注入主沉降器；“重油MIP与劣质催化裂化柴油LTAG组合工艺催化裂化装置运行分析”公开了重油MIP反应器及劣质催化裂化柴油LTAG反应器组合的双反应器工艺。但是，仍存在重质原料和轻烃裂化反应的效果不够理想，产品性质及产品分布不易灵活调控等缺陷。
发明内容
本发明的目的是提供一种催化裂解的装置和方法，该装置和方法能够强化重油裂解过程与轻烃裂解过程中的气固混合与接触，提高催化反应效率，可实现重油反应器和轻烃反应器的紧密耦合，提高调控产品性质及产品分布的操作灵活 性。
本发明的发明人在催化裂解的研究中，意外地发现，将采用扩径提升管的重油反应器与轻烃裂解反应器并联耦合后，从裂解催化剂出口将轻烃裂解后的催化剂回用至重油反应器的扩径段，不仅可将重油裂解过程与轻烃裂解过程独立进行，提高轻烃的回炼比例，而且有助于强化重油的催化反应；通过在重油反应器的扩径段内部布置内输送管，能够改变并优化扩径段的气固流化状态，拓宽扩径段内气体线速和催化剂床层密度的调整范围；而在轻烃裂解反应器的裂解沉降器内设置湍流反应段，也有利于轻烃在该反应段与裂解催化剂的充分强化接触，能够实现在较低温度下、低空速下轻烃与催化剂的高效接触反应，避免反应温度过高导致轻烃过裂化反应；此时，重油反应器与轻烃裂解反应器的耦合，可同时对重油和轻烃进行单独且高效的裂化/裂解操作控制，具有更多种的操作手段，能够实现调控产品性质及产品分布的各种灵活操作。
为了实现上述目的，本发明第一方面提供一种催化裂解的装置，该装置包括多级反应单元和裂解反应单元，所述多级反应单元包括多变径提升管和多级反应沉降器，所述裂解反应单元包括裂解提升管和裂解沉降器；所述多变径提升管包括沿其长度方向依次串联的第一反应区、第二反应区和第三反应区，所述第三反应区的出口与所述多级反应沉降器连通，所述第二反应区的内径大于所述第一反应区和第三反应区的内径，所述第三反应区与所述第二反应区的连接处设置有向所述第二反应区内部延伸的内输送管；所述裂解沉降器包括沿其长度方向依次设置的裂解汽提段、湍流反应段和裂解沉降段，所述裂解提升管的出口位于所述湍流反应段内，所述裂解汽提段设置有与所述第二反应区连通的裂解催化剂出口。
优选地，所述内输送管与所述第三反应区的内径之比为0.9-1.2:1，所述内输送管与所述第二反应区的长度之比为0.15-0.8:1。
更优选地，所述第一反应区、第二反应区和第三反应区的内径之比为1:2-5:0.7-1.5、长度之比为1:0.3-2:0.5-2；所述湍流反应段、裂解汽提段和裂解沉降段的内径之比为1:0.4-0.8:2.5-5、长度之比为1:1.5-3:1.5-4；所述多变径提升管与裂解提升管的长度之比为1:0.5-1.2。
进一步优选地，所述内输送管的侧壁上设置有至少一个内输送孔，所述内输送孔与所述第三反应区的入口之间的距离为0.5-2m。
具体地，至少一个所述内输送孔的总面积与所述内输送管的截面积之比为0.05-0.3:1。
作为一种优选的实施方式，所述第二反应区内设置有与所述第一反应区连接的分布内构件，所述分布内构件上设有多个分布孔。
优选地，所述分布内构件包括与所述第一反应区连接的分布筒和与所述分布筒连接的端部分布板，所述分布筒和端部分布板上均设置有所述分布孔，所述端部分布板位于所述分布筒远离所述第一反应区的一端。
更优选地，所述分布筒上分布孔的面积之和与所述第一反应区的截面积之比为0.6-4:1；所述端部分布板上分布孔的面积之和与该端部分布板的垂直投影面积之比为0.1-0.3:1。
更优选地，所述分布筒上的分布孔处设置有向该分布筒的外侧延伸的侧通道，所述侧通道远离所述分布筒的开口断面为朝向所述第一反应区的斜面。
进一步优选地，所述端部分布板的面积与所述第二反应区的截面积之比为0.7-0.85:1。
具体地，所述第一反应区、第二反应区、第三反应区和内输送管为同轴设置，所述裂解催化剂出口位于所述裂解汽提段的底部且与所述第二反应区的底部连通。
更具体地，所述多变径提升管还包括与所述第一反应区的入口连接的多级反应预提升段，所述裂解提升管包括沿其长度方向依次串联的裂解预提升段和裂解反应段，所述裂解反应段的出口位于所述湍流反应段靠近所述裂解汽提段的一端，所述多级反应预提升段和所述第一反应区之间的过渡段与多级反应进料单元连通，所述裂解预提升段与所述裂解反应段之间的过渡段与裂解进料单元连通。
进一步具体地，所述裂解反应段的出口处设置有裂解分布器，所述裂解沉降段的气相出口与所述多级反应沉降器连通。
典型地，该装置还包括再生单元，所述多级反应沉降器上设置有与所述再生单元的入料口连通的多级反应催化剂出口，所述再生单元的再生催化剂出口与所述多变径提升管和/或所述裂解提升管连通。
本发明第二方面提供一种催化裂解的方法，采用上述的装置，该方法包括以下步骤：
重质原料与多级反应催化剂混合形成的多级反应混合料从所述第一反应区的入口输入，依次进入所述第一反应区、所述第二反应区，再从所述内输送管进入所述第三反应区，以进行多级催化反应得到多级反应物，所述多级反应物经所述多级反应沉降器进行多级反应分离得到多级反应油气产物和多级反应待生催化剂；
轻质原料与裂解催化剂混合形成的裂解混合料从所述裂解提升管的入口输入，经所述裂解提升管进行裂解催化反应得到裂解反应物，所述裂解反应物进入所述湍流反应段进行湍流反应后，经所述裂解沉降段进行裂解分离、经所述裂解汽提段进行裂解汽提得到裂解油气产物和裂解待生催化剂；
将所述裂解待生催化剂从所述裂解催化剂出口输入所述第二反应区内，与所述多级反应混合料混合。
优选地，所述内输送管的入口温度为510-540℃，所述裂解提升管的出口温度为540-620℃。
优选地，所述内输送管的入口温度为550-640℃，所述裂解提升管的出口温度为560-660℃。
更优选地，所述多级反应混合料在所述第一反应区的停留时间为0.8-2s，在所述第二反应区的停留时间为5-15s，在所述第三反应区的停留时间为0.5-1.5s；所述裂解混合料在所述裂解提升管的停留时间为0.8-3.5s，在所述湍流反应段的停留时间为3-13s。
通过上述技术方案，本发明提供的催化裂解的装置通过并列设置的多级反应单元和裂解反应单元，可将重油裂解过程与轻烃裂解过程独立进行，提高轻烃的回炼比例，并从裂解催化剂出口将轻烃裂解后的催化剂回用至第二反应区，有助于提高第二反应区内催化剂的流量及活性，强化重油的催化反应，实现重油裂解过程与轻烃裂解过程的紧密耦合；通过在多级反应单元的多变径提升管内部布置内输送管，能够改变并优化第二反应区的气固流化状态，拓宽第二反应区内气体线速和催化剂床层密度的调整范围，配合以在裂解反应单元的裂解沉降器内设置湍流反应段，基于湍流反应段的裂解催化剂密度较大，有利于轻烃在该反应段与裂解催化剂的充分强化接触，能够实现在较低温度下、低空速下轻烃与催化剂的高效接触反应，避免反应温度过高导致轻烃过裂化反应；本发明提供的装置中 两个反应单元的耦合，可同时对重油和轻烃进行单独裂化/裂解操作控制，具有更多种的操作手段，实现调控产品性质及产品分布的各种灵活操作。
有关本发明的其他优点以及优选实施方式的技术效果，将在下文的具体实施方式中进一步说明。
附图说明
构成本发明的一部分的附图用来提供对本发明的进一步理解，本发明的示意性实施方式及其说明用于解释本发明，并不构成对本发明的不当限定。在附图中：
图1是本发明中催化裂解的装置的一种具体实施方式的结构示意图；
图2是本发明中内输送管和分布内构件的一种具体实施方式的结构示意图。
附图标记说明
1-多级反应单元，11-多变径提升管，11a-第一反应区，11b-第二反应区，
11c-第三反应区，11d-多级反应预提升段，12-多级反应沉降器，12a-多级反应汽提段，12b-多级反应沉降段，12c-多级反应旋风分离器，13-内输送管，13a-内输送孔，14-分布内构件，14a-分布孔，14b-分布筒，14c-端部分布板，14d-侧通道，15-多级反应催化剂出口；
2-裂解反应单元，21-裂解提升管，21a-裂解预提升段，21b-裂解反应段，
22-裂解分布器，23-裂解沉降器，23a-裂解汽提段，23b-湍流反应段，23c-裂解沉降段，23d-裂解旋风分离器，24-裂解气相输送管，25-裂解催化剂出口；
3-再生单元，31-再生催化剂出口，32-再生旋风分离器，33-再生空气入口，
34-再生烟气出口；
4-催化剂取热器；5-待生管路；6-再生管路；7-多级反应进料单元；8-裂解
进料单元；9-急冷油气单元；10-汽提蒸汽单元；
A-多级反应预提升气；B-裂解预提升气；C-再生空气；D-再生烟气；E-反
应后油气；F-多级反应汽提蒸汽；G-低温水蒸气；H-高温水蒸气。
具体实施方式
以下结合附图对本发明的具体实施方式进行详细说明。应当理解的是，此 处所描述的具体实施方式仅用于说明和解释本发明，并不用于限制本发明。
在本发明的描述中，需要说明的是，除非另有明确的规定和限定，术语“连接”、“相连”应做广义理解，例如，可以是固定连接，也可以是可拆卸连接，或者是一体连接；可以是直接连接，也可以是通过中间媒介间接连接，可以是抵接，可以是两个元件内部的连通或两个元件的相互作用关系。对于本领域的普通技术人员而言，可以根据具体情况理解上述术语在本发明中的具体含义。
需要理解的是，术语“上”、“下”、“前”、“后”、“内”、“外”等指示的方位或位置关系为基于附图所示的方向或位置关系，仅是为了便于描述本发明和简化描述，而不是指示或暗示所指的装置或元件必须具有特定的方位、以特定的方位构造和操作，因此不能理解为对本发明的限制。此外，术语“第一”、“第二”仅用于描述的目的，而不能理解为指示或暗示相对重要性或隐含指明所指示的技术特征的数量，因此，限定有“第一”、“第二”的特征可以明示或隐含地包括一个或更多个所述特征。
在本发明的一个基本实施方式中，参见图1，催化裂解的装置包括多级反应单元1和裂解反应单元2，多级反应单元1包括多变径提升管11和多级反应沉降器12，裂解反应单元2包括裂解提升管21和裂解沉降器23；多变径提升管11包括沿其长度方向依次串联的第一反应区11a、第二反应区11b和第三反应区11c，第三反应区11c的出口与多级反应沉降器12连通，第二反应区11b的内径大于第一反应区11a和第三反应区11c的内径，第三反应区11c与第二反应区11b的连接处设置有向第二反应区11b内部延伸的内输送管13；裂解沉降器23包括沿其长度方向依次设置的裂解汽提段23a、湍流反应段23b和裂解沉降段23c，裂解提升管21的出口位于湍流反应段23b内，裂解汽提段23a设置有与第二反应区11b连通的裂解催化剂出口25。
本发明提供的催化裂解的装置可以应用于多种化工生产领域的催化裂解或者催化裂化过程，尤其适用于重油和轻烃的催化裂化/催化裂解，例如，生产超低烯烃汽油、多产裂化气；其反应单元1用作重质原料(重油)进行催化裂化/催化裂解，裂解反应单元2用作轻质原料(轻烃)进行催化裂化/催化裂解。需要强调的是，本申请中催化裂解的装置基于多级反应单元1和裂解反应单元2的并联设置，通过更换催化剂及调控操作条件可以实现超低烯烃汽油和多产裂化 气两种操作模式的切换，还能够增大两种模式的调整深度和变化幅度，实现相对于重油的较大轻烃比例的加工量；其中，轻烃可以来自该装置中多级反应单元1以外乃至使用车间或工厂以外。另外，多级反应单元1和裂解反应单元2并联的同时，从裂解催化剂出口25将轻烃裂解后的催化剂回用至第二反应区11b，可充分利用轻烃反应过后的高温待生剂具有较高裂解活性，将其补充至多级反应单元1的第二反应区11b，有效提升了其内部的催化剂浓度和催化剂活性，实现多级反应单元1和裂解反应单元2的紧密耦合。
发明人在研究过程中发现，常规的MIP-CGP反应器的扩径段内表观气速过低，则不足以将催化剂带走至第三反应区11c，易发生噎塞乃至催化剂在扩径段内积聚而造成装置事故；如果表观气速过高，则扩径段内不能形成催化剂密相段，不能满足工艺要求。通过在多变径提升管11中内径较大的第二反应区11b(即扩径段)内设置内输送管13，可有效克服上述缺陷，产生以下几个方面的效果：第一，当第二反应区11b内表观气速降低形成了湍流床，当湍流床的上界面淹没内输送管13的底部端口时，则可将湍流床顶部的催化剂吸走，即可控制第二反应区11b内湍流床的高度。第二，在第二反应区11b上部与内输送管13同轴等高的扩径段上部即可作为湍流床的稀相段，气体携带大量的颗粒固体抛撒至湍流床上方的稀相段，由于气体流动速度变小，形成固体颗粒沉降的自由空间，湍流床夹带上来的催化剂颗粒，会再次落入第二反应区11b下部，强化催化剂的返混；由于自由空间内储存了相当数量的气体，可为多变径提升管11内的气体脉动提供缓冲空间；以上两方面的效果，可在一定程度上避免多变径提升管11内催化剂的大量急剧和气体压力快速变大和急剧波动，拓宽了第二反应区11b内最佳操作的气体线速范围。第三，第二反应区11b的气固流化呈湍流床，在湍动床的底部，固体相湍动剧烈且呈连续相，单位空间内固气比大，固体颗粒在湍动床内气泡直径较小，气固接触效率加强，一般气固接触效率高于各种形式的输送床，适用一些需要气固接触效率进一步加强的领域。第四，若采用大于2m/s的气速时，由于与内输送管13等高的扩径段内气速较小，可以使被下方高速气体抛撒上来的催化剂降落在底部的第二反应区11b，提高了第二反应区11b底部的催化剂密度，强化气固混合和接触。
与常规的MIP-CGP反应器相比，本发明提供的内输送管13在一定程度上 改变了第二反应区11b内气体与催化剂的流动状态。第二反应区11b内的气体表观速度范围更宽，常规MIP-CGP种扩径段的表观气速在1.4-2.8m/s，工业运行装置一般优选在1.5-2.0m/s；本发明提供的多级反应单元1中，第二反应区11b内表观气速在0.7-3.5m/s均能在第二反应区11b的底部形成密相段，表观气速优选的操作范围为0.9-2.5m/s；与之对应的是本发明中第二反应区11b与第一反应区11a的内径之比更大，优选情况下，第二反应区11b与第一反应区11a的内径之比为2-5:1，更优选为1.7-3.5:1。
本发明中，催化裂解的装置在使用时，多级反应单元1和裂解反应单元2通常为立式设置，多变径提升管11的长度方向为竖直方向，第一反应区11a、第二反应区11b和第三反应区11c由下至上依次连接，裂解沉降器23的长度方向也为竖直方向，裂解汽提段23a、湍流反应段23b和裂解沉降段23c由下至上依次连接；多级反应沉降器12与多变径提升管11可以同轴布置或者并列布置。
本发明中，多级反应沉降器12内一般包括由下至上设置的多级反应汽提段12a和多级反应沉降段12b，多级反应沉降段12b内设置有一级或多级反应旋风分离器12c，以使得多变径提升管11的物料可进入多级反应沉降段12b内经多级反应旋风分离器12c进行分离，使得分离出的待生催化剂进入多级反应汽提段12a经多级反应汽提蒸汽F进行汽提再排出，可进入催化剂再生器中进行空气烧焦再生处理后返回多级反应单元1进行循环利用。
上述提供的催化裂解的装置的基本实施方式，应用于催化裂解时，包括以下步骤：
将重质原料经蒸汽雾化后在第一反应区11a底部与多级反应催化剂混合形成多级反应混合料，多级反应混合料从第一反应区11a的入口输入，依次进入第一反应区11a、第二反应区11b上行，第二反应区11b的扩径可使多级反应混合料停留时间变长并可充分混合接触，充分完成重油及其中间馏分的二次反应，再从内输送管13进入第三反应区11c，以进行多级催化反应得到多级反应物，由于内输送管13与第三反应区11c的提升输送管缩径以提高该混合料的流速，使该混合料快速进入多级反应沉降器12进行多级反应分离，得到多级反应油气产物和多级反应待生催化剂，反应后带炭的待生多级反应催化剂经过或不经水蒸汽汽提后排出；
轻质原料与裂解催化剂混合形成的裂解混合料从裂解提升管21的入口输入，经裂解提升管21进行裂解催化反应得到裂解反应物，裂解反应物进入湍流反应段23b进行湍流反应后，经裂解沉降段23c进行裂解分离、经裂解汽提段23a进行裂解汽提得到裂解油气产物和裂解待生催化剂；
将裂解待生催化剂从裂解催化剂出口25输入第二反应区11b内，与多级反应混合料混合，提高第二反应区11b内的多级反应催化剂活性的和密度。
本发明中，所述重质原料为石油烃和/或其它矿物油，石油烃选自减压瓦斯油(VGO)、常压瓦斯油(AGO)、焦化瓦斯油(CGO)、脱沥青油(DAO)、减压渣油(VR)、常压渣油(AR)和加氢重油中的至少一种；其它矿物油选自煤液化油、油砂油和页岩油中的至少一种。优选地，重质原料选自减压瓦斯油、常压瓦斯油、焦化瓦斯油、脱沥青油、减压渣油、常压渣油和加氢重油中的至少一种；其中，VGO、AGO、CGO、DAO、VR、AR为未加氢的全馏分或部分馏分，或为加氢后的全馏分或部分馏分。
本发明中，所述轻质原料选自富含烯烃汽油、柴油、碳原子数为4-8的不富含/富含烯烃的轻烃中的至少一种，其中，富含烯烃汽油可以为采用本发明提供的装置进行催化裂解获得的汽油、常规催化裂化汽油、其他催化裂解汽油、焦化汽油、热裂解汽油、热裂化汽油中一种或其中一种以上的混合物；所述不富含烯烃的轻质可以选自直馏石脑油、直馏汽油、加氢石脑油、碳原子数为4-8的烷烃、抽余油中一种或其中一种以上的混合物。
本发明中，所述方法中多级反应催化剂和裂解催化剂可以采用所有类型的催化裂化或催化裂解催化剂，其活性组分选自含有或不含稀土的Y或HY型沸石、含有或不含稀土的超稳Y型沸石、ZSM-5系列沸石或用其它方法制得的具有五元环结构的高硅沸石中的一种或多种的催化剂，以及无定型硅铝催化剂。其中，多级反应催化剂和裂解催化剂可以采用相同的催化剂，也可以采用不同的催化剂。
上述装置的使用过程中，为了控制多级反应单元1的反应温度，可以在多级反应单元1的第一反应区11a上部与第二反应区11b中部的任意位置注入急冷剂，优选地，在第一反应区11a的顶部至第二反应区11b的底部设置有急冷油气单元9。急冷剂可以选自水、回炼油、汽油、柴油和污油中的至少一种，注入量 为催化原料量(重质原料)的1-10wt％，急冷剂的温度在10-200℃，优选为10-30℃。
本发明中，裂解反应单元2中大部分催化剂落入湍流反应段23b，从而湍流反应段23b及裂解汽提段23a的催化剂密度较大，通过控制裂解沉降器23底部外输催化剂而使裂解沉降器23中部和顶部的催化剂逐步落入底部的裂解汽提段23a，落入裂解汽提段23a的催化剂可经过或不经过水蒸汽汽提后从裂解催化剂出口25输出。裂解催化剂出口25通过管路与第一反应区11a末端至第二反应区11b中部的任意位置连通，以将裂解待生催化剂输入多变径提升管11的第二反应区11b，另外，还可以增设管路将裂解催化剂出口25与催化剂再生器连接进行回收。优选情况下，裂解催化剂出口25位于裂解汽提段23a的底部且通过管路与第二反应区11b的底部(靠近第一反应区11a的一端)连通，该管路上可以设置有滑阀，以通过控制该滑阀的开度来调节输入第二反应区11b的裂解待生催化剂量，也可以调节裂解沉降器23内催化剂存量。
由于裂解催化剂与轻烃反应后焦炭含量低、温度较高，仍具有较高的催化活性，在输入第二反应区11b内，提高其内催化剂的流量及其催化剂活性，有利于强化多变径提升管11内的催化反应；经裂解沉降器23分离后的轻烃反应后油气不进入重油反应的多变径提升管11的各个反应区，不影响第二反应区11b的油气流量及流速，进而不影响第二反应区11b的气固流态化，同时缩短轻烃在裂解反应单元2内高温反应区的停留时间进而降低热裂化反应的程度。
本发明中，内输送管13的内径小于第二反应区11b，优选情况下，内输送管13与第三反应区11c的内径之比为0.9-1.2:1，更优选为，内输送管13与第三反应区11c的内径之比为0.95-1.05:1，使得两者的内径相当，以进一步拓宽第二反应区11b内气体表观线速的调节范围，使多级反应单元1能承受的反应油气流速(代表装置加工能力)在更大的范围内调节或波动。
本发明中，内输送管13的长度小于第二反应区11b的长度，优选情况下，内输送管13与第二反应区11b的长度之比为0.15-0.8:1，更优选为0.2-0.6:1，能够更好地使得第二反应区11b内的气固流化状态及湍流效果更佳。需要说明的是，第二反应区11b可以为圆柱形，也可为从下到上逐渐扩径的圆台形，也可为两级或三级扩径的圆柱形；第二反应区11b的高度指的是其本身的区域段、其与第一反应区11a之间的过渡段、与第三反应区11c之间的过渡段的高度总和。
本发明中，作为内输送管13的一种优选实施方式，内输送管13底部端口与第二反应区11b底部(即第一反应区11a的出口)的距离为2-8m，优选为3-6m；内输送管13的长度为2-7m，优选为3-5m；此时，能够更好地控制控制第二反应区11b内湍流床的高度。
本发明的发明人通过对变径提升管特别是扩径段催化剂颗粒的研究发现，从轴向分布看，扩径段颗粒轴向密度成“S形”上稀下浓分布，浓相段高度虽然受气速和颗粒输送量的影响，但是在浓相段集中于扩径段高度60％以下的底部段。从催化剂颗粒的径向分布看，在反应器任一截面出，催化剂颗粒速度最大值在提升管中心，颗粒速度在中心区维持一个较大且相对恒定的数值，从中心向管壁的下降程度较为缓慢，在r/R(r代表离中心轴的距离，R为管道半径)到达一定数值后颗粒速度突然降低，在边壁附近几乎为零。边壁效应存在于整个变径提升管，在下部的气力输送提升管，气体速度在10m/s以上，r/R<0.8的区域内催化剂颗粒轴向速度均较高，但在扩径段，由于气体表观速度在1.5-3m/s，边壁效益更为明显；在扩径段的下部，r/R<0.7的区域内催化剂颗粒速度均较高，但靠近边壁的催化剂颗粒轴向速度为零乃至为负；在扩径段的上部，r/R<0.6的区域内催化剂颗粒速度均较高，但靠近边壁的催化剂颗粒速度维持在0附近，即此段有很大部分催化剂在悬浮，此部分催化剂可能在扩径段停留可能过长，影响最佳工艺效果。基于此，优选情况下，内输送管13的侧壁上设置有至少一个内输送孔13a，内输送孔13a与第三反应区11c的入口之间的距离为0.5-2m。内输送孔13a的设置能够使得扩径段第二反应区11b内上部与内输送管13同轴部分内气体处于流动状态，以防止在第二反应区11b上部出现没有气体流动的死区，以进一步提高第二反应区11b内的湍流效果和催化反应效率。
本发明中，内输送孔13a优选为设置多个，可以周向均匀地分布在内输送管13上。进一步优选地，至少一个内输送孔13a的总面积与内输送管13的截面积之比为0.05-0.3:1。其中，内输送管13的截面积指的是与该内输送管13的长度方向相垂直的截面的面积。
本发明中，第一反应区11a与第二反应区11b之间的过渡段、第二反应区11b与第三反应区11c之间的过渡段为穹形或圆台形，且与整个多级反应单元1同轴布置。优选情况下，第一反应区11a、第二反应区11b、第三反应区11c和 内输送管13为同轴设置。
作为一种优选实施方式，第一反应区11a、第二反应区11b和第三反应区11c的内径之比为1:2-5:0.7-1.5、长度之比为1:0.3-2:0.5-2。此时，更加有利于延长油剂在第二反应区11b的停留时间并可充分混合接触，充分完成重油及其中间馏分的二次反应。
本发明中，裂解沉降器23呈两级扩径结构，裂解汽提段23a、湍流反应段23b和裂解沉降段23c的筒体外径依次增大，裂解提升管21位于湍流反应段23b内或者湍流反应段23b与裂解沉降段23c之间的过渡段。作为裂解分离器23的一种优选实施方式，湍流反应段23b、裂解汽提段23a和裂解沉降段23c的内径之比为1:0.4-0.8:2.5-5、长度之比为1:1.5-3:1.5-4。此时，更加有利于湍流反应段23b中轻烃与催化剂的高效充分反应，实现在较低温度下、低空速下轻烃与催化剂的高效接触反应。
本发明中，裂解反应单元2的裂解提升管21与裂解沉降器23优选为同轴设置。
作为本发明一种优选的实施方式，多变径提升管11与裂解提升管21的长度之比为1:0.5-1.2，更优选为1:0.6-1。多变径提升管11中第一反应区11a的内径与裂解提升管21的内径之比为1:0.2-1。此时，能够更好地实现多级反应单元1与裂解反应单元2之间的耦合效果。一般情况下，实验室的中型装置中多变径提升管11的高度约为20米；工业生产中变径提升管11的高度约为40米。
作为本发明另一种优选的实施方式，参见图2，第二反应区11b内设置有与第一反应区11a连接的分布内构件14，分布内构件14上设有多个分布孔14a，也就是，分布内构件14为开孔分布件。分布内构件14的主体位于第一反应区11a与第二反应区11b之间的过渡段，其作用是将第一反应区11a高速气固颗粒迅速降至第二反应区11b所需的气固流化状态。
本发明中，分布内构件14可以采用开口分布板，其开口率为8-35％；也可以采用顶部为蘑菇头、圆锥型或伞帽型并从侧通道输出油气的帽头型分布器，帽头型分布器可以采用常规的伞帽型分布器、蘑菇头型分布器或者圆锥形帽头型分布器。具体地，分布内构件14包括与第一反应区11a连接的分布筒14b和与分布筒14b连接的端部分布板14c，分布筒14b和端部分布板14c上均设置有分布 孔14a，端部分布板14c位于分布筒14b远离第一反应区11a的一端。其中，端部分布板14c可以呈伞帽型者穹顶型，例如，采用拱形(凸式)分布板、异型(盆式)分布板、凹式分布板。
作为端部分布板14c的一种优选实施方式，端部分布板14c上开设的多个分布孔14a的面积之和(即总孔面积)与该端部分布板14c的垂直投影面积之比为0.1-0.3:1，可使端部分布板14c上方的气体平缓流动，并防止在端部分布板14c上堆积催化剂。其中，端部分布板14c的垂直投影面积指的是水平平行光线照射到端部分布板14c后，端部分布板14c的外轮廓在垂直平面上的影子所占的面积。
本发明中，分布筒14b上的分布孔14a可以为长方形或椭圆形，优选情况下，分布筒14b上分布孔14a的面积之和与第一反应区11a的截面积之比为0.6-4:1；能够更好地提升对第一反应区11a内高速气固颗粒的减速、分布效果。其中，第一反应区11a的截面积指的是与该第一反应区11a的长度方向相垂直的截面的面积。
作为分布筒14b一种优选的实施方式，分布筒14b上的分布孔14a处设置有向该分布筒14b的外侧延伸的侧通道14d，侧通道14d远离分布筒14b的开口断面为朝向第一反应区11a的斜面。侧通道14d的横截面(与侧通道14d长度方向垂直的截面)形状与分布筒14b上的分布孔14a相对应，可以为长方形或椭圆形。侧通道14d远离分布筒14b的开口断面一般不与该侧通道14d的轴线垂直，该开口断面为朝向第一反应区11a的斜面，即，其断面延长线与多级反应单元1的轴线呈一定斜向上的夹角；该开口断面的上部为该侧通道14d的一部分，以尽可能地避免侧通道14d的出口被催化剂堵住，保证侧通道14d出口的通畅性。
作为分布筒14b另一种优选的实施方式，端部分布板14c的面积与第二反应区11b的截面积之比为0.7-0.85:1，使得端部分布板14c边缘的气速在2.5m/s以上，最好在3.5m/s以上，保证在第二反应区11b底部乃至过渡段不发生催化剂堆积而造成装置操作事故。其中，端部分布板14c的面积为端部分布板14c外轮廓的投影面积，第二反应区11b的截面积为与该第二反应区11b的长度方向相垂直的截面的面积。
本发明中，内输送孔13a和/或分布孔14a内安装或涂抹有耐磨衬里，例如陶瓷耐磨衬里。
本发明中，为了使得原料与催化剂充分混合，优选情况下，多变径提升管11还包括与第一反应区11a的入口连接的多级反应预提升段11d，裂解提升管21包括沿其长度方向依次串联的裂解预提升段21b和裂解反应段21a，裂解反应段21a的出口位于湍流反应段23b靠近裂解汽提段23a的一端，多级反应预提升段11d和第一反应区11a之间的过渡段与多级反应进料单元7连通，裂解预提升段21b与裂解反应段21a之间的过渡段与裂解进料单元8连通。通过多级反应预提升段11d注入多级反应预提升气A，与多级反应催化剂混合后上行，经水蒸气雾化的重质原料从多级反应进料单元7进入多变径提升管11，使得重质原料与多级反应催化剂在第一反应区11a内接触反应；通过裂解预提升段21b注入裂解预提升气B，与裂解催化剂混合后上行，经水蒸气雾化的轻质原料从裂解进料单元8进入裂解提升管21，使得轻质原料与裂解催化剂在裂解反应段21a内接触反应。
本发明中，为了提高轻烃反应混合物进入湍流反应段23b后催化剂与轻烃的接触效果，提高催化反应效率，优选情况下，裂解反应段21a的出口处设置有裂解分布器22。
本发明中，裂解沉降段23c的气相出口可以汇入主油气管道，也可以与多级反应沉降器12连通，例如，多级反应沉降器12顶部设有集气室，裂解沉降段23c的气相出口通过与裂解气相输送管24与该集气室连通，使得反应后油气E(包含重质原料反应后油气和轻质原料反应后油气)从集气室排出。裂解汽提段23a的底部可以设置汽提蒸汽单元10，以通入蒸汽对裂解汽提段23a内的催化剂进行汽提处理。
本发明提供的催化裂解的装置在使用时，可以多级反应单元1和裂解反应单元分别配置相应的催化剂再生器，在多级反应催化剂和裂解催化剂相同时，也可以两者共用一个催化剂再生器。作为本发明的一种优选实施方式，该装置还包括再生单元3，多级反应沉降器12上设置有与再生单元3的入料口连通的多级反应催化剂出口15，再生单元3的再生催化剂出口31与多变径提升管11和/或裂解提升管21连通。其中，裂解催化剂出口25不仅与第二反应区11b连通，也与再生单元3的入料口连通，以将裂解反应单元2的部分待生催化剂直接输入再生单元3中。
本发明中，再生单元3内设置有再生旋风分离器32，且底部设有再生空气入口33、顶部设有再生烟气出口34，待生催化剂经待生管路5进入再生单元3，与再生空气入口33进入的再生空气C接触发生氧化反应，再生烟气D从再生烟气出口34输出进入烟气轮机等进一步回收能量和处理；烧掉焦炭从而生成的高温再生催化剂经再生管路6进入催化剂取热器4，将低温水蒸气G从取热蒸汽入口进入催化剂取热器4与再生催化剂取热后变为高温水蒸气H，从取热蒸汽出口输出，经冷却后的再生催化剂经再生管路6回流至第一反应区11a底部和裂解提升管21底部循环使用。
本发明中，部分或全部的多级反应催化剂、部分或全部的裂解催化剂采用再生催化剂出口31输出的再生催化剂。
基于上述提供的催化裂解的装置，本发明第二方面提供该装置应用于催化裂解的方法，该方法包括以下步骤：
重质原料与多级反应催化剂混合形成的多级反应混合料从第一反应区11a的入口输入，依次进入第一反应区11a、第二反应区11b，再从内输送管13进入第三反应区11c，以进行多级催化反应得到多级反应物，多级反应物经多级反应沉降器12进行多级反应分离得到多级反应油气产物和多级反应待生催化剂；
轻质原料与裂解催化剂混合形成的裂解混合料从裂解提升管21的入口输入，经裂解提升管21进行裂解催化反应得到裂解反应物，裂解反应物进入湍流反应段23b进行湍流反应后，经裂解沉降段23c进行裂解分离、经裂解汽提段23a进行裂解汽提得到裂解油气产物和裂解待生催化剂；
将裂解待生催化剂从裂解催化剂出口25输入第二反应区11b内，与多级反应混合料混合。
本发明提供的装置及方法具有多种操作模式，例如，超低烯烃汽油操作模式下，多级反应单元1采用多产汽油及具有降烯烃功能催化剂，控制催化剂循环量、急冷油气单元9的急冷剂注入量、在第二反应区11b底部或中部不补充或补适当量的再生剂或经冷却后的再生剂(温度一般在560-660℃)或含碳量<0.5％的含碳待生催化剂等措施，控制第二反应区11b内温度在510-540℃，优化范围为515-530℃，内输送管13的入口温度控制在510-540℃，优化范围为515-530℃，可实现生产超低烯烃汽油模式；
同时，裂解反应单元2裂解各种富含烯烃的催化轻汽油、焦化汽油等，可进一步降低全厂范围的高烯烃汽油量，降低全厂汽油池的烯烃含量，裂解提升管21的出口温度控制540-620℃，优选560-600℃。
又例如，多产裂化气模式下，多级反应单元1采用多产裂化气的催化裂解催化剂，在第二反应区11b的底部或中部不补充或补充适当量的再生剂或经冷却后的再生剂(温度一般在580-680℃)或含碳量<0.5％的含碳待生催化剂(温度一般在580-680℃)，控制第二反应区11b内温度在560-660℃，优化范围为560-600℃，内输送管13的入口温度控制在550-640℃，优化范围为560-600℃，可实现多产裂化气操作模式；
同时，裂解反应单元2回炼富含C4-C8烯烃的轻烃，或汽油或柴油等，裂解提升管21的出口温度控制560-660℃，优选580-640℃；如此，可进一步提高全厂的裂化气产量，即提高MTBE、轻汽油醚化、聚丙烯等轻烃加工装置所需原料的产量，调控全厂的炼化产品结构。
本发明中，重质原料或轻质原料与催化剂(即油剂)接触前的再生催化剂的温度为660-760℃。
基于本发明提供的装置和方法，多级反应混合料在第一反应区11a的油气表观停留时间为0.8-2s，在第二反应区11b的油气表观停留时间为5-15s，在第三反应区11c的油气表观停留时间为0.5-1.5s；裂解混合料在裂解提升管21的停留时间为0.8-3.5s，更优选为1-2.5s；在湍流反应段23b的停留时间为3-13s。
作为本发明中催化裂解的装置一种相对优选的具体实施方式，催化裂解的装置包括多级反应单元1、裂解反应单元2和再生单元3，多级反应单元1包括多变径提升管11和多级反应沉降器12；多变径提升管11包括沿其长度方向依次串联且同轴的多级反应预提升段11d、第一反应区11a、第二反应区11b和第三反应区11c，第三反应区11c的出口与多级反应沉降器12连通，在第一反应区11a的顶部至第二反应区11b的底部设置有急冷油气单元9，第一反应区11a、第二反应区11b和第三反应区11c的内径之比为1:2-5:0.7-1.5、长度之比为1:0.3-2:0.5-2，第三反应区11c与第二反应区11b的连接处设置有向第二反应区11b内部延伸的内输送管13，内输送管13与第三反应区11c的内径之比为0.9-1.2:1，内输送管13与第二反应区11b的长度之比为0.15-0.8:1，内输送管13 的侧壁上设置有多个内输送孔13a，内输送孔13a与第三反应区11c的入口之间的距离为0.5-2m，多个内输送孔13a的总面积与内输送管13的截面积之比为0.05-0.3:1，第二反应区11b内设置有与第一反应区11a连接的分布内构件14，分布内构件14包括与第一反应区11a连接的分布筒14b和与分布筒14b连接的端部分布板14c，分布筒14b和端部分布板14c上均设置有多个分布孔14a，端部分布板14c位于分布筒14b远离第一反应区11a的一端，端部分布板14c的面积与第二反应区11b的截面积之比为0.7-0.85:1，端部分布板14c上开设的多个分布孔14a的面积之和与该端部分布板14c的垂直投影面积之比为0.1-0.3:1，分布筒14b上分布孔14a的面积之和与第一反应区11a的截面积之比为0.6-4:1，分布筒14b上的分布孔14a处设置有向该分布筒14b的外侧延伸的侧通道14d，侧通道14d远离分布筒14b的开口断面为朝向第一反应区11a的斜面，多级反应预提升段11d和第一反应区11a之间的过渡段与多级反应进料单元7连通，多级反应沉降器12包括由下至上设置的多级反应汽提段12a和多级反应沉降段12b，多级反应沉降段12b内设置有一级或多级反应旋风分离器12c，多级反应汽提段12a的底部设置有多级反应催化剂出口15；
裂解反应单元2包括同轴设置的裂解提升管21和裂解沉降器23，裂解提升管21包括沿其长度方向依次串联的裂解预提升段21b和裂解反应段21a，裂解预提升段21b与裂解反应段21a之间的过渡段与裂解进料单元8连通，裂解沉降器23包括沿其长度方向依次设置的裂解汽提段23a、湍流反应段23b和裂解沉降段23c，裂解反应段21a的出口位于湍流反应段23b靠近裂解汽提段23a的一端，裂解反应段21a的出口处设置有裂解分布器22，裂解汽提段23a的底部设置有与第二反应区11b底部连通的裂解催化剂出口25，湍流反应段23b、裂解汽提段23a和裂解沉降段23c的内径之比为1:0.4-0.8:2.5-5、长度之比为1:1.5-3:1.5-4，在裂解沉降段23c内布置一级或两级裂解旋风分离器23d，裂解沉降段23c的气相出口通过裂解气相输送管24与多级反应沉降器12顶部的集气室连通，裂解汽提段23a的底部设置有汽提蒸汽单元10；
再生单元3内设置有再生旋风分离器32，且侧壁上设有再生催化剂出口31、底部设有再生空气入口33、顶部设有再生烟气出口34，多级反应催化剂出口15、裂解催化剂出口25均通过待生管路5与再生单元3的入料口连通，再生催化剂 出口31与多级反应预提升段11d和裂解预提升段21b通过再生管路6连通，再生管路6上设置有催化剂取热器4；
多变径提升管11与裂解提升管21的长度之比为1:0.5-1.2，内输送孔13a和分布孔14a内安装或涂抹有耐磨衬里。
以下将通过实施例对本发明进行详细描述。
以下实施例中，重质原料(重油)、轻烃I、轻烃II均由中国石油化工集团有限公司某炼厂提供，催化剂I、催化剂II均由中国石化催化剂有限公司提供，产品分布和汽油组成通过色谱法检测；在无特殊说明的情况下，其余原料均采用常规的市售品，所涉及的参数均采用现有的相应方法。
实施例1
采用上述提供的催化裂解的装置相对优选的具体实施方式，使用表2中的催化剂I加工表1中的重质原料油，生产超低烯烃汽油的过程情况。实施例1中原料性质见表1，回炼轻汽油采用轻烃A为轻烃原料，回炼轻汽油占重质原料的30％，其中一半为装置自产轻汽油，一半来自本装置外的其它装置。
实施例1提供的装置中，多变径提升管11总长约18米，多级反应预提升段11d的直径为0.25米、高度为1.5米，第一反应区11a的直径为0.30米、高度为4米；第二反应区11b的直径为1.2米、高度为7.6米，内输送管13的直径为0.30米、高度为2.5米；第三反应区11c的直径为0.30米、高度为4米，第一反应区11a与第二反应区11b结合部位的纵剖面为等腰梯形，等腰梯形的侧线与轴线的夹角为45°，等腰梯形高度为0.45米；第二反应区11b与第三反应区11c结合部位的纵剖面为等腰梯形，等腰梯形的侧线与轴线的夹角为45°，等腰梯形高度为0.45米；内输送孔13a与第三反应区11c的入口之间的距离为1m，多个内输送孔13a的总面积与内输送管13的截面积之比为0.2:1，分布内构件14形成帽头型分布内构件，端部分布板14c的面积与第二反应区11b的截面积之比为0.8:1，端部分布板14c上开设的多个分布孔14a的面积之和与该端部分布板14c的垂直投影面积之比为0.2:1，分布筒14b上分布孔14a的面积之和与第一反应区11a的截面积之比为2:1；
裂解提升管21总高度13米，其中裂解预提升段21b的高度为2米、直径为0.12米，裂解反应段21a的高度为11米、直径为0.15米，裂解汽提段23a的高度为3米、直径为0.35米，湍流反应段23b的高度为1.5米、直径为0.5米，裂解沉降段23c的高度为2.5米、直径为1.3米，裂解气相输送管24的直径为0.2米。
从多级反应预提升段11d底部进入的多级反应预提升气A与从催化剂取热器4或再生单元3直接过来的经再生管路6输送过来的再生催化剂(催化剂I)混合上行，在多级反应预提升段11d顶部与第一反应区11a主体的连接处，经水蒸气雾化的重质原料经重油进料系统7进入多变径提升管11，油剂在第一反应区11a内接触反应，沿多变径提升管11上行经分布内构件14进入第二反应区11b；由于油气表观速度的下降及油气转向等因素，在第二反应区11b内催化剂呈湍流床或浓相输送床状态，强化油剂的接触；在第二反应区11b与第一反应区11a过渡段内可从急冷油气单元9注入急冷剂，从第二反应区11b的底部注入裂解催化剂出口25输出的待生催化剂，以补充第二反应区11b内的催化剂活性；
从裂解预提升段21b的底部注入裂解预提升气B与从催化剂取热器4或再生单元3直接过来的再生催化剂(催化剂I)混合上行，在裂解预提升段21b顶部与裂解反应段21a的连接处，经水蒸气雾化或直接机械雾化的轻质原料(轻烃I)经裂解进料单元8进入裂解提升管21，轻质原料与催化剂在裂解反应段21a接触反应并上行，经裂解反应段21a顶部末端的裂解分布器22进入湍流反应段23b的底部，将反应后的油气与催化剂实现分离；催化剂在裂解汽提段23a的底部经过/不经汽提蒸汽单元10通入的蒸汽汽提后，再通过裂解催化剂出口25送入第二反应区11b内，经分离后的轻烃反应后油气经裂解沉降段23c上的轻烃输送管进入多级反应沉降器12顶部的集气室或直接进入多级反应沉降器12；第二反应区11b内的多级反应混合料与裂解提升管21内反应后的待生催化剂经过充分混合接触及再分配后，通过相对于第二反应区11b缩径的内输送管13进入第三反应区11c，在第三反应区11c顶部与多级反应沉降段12b内的多级反应旋风分离器12c相连通进行多级反应分离，重质原料反应后油气从多级反应旋风分离器12c的顶部输出，进而多级反应沉降器12底部经大转油线等进入油气分馏塔；重质原料反应后含有焦炭的催化剂从多级反应旋风分离器12c下部的输送管进 入多级反应沉降器12底部的多级反应汽提段12a，经多级反应汽提蒸汽F进行汽提后的待生催化剂经待生管路5进入再生单元3，与再生单元3底部进入的再生空气C接触发生氧化反应，再生烟气D从再生单元3顶部输出进入烟气轮机等进一步回收能量和处理；烧掉焦炭从而生成的高温再生催化剂经进再生管路6进入催化剂取热器4，将低温水蒸气G取热后变为高温水蒸气H，经冷却后的再生催化剂经出催化剂取热器4的再生管路6回流至第一反应区11a底部和裂解提升管21的底部循环使用。
其余的操作条件，获得的产品分布和汽油主要性质见表3。
对比例1
催化裂解的装置包括实施例1中采用的多级反应单元1和常规的裂解单元，裂解单元采用常规的提升管、不含有裂解沉降器，且提升管直接连接旋风分离器，使用表2中的催化剂I加工表1中重质原料油，生产超低烯烃汽油的过程情况。多级反应单元1的各项参数与实施例1相同，裂解单元的提升管出口温度与实施例1中裂解提升管21的出口温度相同。
其余的操作条件，获得的产品分布和汽油主要性质见表3。
对比例2
催化裂解的装置包括实施例1中所述的多级反应单元1、不设置裂解反应单元2，不回炼轻烃，使用表2中的催化剂I加工表1中重质原料油，生产超低烯烃汽油的过程情况；多级反应单元1的各项参数与实施例1相同。
其余的操作条件，获得的产品分布和汽油主要性质见表3。
对比例3
催化裂解的装置包括实施例1中采用的多级反应单元1、不设置裂解反应单元2，不回炼轻烃，且第二反应区11b的直径为0.9米，第二反应区11b内不设置内输送管13，多级反应单元1的其余各项参数与对比例2相同。使用催化剂I加工表1中重质原料油，进行生产低烯烃汽油的操作模式。
其余的操作条件，获得的产品分布和汽油主要性质见表3。
实施例2
与实施例1相比，本实施例采用的催化裂解的装置除第一反应区11a与第二反应区11b连接的分布内构件14为蘑菇头型分布器、催化剂采用表2中的催化剂II、操作条件不同外，其余的装置结构参数等均与实施例1相同。本实施例中原料性质见表1，其中回炼轻汽油采用轻烃II为轻质原料，回炼轻汽油占重质原料的30％，其中一半为装置自产轻汽油，一半来自本装置外的其它装置。
其余的操作条件，获得的产品分布和汽油主要性质见表3。
对比例4
催化裂解的装置包括实施例1中采用的多级反应单元1和常规的裂解单元，裂解单元采用常规的提升管、不含有裂解沉降器，且提升管直接连接旋风分离器，使用表2中的催化剂II加工表1中重质原料油，进行多产裂化气的操作模式。多级反应单元1的各项参数与实施例2相同，裂解单元的提升管出口温度与实施例1中裂解提升管21的出口温度相同。
其余的操作条件，获得的产品分布和汽油主要性质见表3。
对比例5
催化裂解的装置包括实施例1中采用的多级反应单元1、不设置裂解反应单元2，不回炼轻烃，使用表2中的催化剂II加工表1中重质原料油，进行多产裂化气的操作模式。多级反应单元1的各项参数与实施例2相同。
其余的操作条件，获得的产品分布和汽油主要性质见表3。
对比例6
催化裂解的装置包括实施例1中采用的多级反应单元1、不设置裂解反应单元2，不回炼轻烃，且第二反应区11b的直径为0.9米，第二反应区11b内不设置内输送管13，多级反应单元1的其余各项参数与对比例5相同。
其余的操作条件，获得的产品分布和汽油主要性质见表3。
表1原料性质
表2催化剂性质
表3操作条件及产品分布
从表3可以看出，实施例1、对比例1、对比例2和对比例3四者均是生产超低烯烃汽油模式，实施例1和对比例1回炼30％的轻烃I，其中一半来自本装置外的富含烯烃的轻汽油；对比例2和对比例3都不回炼汽油；实施例1相对于对比例1、对比例2和对比例3，汽油烯烃分别降低2.6、3.6和6.4个百分点，丙烯产率分别增加0.5、0.7和1.1个百分点；表明采用本发明提供的装置和方法能明显降低汽油烯烃含量，并适度增加丙烯产率。
实施例2、对比例4、对比例5和对比例6四者均是多产裂化气操作模式，实施例2和对比例4回炼30％的轻烃II，其中一半来自本装置外的富含烯烃的轻汽油；对比例5和对比例6都不回炼汽油；实施例2相对于对比例4、对比例5和对比例6，汽油烯烃分别降低3.3、11.2和13.8个百分点，丙烯产率分别增加1.3、2.6和4.5个百分点；表明采用发明提供的装置和方法能明显降低汽油烯烃含量，并增加丙烯产率。
本发明提供的装置和方法可促进对整个炼厂汽油池的汽油烯烃的降低，裂化气收率明显提高，但是对整个炼厂而言，裂化气的产率或产量都有明显提升。
以上结合附图详细描述了本发明的优选实施方式，但是，本发明并不限于上述实施方式中的具体细节，在本发明的技术构思范围内，可以对本发明的技术方案进行多种简单变型，这些简单变型均属于本发明的保护范围。
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，本发明对各种可能的组合方式不再另行说明。此外，本发明的各种不同的实施方式之间也可以进行任意组合，只要其不违背本发明的思想，其同样应当视为本发明所公开的内容。

Claims (18)

  1. 一种催化裂解的装置，其特征在于，该装置包括多级反应单元(1)和裂解反应单元(2)，所述多级反应单元(1)包括多变径提升管(11)和多级反应沉降器(12)，所述裂解反应单元(2)包括裂解提升管(21)和裂解沉降器(23)；所述多变径提升管(11)包括沿其长度方向依次串联的第一反应区(11a)、第二反应区(11b)和第三反应区(11c)，所述第三反应区(11c)的出口与所述多级反应沉降器(12)连通，所述第二反应区(11b)的内径大于所述第一反应区(11a)和第三反应区(11c)的内径，所述第三反应区(11c)与所述第二反应区(11b)的连接处设置有向所述第二反应区(11b)内部延伸的内输送管(13)；
    所述裂解沉降器(23)包括沿其长度方向依次设置的裂解汽提段(23a)、湍流反应段(23b)和裂解沉降段(23c)，所述裂解提升管(21)的出口位于所述湍流反应段(23b)内，所述裂解汽提段(23a)设置有与所述第二反应区(11b)连通的裂解催化剂出口(25)。
  2. 根据权利要求1所述的装置，其特征在于，所述内输送管(13)与所述第三反应区(11c)的内径之比为0.9-1.2:1，所述内输送管(13)与所述第二反应区(11b)的长度之比为0.15-0.8:1。
  3. 根据权利要求1所述的装置，其特征在于，所述第一反应区(11a)、第二反应区(11b)和第三反应区(11c)的内径之比为1:2-5:0.7-1.5、长度之比为1:0.3-2:0.5-2；
    所述湍流反应段(23b)、裂解汽提段(23a)和裂解沉降段(23c)的内径之比为1:0.4-0.8:2.5-5、长度之比为1:1.5-3:1.5-4；
    所述多变径提升管(11)与所述裂解提升管(21)的长度之比为1:0.5-1.2。
  4. 根据权利要求1所述的装置，其特征在于，所述内输送管(13)的侧壁上设置有至少一个内输送孔(13a)，所述内输送孔(13a)与所述第三反应区(11c)的入口之间的距离为0.5-2m。
  5. 根据权利要求4所述的装置，其特征在于，至少一个所述内输送孔(13a)的总面积与所述内输送管(13)的截面积之比为0.05-0.3:1。
  6. 根据权利要求1至5中任意一项所述的装置，其特征在于，所述第二反应区(11b)内设置有与所述第一反应区(11a)连接的分布内构件(14)，所述分布内构件(14)上设有多个分布孔(14a)。
  7. 根据权利要求6所述的装置，其特征在于，所述分布内构件(14)包括与所述第一反应区(11a)连接的分布筒(14b)和与所述分布筒(14b)连接的端部分布板(14c)，所述分布筒(14b)和所述端部分布板(14c)上均设置有所述分布孔(14a)，所述端部分布板(14c)位于所述分布筒(14b)远离所述第一反应区(11a)的一端。
  8. 根据权利要求7所述的装置，其特征在于，所述分布筒(14b)上分布孔(14a)的面积之和与所述第一反应区(11a)的截面积之比为0.6-4:1；
    所述端部分布板(14c)上分布孔(14a)的面积之和与该端部分布板(14c)的垂直投影面积之比为0.1-0.3:1。
  9. 根据权利要求7所述的装置，其特征在于，所述分布筒(14b)上的分布孔(14a)处设置有向该分布筒(14b)的外侧延伸的侧通道(14d)，所述侧通道(14d)远离所述分布筒(14b)的开口断面为朝向所述第一反应区(11a)的斜面。
  10. 根据权利要求7所述的装置，其特征在于，所述端部分布板(14c)的面积与所述第二反应区(11b)的截面积之比为0.7-0.85:1。
  11. 根据权利要求1至5中任意一项所述的装置，其特征在于，所述第一反应区(11a)、第二反应区(11b)、第三反应区(11c)和内输送管(13)为同轴设置，所述裂解催化剂出口(25)位于所述裂解汽提段(23a)的底部且与 所述第二反应区(11b)的底部连通。
  12. 根据权利要求1至5中任意一项所述的装置，其特征在于，所述多变径提升管(11)还包括与所述第一反应区(11a)的入口连接的多级反应预提升段(11d)，所述裂解提升管(21)包括沿其长度方向依次串联的裂解预提升段(21b)和裂解反应段(21a)，所述裂解反应段(21a)的出口位于所述湍流反应段(23b)靠近所述裂解汽提段(23a)的一端，所述多级反应预提升段(11d)和所述第一反应区(11a)之间的过渡段与多级反应进料单元(7)连通，所述裂解预提升段(21b)与所述裂解反应段(21a)之间的过渡段与裂解进料单元(8)连通。
  13. 根据权利要求12所述的装置，其特征在于，所述裂解反应段(21a)的出口处设置有裂解分布器(22)，所述裂解沉降段(23c)的气相出口与所述多级反应沉降器(12)连通。
  14. 根据权利要求1至5中任意一项所述的装置，其特征在于，该装置还包括再生单元(3)，所述多级反应沉降器(12)上设置有与所述再生单元(3)的入料口连通的多级反应催化剂出口(15)，所述再生单元(3)的再生催化剂出口(31)与所述多变径提升管(11)和/或所述裂解提升管(21)连通。
  15. 一种催化裂解的方法，其特征在于，采用权利要求1至14中任意一项所述的装置，该方法包括以下步骤：
    重质原料与多级反应催化剂混合形成的多级反应混合料从所述第一反应区(11a)的入口输入，依次进入所述第一反应区(11a)、所述第二反应区(11b)，再从所述内输送管(13)进入所述第三反应区(11c)，以进行多级催化反应得到多级反应物，所述多级反应物经所述多级反应沉降器(12)进行多级反应分离得到多级反应油气产物和多级反应待生催化剂；
    轻质原料与裂解催化剂混合形成的裂解混合料从所述裂解提升管(21)的入口输入，经所述裂解提升管(21)进行裂解催化反应得到裂解反应物，所述裂 解反应物进入所述湍流反应段(23b)进行湍流反应后，经所述裂解沉降段(23c)进行裂解分离、经所述裂解汽提段(23a)进行裂解汽提得到裂解油气产物和裂解待生催化剂；
    将所述裂解待生催化剂从所述裂解催化剂出口(25)输入所述第二反应区(11b)内，与所述多级反应混合料混合。
  16. 根据权利要求15所述的方法，其特征在于，所述内输送管(13)的入口温度为510-540℃，所述裂解提升管(21)的出口温度为540-620℃。
  17. 根据权利要求15所述的方法，其特征在于，所述内输送管(13)的入口温度为550-640℃，所述裂解提升管(21)的出口温度为560-660℃。
  18. 根据权利要求15至17中任意一项所述的方法，其特征在于，所述多级反应混合料在所述第一反应区(11a)的停留时间为0.8-2s，在所述第二反应区(11b)的停留时间为5-15s，在所述第三反应区(11c)的停留时间为0.5-1.5s；
    所述裂解混合料在所述裂解提升管(21)的停留时间为0.8-3.5s，在所述湍流反应段(23b)的停留时间为3-13s。
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