A method for preparing low carbon olefins and aromatic hydrocarbons by processing LCO with a downflow bed

Abstract

The present application relates to a method for preparing low carbon olefins and aromatics by processing LCO in a downflow bed, which comprises: S1, contacting LCO raw material with an aromatization catalytic cracking catalyst in a hydrogen atmosphere, and carrying out a catalytic cracking reaction in a downflow bed reactor, to obtain a reaction mixture; S2, separating the reaction mixture to obtain product oil gas containing low carbon olefins and aromatics and spent catalyst. The method of the present application can improve the yield of low carbon olefins and aromatics, and reduce the yield of coke.

Description

TECHNICAL FIELD

[0001] The present application relates to a method for preparing low-carbon olefins and aromatics by processing LCO with a downflow bed. BACKGROUND

[0002] With the upgrading of gasoline and diesel quality and the change of demand structure, the domestic consumption of gasoline to diesel ratio continues to decline, and the problem of excess diesel in refineries will become increasingly prominent. Low-carbon olefins and aromatics are important organic chemical raw materials, and the domestic market demand is strong, with high dependence on foreign countries. On the one hand, there is an excess of oil refining capacity, and on the other hand, there is a shortage of chemical raw materials, so the transformation of oil refining to chemical industry has become an inevitable trend.

[0003] Catalytic cracking diesel (LCO) is rich in aromatics and is an ideal raw material for producing aromatics. Single-ring aromatics have crackability and can be directly catalytically cracked to produce aromatics. Polycyclic aromatics are easily saturated into single-ring aromatics after hydrotreating. Therefore, by hydrotreating LCO and combining catalytic cracking process, it is expected to improve the production capacity of aromatics.

[0004] The current catalytic cracking process for producing low-carbon olefins and aromatics mainly uses a riser reactor and a downflow bed reactor. The most widely used is the DCC technology developed by Sinopec Petroleum Chemical Research Institute, which uses a riser reactor. Patents CN101362963, CN101747928, CN1667089, etc. disclose a method for producing more propylene and aromatics from heavy feedstock by catalytic cracking, which can make the propylene yield reach more than 40 wt% by recycling difficult cracking feedstock or circulating cracking feedstock, etc. Distillate oil, and through aromatics extraction technology, BTX can be produced by extracting from the aromatics-rich fraction. The above patents all use a riser reactor, but the riser reactor itself has the defects of long gas-solid contact time, non-uniform annular core distribution in the radial direction, and large axial backmixing, which will affect the conversion rate and selectivity of the catalytic cracking reaction.

[0005] In view of the problems existing in the riser reactor, some famous foreign oil companies proposed the concept of gas-solid ultra-short contact down-flow bed reactor, and believed that shortening the residence time of oil and gas could help to improve the selectivity of intermediate products such as propylene, and began to develop new down-flow bed reactor catalytic cracking technology. The down-flow bed reactor adopts gas-solid flow in the gravity field, so it has the characteristics of approximate plug flow distribution and short gas-solid residence time, which helps to improve the uniformity of products, inhibit excessive reaction and improve the selectivity of target products. Patent CN1113659 discloses a flexible folding design using riser coupled down-flow bed, which can produce low carbon olefins and clean gasoline. Patent US6656346B2 discloses a high-severity catalytic cracking process HS-FCC (High-Severity Fluid Catalytic Cracking), which uses a down-flow bed reactor for catalytic cracking at high severity, and the propylene yield can reach more than 20, while the dry gas and coke yield is low.

[0006] The above patents, whether using a riser reactor or a down-flow bed reactor, although suitable for the production of low carbon olefins by saturate cracking in LCO, are not suitable for the dealkylation of aromatics with high activation energy barriers, resulting in low yields of low carbon olefins and aromatics. In order to meet the growing demand for low carbon olefins and aromatics and other organic chemical raw materials, and to solve the problem of excess fuel production capacity, it is necessary to develop a method for processing LCO catalytic cracking to produce low carbon olefins and aromatics. SUMMARY

[0007] The purpose of the present application is to provide a method for processing LCO to produce low carbon olefins and aromatics using a down-flow bed, which has high yield of low carbon olefins and aromatics and low yield of coke.

[0008] In order to achieve the above-mentioned purpose, the present application provides a method for processing LCO to produce low carbon olefins and aromatics using a down-flow bed, which comprises:

[0009] S1, contacting LCO raw material with aromatization catalytic cracking catalyst in the upper part of the down-flow bed reactor in a hydrogen atmosphere and carrying out catalytic cracking reaction in a down-flow direction to obtain a reaction mixture;

[0010] S2, separating the reaction mixture to obtain product oil gas containing low carbon olefins and aromatics and spent catalyst.

[0011] Optionally, the conditions of the catalytic cracking reaction include: reaction temperature of 500-750℃, reaction pressure of 0.1-2.0MPa, hydrogen partial pressure of 0.1-1.0MPa, catalyst to oil weight ratio of 10-50, and residence time of the LCO raw material for hydrogenation of 0.2-5.0 seconds;

[0012] The circulation rate of the aromatization catalytic cracking catalyst in the down-flow bed reactor is 100-500 kg / m 2 s, and the bed particle concentration is 1-5%.

[0013] Optionally, the conditions of the catalytic cracking reaction include that the reaction temperature is 550-650℃, the reaction pressure is 0.4-15 MPa, the hydrogen partial pressure is 0.2-0.8 MPa, the weight ratio of the catalyst to oil is 25-35, and the residence time of the LCO raw material is 0.5-1.5 seconds.

[0014] The circulation rate of the aromatization catalytic cracking catalyst in the down-flow bed reactor is 120-450 kg / m 2 s, and the bed particle concentration is 1.5-4.5%.

[0015] Optionally, the method further comprises: optionally stripping and coke-burning regeneration of the spent catalyst, and returning at least part of the obtained regenerated catalyst to the down-flow bed reactor of step S1 for use as the aromatization catalytic cracking catalyst.

[0016] Optionally, the conditions of the coke-burning regeneration include that the temperature is 600-800℃, the pressure is 0.1-2.0 MPa, and the time is 1-10 minutes.

[0017] Optionally, the separation is carried out in a rapid separation device.

[0018] Optionally, the method further comprises: preheating the LCO raw material before introducing it into the down-flow bed reactor, and the temperature of the preheated LCO raw material is 100-300℃.

[0019] Optionally, the hydrogen atmosphere contains hydrogen and / or dry gas.

[0020] Optionally, the aromatization catalytic cracking catalyst contains 1-60% by weight of zeolite, 5-95% by weight of inorganic oxide, 0-80% by weight of clay, and 0-20% by weight of active metal component based on the total weight of the aromatization catalytic cracking catalyst on a dry basis, wherein the active metal component is calculated as active metal oxide.

[0021] Optionally, the zeolite contains 40-100% by weight of medium-pore zeolite and 0-60% by weight of large-pore zeolite based on the total weight of the zeolite on a dry basis.

[0022] Optionally, the large-pore zeolite is selected from one or more of high-silica Y, rare earth Y, rare earth hydrogen Y, and ultra-stable Y, and the medium-pore zeolite is selected from ZSM series zeolite and / or ZRP zeolite.

[0023] The inorganic oxide is selected from silicon dioxide and / or diatomic aluminum oxide.

[0024] The clay is selected from one or more of kaolin, montmorillonite, diatomite, halloysite, saponite, laponite, sepiolite, attapulgite, hydrotalcite and bentonite;

[0025] The active metal component contains one or more of Fe, Co, Ni, Cu, Zn and rare earth metals.

[0026] Optionally, in step S1, the contacting of the LCO feedstock with the aromatization catalytic cracking catalyst in the upper part of the downflow bed reactor and performing catalytic cracking reaction in a concurrent manner comprises:

[0027] The aromatization catalytic cracking catalyst is distributed through the catalyst inlet distributor and then is sent to the upper part of the downflow bed reactor to contact with the LCO feedstock and perform catalytic cracking reaction in a concurrent manner;

[0028] The catalyst inlet distributor comprises a shell; a solid inlet is arranged at the upper end of the shell, and a mixture outlet is arranged at the lower end of the shell; the mixture outlet is in fluid communication with the catalyst inlet in the upper part of the downflow bed reactor; the inner cavity of the shell is divided into a fluidization zone in the upper part and a gas-solid mixing zone in the lower part by a sealing plate arranged in the radial direction; a fluidization air inlet is arranged at the lower part of the fluidization zone, and a feedstock oil gas inlet is arranged at the upper part of the gas-solid mixing zone; a plurality of Venturi-type overflow pipes for fluid communication between the fluidization zone and the gas-solid mixing zone are arranged on the sealing plate; an upwardly-open first gas distribution plate is arranged between the outer wall of the upper end of the overflow pipe and the shell in the radial direction; and the fluidization air inlet is arranged below the first gas distribution plate.

[0029] Optionally, a fluidization air inlet cavity is formed between the lower surface of the first gas distribution plate, the upper surface of the sealing plate, the inner side wall of the shell and the outer side wall of the overflow pipe, and the fluidization air inlet cavity is in fluid communication with the fluidization air inlet;

[0030] Optionally, the inlet distributor further comprises a second gas distribution plate arranged between the outer wall of the lower end of the overflow pipe and the shell in the radial direction; a feedstock oil gas inlet cavity is formed between the upper surface of the second gas distribution plate, the lower surface of the sealing plate, the inner side wall of the shell and the outer side wall of the overflow pipe, and the feedstock oil gas inlet cavity is in fluid communication with the feedstock oil gas inlet;

[0031] Optionally, the overflow pipe comprises, in sequence, a diameter-reducing section, a straight pipe section and a diameter-increasing section in fluid communication, and the sealing plate is sealingly connected to the outer side wall of the straight pipe section of the overflow pipe;

[0032] Optionally, the overflow pipes are 5-9, and the overflow pipes are uniformly distributed in the shell.

[0033] Optionally, the shell comprises a cylindrical shell and a reverse trapezoidal shell arranged in an axial direction and in fluid communication; the upper end of the cylindrical shell is provided with the solid inlet, the lower end of the reverse trapezoidal shell is provided with the mixture outlet, and the sealing plate is arranged in a radial direction at the bottom of the cylindrical shell.

[0034] Compared with the prior art, the method has the following technical effects:

[0035] (1) The raw material is suitable for not only LCO and other light catalytic cracking raw materials, but also deep catalytic cracking process of poor heavy raw materials with high density and low hydrogen content.

[0036] (2) The down-flow bed reactor is adopted, which has the characteristics of approximate plug flow distribution and short gas-solid residence time, which helps to improve the propylene yield, and is also conducive to inhibiting the generation of dry gas and coke to a certain extent, and the product selectivity is high.

[0037] (3) The hydrogenation operation and the down-flow bed reactor are coupled, the existing hydrogenation processing technology has the problem of high coke generation, the down-flow bed reactor adopted in the present application can reduce the generation of coke, fully play the effect of hydrogenation, be conducive to the dealkylation reaction of aromatics, and increase the yield of aromatics. At the same time, under the hydrogenation atmosphere, it is helpful to reduce the hydrocarbon partial pressure and increase the low-carbon olefin yield.

[0038] (4) The catalytic cracking catalyst with aromatization function is adopted, which can increase the selectivity of aromatics under the hydrogenation atmosphere, and realize the production of aromatics.

[0039] Other features and advantages of the present application will be described in detail in the following specific embodiments. BRIEF DESCRIPTION OF DRAWINGS

[0040] The accompanying drawings are included to provide a further understanding of the present application, and constitute a part of the specification, and are used together with the following specific embodiments to explain the present application, but do not constitute a limitation on the present application. In the drawings:

[0041] Figure 1 is a structural schematic diagram of a system for processing LCO to prepare low-carbon olefins and aromatics according to the present application;

[0042] Figure 2 is a structural schematic diagram of one specific embodiment of the inlet distributor according to the present application;

[0043] Figure 3 is a structural schematic diagram of the overflow pipe according to the present application;

[0044] Figure 4Figure 1 is a schematic diagram of the overflow pipe arrangement of the present application.

[0045] Reference Signs List

[0046] 1 Hydrogen-containing medium 2 Inlet distributor 3 Regeneration slide valve

[0047] 4 Oil inlet nozzle 5 Downflow bed reactor 6 Rapid separator

[0048] 7 Cyclone 8 Product oil gas outlet line 9 Stripper

[0049] 10 Stripper oil gas line 11 Catalyst line 12 Stripper steam

[0050] 13 Spent slide valve 14 Regeneration medium 15 Riser regenerator

[0051] 16 Cyclone 17 Settler 18 Cyclone

[0052] 19 Flue gas outlet line

[0053] 111 Cylindrical shell 112 Fluidization air inlet 113 Overflow pipe

[0054] 114 First gas distribution plate 115 Seal plate 116 Feed oil gas inlet

[0055] 117 Second gas distribution plate 118 Inverted truncated cone shell DETAILED DESCRIPTION

[0056] The specific embodiments of the present application will be described in detail below with reference to the accompanying drawings. It should be understood that the specific embodiments described herein are merely intended to illustrate and explain the present application, and are not intended to limit the present application.

[0057] The present application provides a method for preparing low-carbon olefins and aromatic hydrocarbons by processing LCO in a downflow bed, which comprises: S1, contacting LCO raw material with an aromatic-catalytic cracking catalyst in a hydrogen atmosphere, and performing a catalytic cracking reaction in a downflow bed reactor, to obtain a reaction mixture; S2, separating the reaction mixture to obtain product oil gas containing low-carbon olefins and aromatic hydrocarbons, and spent catalyst.

[0058] The inventors of the present application find through a large amount of literature research and experimental research that harsh reaction conditions and short residence time are beneficial to the production of low-carbon olefins and aromatic hydrocarbons by LCO catalytic cracking. The conventional low-carbon olefin and aromatic hydrocarbon production method mainly uses a riser reactor, and the contact time is long and the back mixing is serious, which affects the selectivity of intermediate products such as propylene, and the effect of aromatic hydrocarbon dealkylation is poor and affects the yield of aromatic hydrocarbons. In the method of the present application, the downflow bed reactor has the characteristics of short residence time and approximate plug flow, and in combination with the hydrogen-containing reaction atmosphere and the catalytic action of the aromatization catalyst, it is beneficial to improve the yield of low-carbon olefins such as propylene and the yield of aromatic hydrocarbons, and effectively inhibit the generation of dry gas and coke.

[0059] According to the present application, the catalytic cracking is well known to those skilled in the art, and in one specific embodiment of the present application, the conditions of the catalytic cracking reaction include: the reaction temperature is 500-750℃, preferably 550-650℃, the reaction pressure is 0.1-2.0MPa, preferably 0.4-1.5MPa, the hydrogen partial pressure is 0.1-1MPa, preferably 0.2-0.8MPa, the weight ratio of catalyst to oil is 10-50, preferably 25-35, and the residence time of the LCO raw material for hydrogenation is 0.2-5.0 seconds, preferably 0.5-1.5 seconds. Among them, the reaction temperature refers to the outlet of the downflow bed reactor, when the catalytic cracking reaction is carried out under the above hydrogen-containing, high-temperature, large catalyst-oil ratio reaction conditions, and in combination with the hydrogen-containing reaction atmosphere and the aromatization catalytic cracking catalyst, the reaction conversion rate can be further improved, and the yield of olefins and aromatic hydrocarbons is increased.

[0060] In one specific embodiment of the present application, the circulation rate of the aromatization catalytic cracking catalyst in the downflow bed reactor is 100-500kg / m 2 s, and the bed particle concentration is 1-5%; the circulation rate is 120-450kg / m 2 s, and the bed particle concentration is 1.5-4.5%.

[0061] In a preferred embodiment of the present application, the method further comprises: optionally stripping and coke-burning regeneration of the spent catalyst, and returning at least part of the obtained regenerated catalyst to the downflow bed reactor of step S1 as the aromatization catalytic cracking catalyst, preferably, the temperature of the downflow bed reactor is optimized and controlled by controlling the temperature of the regenerated catalyst. Among them, the stripping and coke-burning regeneration method is well known to those skilled in the art, and will not be repeated here, the medium used for stripping can be steam, and the weight ratio of steam to spent catalyst can be selected as needed, for example, it can be 0.1-0.3. The conditions of coke-burning regeneration include: the temperature is 600-800℃, the pressure is 0.1-2.0MPa, the time is 1-10 hours, and the regeneration medium can be air. In one embodiment, the temperature of coke-burning regeneration refers to the outlet temperature of the regenerator.

[0062] In one embodiment of the present application, the separation in step S2 is performed in a rapid separation device. In the process of the present application, the reaction mixture obtained by the catalytic cracking reaction in the down-flow bed reactor is introduced into the rapid separation device for separation, which can effectively inhibit the excessive cracking reaction of the raw material, reduce the dry gas and coke yield, and thus improve the selectivity of the reaction.

[0063] In one embodiment of the present application, the process further comprises: introducing the LCO raw material into the down-flow bed reactor after preheating, and the temperature of the preheated LCO raw material is 100-300°C, preferably 200-300°C.

[0064] According to the present application, the hydrogen atmosphere refers to an atmosphere containing hydrogen, and preferably, the hydrogen atmosphere contains hydrogen and / or dry gas. In one embodiment of the present application, the gas containing hydrogen is introduced into the down-flow bed reactor from the upper part of the down-flow bed reactor to form the hydrogen atmosphere, and in this way, the gas containing hydrogen also serves as the fluidizing gas.

[0065] In one embodiment of the present application, the aromatization catalytic cracking catalyst contains 1-60 wt% of zeolite, 5-95 wt% of inorganic oxide, 0-80 wt% of clay and 0-20 wt% of active metal component, based on dry basis and based on the total weight of the aromatization catalytic cracking catalyst; preferably, the aromatization catalyst contains 10-40 wt% of zeolite, 20-80 wt% of inorganic oxide, 5-40 wt% of clay and 1-15 wt% of active metal component; and the zeolite contains 40-100 wt% of mesoporous zeolite and 0-60 wt% of macroporous zeolite, based on dry basis and based on the total weight of the zeolite.

[0066] According to the present application, the large-pore zeolite can be selected from one or more of Y-series zeolites, such as high-silica Y, rare earth Y (REY), rare earth hydrogen Y (REHY), and ultra-stable Y, and the medium-pore zeolite can be selected from ZSM-series zeolites and / or ZRP zeolites. Among them, the ZSM-series zeolite can include, but is not limited to, one or two of ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM-38, ZSM-48, and other zeolites with similar structures. In addition, the zeolite can also include a BETA-type molecular sieve with both large-pore structure and medium-pore structure. The present application does not make any specific limitation on the specific type of inorganic oxide contained in the aromatization catalyst, and the inorganic oxide commonly used by those skilled in the art as a binder can be used, for example, the inorganic oxide can be selected from silicon dioxide and / or aluminum oxide, preferably, the inorganic oxide is a mixture of silicon dioxide and aluminum oxide, and the content of silicon dioxide in the mixture can be 45-95% by weight on a dry weight basis. The present application also does not make any specific limitation on the type of clay contained in the aromatization catalyst, and the clay commonly used by those skilled in the art as a carrier can be used, for example, the clay can be selected from one or more of kaolin, montmorillonite, diatomite, halloysite, saponite, rectorite, sepiolite, attapulgite, hydrotalcite, and bentonite. According to the present application, the active metal component contained in the aromatization catalyst can be a metal having dehydrogenation effect, preferably a transition metal element in the fourth period and the fifth period, which can include, but is not limited to, one or more of Fe, Co, Ni, Cu, Zn, and rare earth metals, preferably, the active metal component exists in the form of metal oxide and / or metal hydroxide.

[0067] As shown in Figure 1 In a preferred embodiment of the present application, the following method is used to process LCO to prepare low-carbon olefins and aromatic hydrocarbons:

[0068] The aromatization catalytic cracking catalyst is regenerated and enters the inlet distributor 2 through the slide valve 3, and is fluidized in the hydrogen-containing medium 1. After the LCO raw material is preheated, it enters the top end of the downflow bed reactor 5 through the oil injection nozzle 4, and is mixed and contacted with the aromatization catalytic cracking catalyst. The LCO raw material oil gas and the aromatization catalytic cracking catalyst flow downward in parallel into the downflow bed reactor 5 to perform catalytic cracking reaction, and a reaction mixture is obtained. The reaction mixture is separated at the bottom end of the downflow bed reactor 5 by the rapid separator 6. The spent catalyst obtained after separation enters the stripper 9 to be stripped, and the product oil gas containing low-carbon olefins and aromatic hydrocarbons obtained is separated again in the cyclone separator 7, and finally enters the fractionation unit through the product oil gas discharge pipeline 8. The spent catalyst separated by the cyclone separator 7 enters the stripper through the catalyst pipeline 11. The spent catalyst separated by the rapid separator 6 enters the stripper 9 and is stripped under the action of the stripping steam 12. After the spent catalyst is stripped, the oil gas stripped out is introduced into the cyclone separator 7 through the stripping oil gas pipeline 10. The stripped catalyst enters the riser regenerator 15 through the spent slide valve 13, is coked and regenerated in the regeneration medium 14, is separated in the cyclone separator 16 at the outlet of the riser, and is settled in the settler 17. The flue gas generated by regeneration is separated by the cyclone separator 18 of the settler, is discharged from the flue gas outlet pipeline 19. The regenerated catalyst enters the inlet distributor 2 of the reactor through the regeneration slide valve 3, and performs the next cycle.

[0069] In a preferred embodiment of the present application, in step S1, the contacting and catalytic cracking downward in parallel of the LCO raw material with the aromatization catalytic cracking catalyst in the upper part of the downflow bed reactor comprises: after the aromatization catalytic cracking catalyst is distributed through the catalyst inlet distributor, it is sent to the upper part of the downflow bed reactor to contact with the LCO raw material and perform catalytic cracking downward in parallel. In an embodiment, as shown in Figure 2As shown, the catalyst inlet distributor comprises a shell; a solid inlet is arranged at the upper end of the shell, and a mixture outlet is arranged at the lower end of the shell; the mixture outlet is in fluid communication with the catalyst inlet of the upper part of the downflow bed reactor; the inner cavity of the shell is divided into a fluidization zone at the upper part and a gas-solid mixing zone at the lower part by a radial sealing plate 115; the lower part of the fluidization zone is provided with a fluidization air inlet 112, and the upper part of the gas-solid mixing zone is provided with a raw oil gas inlet 116; a plurality of Venturi-type overflow pipes 113 for fluid communication between the fluidization zone and the gas-solid mixing zone are arranged on the sealing plate 115; an upwardly-open first gas distribution plate 114 is arranged between the outer wall of the upper end of the overflow pipe and the shell in the radial direction; and the fluidization air inlet 112 is arranged below the first gas distribution plate 114. The above method uses a new type of inlet distributor, which has a more uniform solid content distribution than conventional inlet distributors, can effectively improve the uniformity of catalyst particle distribution in the downflow bed reactor, is more conducive to plug flow reaction, thereby further improving the yield of low-carbon olefins and aromatic hydrocarbons and reducing the yield of coke.

[0070] In one specific embodiment of the present application, a fluidization air inlet cavity is formed between the lower surface of the first gas distribution plate 114, the upper surface of the sealing plate 115, the inner side wall of the shell, and the outer side wall of the overflow pipe 113, the fluidization air inlet cavity is in fluid communication with the fluidization air inlet 112, and is conducive to further reducing the axial disturbance to the catalyst particles. In a preferred embodiment, the fluidization air inlet is a plurality of inlets, preferably 2-6 inlets, and more preferably, a plurality of inlets are uniformly and symmetrically distributed below the first gas distribution plate.

[0071] The present application does not limit the specific structure of the first gas distribution plate, as long as it is upwardly-open and has the effect of distributing gas. In one specific embodiment of the present application, the opening rate of the first gas distribution plate 114 is 10-60%.

[0072] According to the present application, the upper end of the overflow pipe 113 can be higher than the upper surface of the first gas distribution plate 114, so that the catalyst particles overflow into the pipe after being sufficiently and uniformly fluidized, avoiding non-uniform accumulation of the catalyst; in one specific embodiment of the present application, the distance between the upper end of the overflow pipe 113 and the upper surface of the first gas distribution plate 114 can vary within a large range, for example, it can be 10-1000 mm, preferably 50-500 mm.

[0073] In order to improve the uniformity of the fluidizing gas distribution, in one embodiment of the present application, the inlet distributor further comprises a second gas distribution plate 117, which is arranged radially between the outer wall of the lower end of the overflow pipe and the inner wall of the shell; a raw oil gas inlet cavity is formed between the upper surface of the second gas distribution plate 117, the lower surface of the sealing plate 115, the inner wall of the shell and the outer wall of the overflow pipe 113, which is in fluid communication with the raw oil gas inlet 116, which is conducive to the more uniform distribution of the reaction oil gas in the gas-solid mixing zone. In a preferred embodiment of the present application, the lower end of the overflow pipe 113 is flush with the lower surface of the second gas distribution plate 117.

[0074] As shown in Figure 3 In one embodiment of the present application, the overflow pipe 113 comprises a diameter-reducing section, a straight pipe section and a diameter-increasing section in sequence, and the sealing plate 115 is sealingly connected to the outer wall of the straight pipe section of the overflow pipe 113. The diameter-reducing section has a conical funnel structure, which can form a seal and provide a driving force for the downward flow of particles, which is conducive to the downward flow of high-concentration particles, and the diameter-increasing section at the lower part allows particles to be uniformly distributed in the entire distributor at a larger radial migration speed.

[0075] In one embodiment of the present application, the maximum diameter of the diameter-increasing section and the maximum diameter of the diameter-reducing section of the overflow pipe 113 are each independently 1.1-5 times, preferably 1.5-3 times, the diameter of the straight pipe section, and the larger top diameter is conducive to the downward flow of high-concentration particles. In a preferred embodiment, the maximum diameter of the diameter-increasing section and the maximum diameter of the diameter-reducing section are the same; in a preferred embodiment, the length of the diameter-increasing section and the length of the diameter-reducing section are each independently 0.01-0.5 times, preferably 0.1-0.25 times, the height of the shell; in a preferred embodiment, the length of the diameter-increasing section and the length of the diameter-reducing section are the same.

[0076] The number of overflow pipes is not specifically limited in the present application, and in one embodiment of the present application, the overflow pipe 113 is 5-9, which is uniformly distributed in the shell 111, for example, as shown in Figure 4 The overflow pipe is arranged uniformly around the center of the sealing plate as the diffusion origin.

[0077] In one embodiment of the present application, the shell comprises a cylindrical shell 111 and a reverse-tapered shell 118 arranged in axial direction and in fluid communication, in particular, the diameter of the cylindrical shell 111 is the same as the diameter of the top of the reverse-tapered shell 118; the upper end of the cylindrical shell 111 is provided with the solid inlet, the lower end of the reverse-tapered shell 118 is provided with the mixture outlet, and the sealing plate 115 is arranged in radial direction at the bottom of the cylindrical shell 111. The inlet distributor of the present application has a structure combining the cylindrical shell and the reverse-tapered shell, wherein the cavity of the cylindrical shell serves as the fluidization zone of the particles and fluid, which can reduce the influence of wall effect and can be applied to industrial downer reactors; the reverse-tapered shell has a long conical structure, which can reduce the problem of high concentration of particles gathered at the top of the conventional downer and is more conducive to the parallel sequential reactions of catalytic cracking.

[0078] In one embodiment of the present application, the sealing plate 115 is arranged at the bottom of the reverse-tapered shell 118, preferably, the sealing plate 115 is arranged at the joint of the reverse-tapered shell 118 and the cylindrical shell 111.

[0079] According to the present application, the height-diameter ratio of the reverse-tapered shell 118 can vary in a large range, for example, it can be 2-10, preferably 3-8, the height-diameter ratio refers to the ratio of the height of the reverse-tapered shell to the maximum diameter of the top of the reverse-tapered shell, the large height-diameter ratio of the reverse-tapered shell can reduce the concentration of particles, form a dilute phase mixing zone, weaken the axial concentration and temperature gradient, and avoid excessive particle aggregation concentration at the top inlet of the downer, thereby ensuring that the particle flow in the bed presents uniform and stable plug flow, and is more conducive to the downer catalytic cracking reaction process. The inclination angle of the reverse-tapered shell can be 1°-20°, the inclination angle refers to the angle between the generatrix of the reverse-tapered shell and the axis, the height of the reverse-tapered shell can be 1-15 times, preferably 5-10 times of the diameter of the downer reactor, and the ratio of the maximum diameter to the minimum diameter of the reverse-tapered shell can vary in a large range, for example, it can be 2-6, preferably 2-4.

[0080] According to the present application, the height-diameter ratio of the cylindrical shell 111 can vary in a large range, for example, it can be 0.1-5, preferably 1-3, and the diameter of the cylindrical shell 111 can be 2-6 times, preferably 3-5 times of the diameter of the downer catalytic cracking reactor.

[0081] The present application will be further described by way of examples, but the present application is not limited in any way by the examples.

[0082] The catalyst used in the examples is prepared by the following method:

[0083] (1) Dissolve 20 kg of ammonium chloride in 1000 kg of water, add 100 kg (dry basis) of crystalline product DASY zeolite (produced by the Catalyst Factory of Qilu Petrochemical Company, 2.445-2.448 nm, rare earth content RE2O3 2.0 wt%) to the solution, and after 0.6 h of exchange at 90 °C, filter to obtain a filter cake; add 40 kg of Cu(NO3)2-6H2O dissolved in 200 kg of water, and impregnate the filter cake, dry, and then calcine at 560 °C for 2 h to obtain a copper-containing large-pore zeolite with an elemental analysis chemical composition of 0.1 Na2O-5.1 Al2O3-18.2 CuO-3.8 RE2O3-88.1 SiO2.

[0084] (2) Dissolve 20 kg of ammonium phosphate in 600 kg of water, stir until uniform, add 50 kg of MFI-structured medium-pore ZRP-1 zeolite (industrial product of the Catalyst Factory of Qilu Petrochemical Company, SiO2 / Al2O3 30) to the solution, and after 3 h of impregnation and stirring at 60 °C, filter to obtain a filter cake; dry the molecular sieve filter cake at 120 °C, and then calcine at 550 °C for 1 h to obtain a phosphorus-containing medium-pore zeolite.

[0085] (3) Slurry 80 kg of hydrous kaolin (industrial product of Suzhou Porcelain Clay Company, solid content 71.6 wt%) with 250 kg of de-cationized water, then add 50 kg of pseudo-boehmite (industrial product of Shandong Aluminum Factory, solid content 63 wt%), adjust the pH to 2-4 with hydrochloric acid, stir until uniform, and then let stand at 60-70 °C for 1 h to age while maintaining the pH at 2-4, reduce the temperature to below 60 °C, add 40 kg of aluminum sol (product of the Catalyst Factory of Qilu Petrochemical Company, Al2O3 content 21.7 wt%), and stir for 40 min to obtain a mixed slurry.

[0086] (4) Add the copper-containing large-pore zeolite prepared in step (1) (dry basis 33.8 kg) and the phosphorus-containing MFI-structured medium-pore ZRP-1 zeolite prepared in step (2) (dry basis 15.0 kg) to the mixed slurry obtained in step (3), stir until uniform, add 5 g of commercial aluminum trioxide binder, mix, and then place in a binder, add an appropriate amount of water, stir until uniform, and let stand in air for 5 h, spray dry to form, and then dry in a drying oven at 120 °C for 3 h, and then wash with ammonium dihydrogen phosphate solution (phosphorus content 1 wt%) to remove free Na + and then dry again to obtain a catalyst, which is designated as CAT-1. The composition of the catalyst is 18.9 wt% DASY zeolite, 12.0 wt% MFI-structured medium-pore zeolite, 1.9 wt% copper oxide, 0.8 wt% phosphorus pentoxide, 22.8 wt% pseudo-boehmite, 6.0 wt% aluminum sol, and the balance kaolin. The properties of the catalyst are listed in Table 2.

[0087] The DMMC-2 balancing agent used in Comparative Example 1 was purchased from the Catalyst Factory of Qilu Petrochemical Company, and mainly contained 48.1 wt% alumina, 46 wt% silica and 1.51 wt% phosphorus pentoxide. The properties are shown in Table 2.

[0088] Example 1

[0089] This example processed LCO in the system shown in Figure 1 using the LCO in Table 1 as the raw oil, using CAT-1 catalyst, and the catalyst activity was 64.

[0090] The LCO preheating temperature was 200°C, and the catalytic cracking reaction conditions were: the outlet temperature of the downflow bed reactor was 600°C, the reaction pressure was 0.4 MPa, the hydrogen partial pressure was 0.2 MPa, the weight ratio of catalyst to oil was 30, and the residence time was 0.7 seconds.

[0091] The weight ratio of steam to spent catalyst in the stripping process of the spent catalyst was 0.15.

[0092] The conditions for the coke burning regeneration were: the outlet temperature of the regenerator was 700°C, the pressure was 0.6 MPa, and the regeneration medium was air.

[0093] The regenerated aromatic catalytic cracking catalyst entered the inlet distributor 2 through the regeneration slide valve 3, was fluidized in the hydrogen-containing fluidization medium 1, and then entered the downflow bed reactor. The raw oil was preheated, entered the top end of the downflow bed reactor 5 through the oil inlet nozzle 4, mixed with the aromatic catalytic cracking catalyst, and contacted with the catalyst. The oil gas of the raw oil and the catalyst flowed downward into the downflow bed reactor 5 to perform catalytic cracking reaction. The obtained reaction mixture was separated by the rapid separator 6 at the bottom end of the downflow bed, the separated product oil gas entered the cyclone separator 7 for further separation, and finally entered the fractionation unit through the product oil gas discharge pipeline 8. The spent catalyst separated by the cyclone separator 7 entered the stripper 9 through the catalyst pipeline 11, and the spent catalyst separated by the rapid separator 6 also entered the stripper 9 to be stripped under the action of stripping steam. The regenerated catalyst after stripping entered the inlet distributor 2 of the reactor through the spent slide valve 13, was subjected to coke burning regeneration under the action of the regeneration medium 14, was separated by the cyclone separator 16 at the outlet of the riser and settled in the settler, the flue gas generated by the regeneration was separated by the cyclone separator 18 of the settler 17, was discharged from the flue gas outlet pipeline 19, and the regenerated catalyst entered the inlet distributor 2 of the reactor from the regeneration slide valve 3 to perform the next cycle. The operating conditions and the products are shown in Table 3.

[0094] Example 2

[0095] LCO was processed in the same manner as in Example 1, except that the conditions for the catalytic cracking reaction were as follows: reaction outlet temperature, 650°C; reaction pressure, 0.6 MPa; hydrogen partial pressure, 0.4 MPa; weight ratio of catalyst to oil, 30; and residence time, 0.7 seconds. The conditions for the coke burning and regeneration were as follows: outlet temperature of the regenerator, 750°C; and pressure, 0.6 MPa.

[0096] Example 3

[0097] LCO was processed in the same manner as in Example 1, except that the conditions for the catalytic cracking reaction were as follows: reaction outlet temperature, 500°C; reaction pressure, 0.4 MPa; hydrogen partial pressure, 0.2 MPa; weight ratio of catalyst to oil, 30; and residence time, 0.7 seconds.

[0098] Example 4

[0099] LCO was processed in the same manner as in Example 1, except that the CAT-1 catalytic cracking catalyst was distributed through the catalyst inlet distributor and then introduced into the upper portion of the downflow bed reactor to contact the LCO feedstock and flow down in parallel to perform the catalytic cracking reaction in the downflow bed reactor 5.

[0100] As Figure 2As shown, the catalyst inlet distributor comprises a shell, a fluidization air inlet 112, an overflow pipe 113, a first gas distribution plate 114, a sealing plate 115, a raw oil gas inlet 116, and a second gas distribution plate 117. The shell comprises an axially arranged and fluidly connected cylindrical shell 111 and a reverse tapered shell 118. The upper end of the cylindrical shell 111 is provided with a solid inlet, and the lower end of the reverse tapered shell 118 is provided with a mixture outlet. The height-diameter ratio of the reverse tapered shell 118 is 3.33, the inclination angle is 6°, the ratio of the maximum diameter to the minimum diameter of the reverse tapered shell 118 is 3, the height-diameter ratio of the cylindrical shell 111 is 2, and the diameter of the cylindrical shell 111 is 3 times the diameter of the downflow bed catalytic cracking reactor. The sealing plate 115 is arranged at the bottom of the reverse tapered shell 118, which divides the inner cavity of the shell into a fluidization zone in the upper part and a gas-solid mixing zone in the lower part. The lower part of the fluidization zone is provided with the fluidization air inlet 112, and the upper part of the gas-solid mixing zone is provided with the raw oil gas inlet 116. The sealing plate 115 is provided with 9 Venturi type overflow pipes 113 for fluidly connecting the fluidization zone and the gas-solid mixing zone. The 9 overflow pipes 113 are uniformly arranged in the shell in a radial manner. The overflow pipe 113 comprises a fluidly connected diameter-reducing section, a straight pipe section, and a diameter-increasing section in sequence. The sealing plate 115 is sealingly connected to the outer side wall of the straight pipe section of the overflow pipe 113. The maximum diameters of the diameter-reducing section and the diameter-increasing section are the same, and the heights of the diameter-reducing section and the diameter-increasing section are the same. The maximum diameter of the diameter-increasing section is 1.5 times the diameter of the straight pipe section, and the length of the diameter-increasing section is 0.15 times the height of the shell. The outer wall of the upper end of the overflow pipe is provided with the first gas distribution plate 114 with upwardly open holes between the shell in the radial direction. The outer wall of the lower end of the overflow pipe is provided with the second gas distribution plate 117 between the shell in the radial direction. The distance between the upper port of the overflow pipe 113 and the upper surface of the first gas distribution plate 114 is 100 mm, and the lower port of the overflow pipe 113 is flush with the lower surface of the second gas distribution plate 117. The fluidization air inlet 112 is arranged below the first gas distribution plate 114. A fluidization air inlet cavity is formed between the lower surface of the first gas distribution plate 114, the upper surface of the sealing plate 115, the inner side wall of the shell, and the outer side wall of the overflow pipe 113. The fluidization air inlet cavity is in fluid communication with the fluidization air inlet 112. A raw oil gas inlet cavity is formed between the upper surface of the second gas distribution plate 117, the lower surface of the sealing plate 115, the inner side wall of the shell, and the outer side wall of the overflow pipe 113. The raw oil gas inlet cavity is in fluid communication with the raw oil gas inlet 116.

[0101] Comparative Example 1

[0102] The comparative example uses the existing riser catalytic cracking process to process LCO in a small riser reactor, using a non-hydrogen operation mode, and the catalyst is DMMC-2 equilibrium catalyst, which is the same as the raw material oil in Example 1. Among them, the LCO preheating temperature is 200°C, the catalytic cracking reaction conditions are: the reaction outlet temperature is 550°C, the reaction pressure is 0.2 MPa, the catalyst to oil ratio is 10, and the residence time is 3.5 seconds. The coking regeneration conditions are: the regenerator outlet temperature is 700°C, the regenerator pressure is 0.6 MPa, the regeneration medium is air, the weight ratio of water vapor to total raw material is 0.15, and the specific operation conditions and product distribution are listed in Table 3.

[0103] Comparative Example 2

[0104] The same method as Example 1 is used to process LCO, except that the catalyst used is not the aromatization catalyst in Example 1, but the DMMC-2 equilibrium catalyst used in Comparative Example 1.

[0105] Comparative Example 3

[0106] The same method as Example 1 is used to process LCO, except that the catalytic cracking reaction is not carried out under the condition of hydrogen.

[0107] Comparative Example 4

[0108] The same method as Comparative Example 1 is used to process LCO, except that the catalytic cracking reaction is carried out in the hydrogen atmosphere of the riser reactor.

[0109] Table 1

[0110] Feed oil name LCO Density (20°C), kg / m3 871.7 Carbon content, wt% 87.37 Hydrogen content, wt% 12.6 Hydrocarbon group composition Paraffin, wt% 17.5 Total naphthene, wt% 27.2 Total aromatic, wt% 52.3 Gum, wt% 0 Total weight, wt% 100

[0111] Table 2

[0112] Catalyst No. DMMC-2 CAT-1 Micro-Reverse activity 60 64 Specific surface area, m2 / g 124 178 Pore volume, ml / g 0.26 0.36 Fractionation composition, wt% 0-40 microns 11.8 18.8 40-80 microns 53.2 58.6 > 80 microns 35 22.6

[0113] Table 3

[0114]

[0115]

[0116] As can be seen from Table 3, compared with the comparative examples, the yield of trienes (ethylene + propylene + butene) and the yield of BTX are high, and the dry gas and coke yields can be ensured to be low. The above results show that the use of a high-temperature large-dosage-oil-ratio downflow bed reactor has higher yields of low-carbon olefins and aromatic hydrocarbons. Although high-severity reaction conditions are prone to generate dry gas and coke, the downflow bed ultra-short contact time is conducive to inhibiting the generation of dry gas and coke to some extent, so as to fully exert the advantages of high severity and increase the selectivity of products.

[0117] The preferred embodiments of the present application are described in detail above with reference to the drawings, but the present application is not limited to the specific details of the above-described embodiments. Various simple modifications can be made to the technical solutions of the present application within the technical concept of the present application, and these simple modifications all belong to the protection scope of the present application.

[0118] In addition, it should be noted that each specific technical feature described in the above specific embodiments can be combined in any appropriate manner without contradiction. In order to avoid unnecessary repetition, various possible combinations are not described again by the present application.

[0119] In addition, various different embodiments of the present application can also be combined in any appropriate manner, as long as they do not deviate from the idea of the present application, and they should also be considered as disclosed by the present application.

Claims

1. A process for producing light olefins and aromatics from LCO in a downflow bed, which comprises: S1. contacting an LCO feedstock with an aromatization catalytic cracking catalyst in a downflow bed reactor in an upper portion thereof and performing a catalytic cracking reaction in a concurrent downward flow in a hydrogen atmosphere to obtain a reaction mixture; S2. separating the reaction mixture to obtain a product oil gas containing light olefins and aromatics and a spent catalyst; the aromatization catalytic cracking catalyst contains 1-60 wt% of a zeolite, 5-95 wt% of an inorganic oxide, 0-80 wt% of clay and 0-20 wt% of an active metal component, based on the total weight of the aromatization catalytic cracking catalyst on a dry basis and based on the total weight of the zeolite on a dry basis, wherein the active metal component is based on the active metal oxide; the zeolite contains 40-100 wt% of a medium pore zeolite and 0-60 wt% of a large pore zeolite, based on the total weight of the zeolite on a dry basis.

2. The method of claim 1, wherein, the conditions of the catalytic cracking reaction include a reaction temperature of 500-750℃, a reaction pressure of 0.1-2.0 MPa, a hydrogen partial pressure of 0.1-1.0 MPa, a weight ratio of catalyst to oil of 10-50, and a residence time of the LCO feedstock of 0.2-5.0 seconds; The circulation rate of the aromatization catalytic cracking catalyst in the downflow bed reactor is 100-500 kg / m 2 s, and the bed particle concentration is 1-5%.

3. The method of claim 1, wherein, the conditions of the catalytic cracking reaction include a reaction temperature of 550-650℃, a reaction pressure of 0.4-1.5 MPa, a hydrogen partial pressure of 0.2-0.8 MPa, a weight ratio of catalyst to oil of 25-35, and a residence time of the LCO feedstock of 0.5-1.5 seconds; The circulation rate of the aromatization catalytic cracking catalyst in the downflow bed reactor is 120-450 kg / m 2 s, the bed particle concentration is 1.5-4.5%.

4. The method of claim 1, wherein, the process further comprises: optionally stripping and coke burning regeneration of the spent catalyst, and returning at least part of the obtained regenerated catalyst to the downflow bed reactor of step S1 for use as the aromatization catalytic cracking catalyst.

5. The method of claim 4, wherein, the conditions of the coke burning regeneration include a temperature of 600-800℃, a pressure of 0.1-2.0 MPa, and a time of 1-10 minutes.

6. The method of claim 1, wherein, the separation is performed in a rapid separation device.

7. The method of claim 1, wherein, the process further comprises: preheating the LCO feedstock before introducing it into the downflow bed reactor, and the preheated LCO feedstock has a temperature of 100-300℃.

8. The method of claim 1, wherein, the hydrogen atmosphere contains hydrogen and / or dry gas.

9. The method of claim 8, wherein, the large pore zeolite is selected from one or more of high-silica Y, rare earth Y, rare earth hydrogen Y and ultra-stable Y, and the medium pore zeolite is selected from a ZSM series zeolite and / or a ZRP zeolite; the inorganic oxide is selected from silicon dioxide and / or diatomic aluminum oxide; the clay is selected from one or more of kaolin, montmorillonite, diatomite, halloysite, saponite, laponite, sepiolite, attapulgite, hydrotalcite and bentonite; the active metal component contains one or more of Fe, Co, Ni, Cu, Zn and rare earth metals.

10. The method of claim 1, wherein, in step S1, the contacting of the LCO feedstock with the aromatization catalytic cracking catalyst in the upper portion of the downflow bed reactor and the catalytic cracking reaction in a concurrent downward flow include: the aromatization catalytic cracking catalyst is distributed through a catalyst inlet distributor before being introduced into the upper portion of the downflow bed reactor to contact the LCO feedstock and perform the catalytic cracking reaction in a concurrent downward flow; The catalyst inlet distributor comprises a shell; a solid inlet is arranged at the upper end of the shell, and a mixture outlet is arranged at the lower end of the shell; the mixture outlet is in fluid communication with the catalyst inlet of the upper part of the downflow bed reactor; the inner cavity of the shell is divided into an upper fluidization zone and a lower gas-solid mixing zone by a radial sealing plate (115), the lower part of the fluidization zone is provided with a fluidization air inlet (112), and the upper part of the gas-solid mixing zone is provided with a raw oil gas inlet (116); a plurality of Venturi-type overflow pipes (113) for fluid communication between the fluidization zone and the gas-solid mixing zone are arranged on the sealing plate (115); a first gas distribution plate (114) with upwardly opening holes is arranged between the outer wall of the upper end of the overflow pipe and the shell in the radial direction; and the fluidization air inlet (112) is arranged below the first gas distribution plate (114).

11. The method of claim 10, wherein, A fluidization air inlet cavity is formed between the lower surface of the first gas distribution plate (114), the upper surface of the sealing plate (115), the inner side wall of the shell, and the outer side wall of the overflow pipe (113), and the fluidization air inlet cavity is in fluid communication with the fluidization air inlet (112).

12. The method of claim 10, wherein, The inlet distributor further comprises a second gas distribution plate (117) arranged between the outer wall of the lower end of the overflow pipe and the shell in the radial direction; a raw oil gas inlet cavity is formed between the upper surface of the second gas distribution plate (117), the lower surface of the sealing plate (115), the inner side wall of the shell, and the outer side wall of the overflow pipe (113), and the raw oil gas inlet cavity is in fluid communication with the raw oil gas inlet (116).

13. The method of claim 10, wherein, The overflow pipe (113) comprises, in sequence, a diameter-reducing section, a straight pipe section, and a diameter-enlarging section in fluid communication, and the sealing plate (115) is sealingly connected to the outer side wall of the straight pipe section of the overflow pipe (113).

14. The method of claim 10, wherein, The overflow pipe (113) is 5-9 in number and is uniformly distributed in the shell (111).

15. The method of claim 10, wherein, The shell comprises a cylindrical shell (111) and an inverted circular truncated cone shell (118) arranged in the axial direction and in fluid communication; the solid inlet is arranged at the upper end of the cylindrical shell (111), the mixture outlet is arranged at the lower end of the inverted circular truncated cone shell (118), and the sealing plate (115) is arranged at the bottom of the cylindrical shell (111) in the radial direction.

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