A method and device for catalyzing diesel to produce high-light aromatic naphtha

Abstract

The present invention discloses a method and device for catalytically producing high-light aromatic naphtha from diesel. A mode combining hydroprocessing and hydrocracking is adopted. The raw materials are divided into light and heavy components in a hydroprocessing pretreatment reactor, the light component passes upward through a hydrorefining catalyst bed, and the heavy component passes downward through a multi-bed reaction chamber to obtain refined diesel; the refined diesel enters a hydrocracking reactor, the heavy component countercurrently contacts with hydrogen to perform a cracking reaction, the lighter part and the light product obtained by cracking enter upward into a distillation chamber, and high-light aromatic naphtha is obtained in a side line. The method of the present invention enables the light component to quickly leave the system and no longer excessively participate in saturation and cracking. Since the product can quickly leave the reaction system, while increasing the forward reaction rate, it also eliminates the hidden danger of by-products clogging the catalyst, thereby improving the yield of the target product.

Description

Technical Field

[0001] The invention belongs to the field of hydrogenation reaction, and in particular relates to a method and a device for hydrogenation treatment of inferior diesel. Background Art

[0002] With the improvement of people's quality of life and the stricter global environmental protection requirements in the future, the proportion of new energy vehicles is getting higher and higher. Instead, the demand for gasoline and diesel is getting lower and lower, and the market competition in this part is becoming more and more brutal. In particular, the demand for diesel has gradually declined since 2016. At the same time, new large-scale oil refineries are constantly increasing, diesel production capacity is constantly increasing, diesel product surplus is constantly intensifying, and the competitive environment is deteriorating. Among diesel, catalytic diesel accounts for the largest proportion, but the quality is the worst, with an aromatic content of 60% to 85%. Generally, the excess catalytic diesel is reduced by modifying catalytic diesel to produce more gasoline or light aromatics (BTX).

[0003] The FD2G technology developed by Fushun Petrochemical Research Institute uses a two-stage reactor. The catalytic diesel first enters the hydrotreating reactor. After impurities are removed and moderate hydrogenation is performed, it is passed through a special strong cracking and weak hydrogenation light oil catalyst to control the depth of the hydrogenation reaction to avoid hydrogenation saturation of monocyclic aromatics, retain more monocyclic aromatics, and partially convert low cetane diesel into high-octane gasoline or aromatic raw materials. The research octane number (RON) of the gasoline product using this process is 90.7-94.6%, and the BTX mass fraction is 20.29%-31.57%.

[0004] The RLG technology of Sinopec Research Institute of Petroleum Processing (RIPP) uses the same two-stage process, with RN-411 as the hydrofining catalyst and RHC-100 with a molecular sieve channel structure and a gradient distribution of acidity as the hydrocracking catalyst. When this technology processes catalytic diesel with an aromatic mass fraction of 74%, the gasoline yield is 38.44% and the RON reaches 93.5.

[0005] The LTAG combined process developed by RIPP uses selective hydrogenation-catalytic cracking to produce high-octane gasoline or light aromatics, which can increase BTX production while reducing hydrogen consumption. Chiyoda Corporation of Japan has announced a non-hydrogenated aromatics production technology (FCA), which does not require hydrogenation saturation, but directly produces light aromatics such as BTX through fluidized catalytic cracking, with a BTX yield of up to 35%.

[0006] UOP's LCO-X process and Sinopec Shanghai Research Institute of Chemical Industry's PAC platform technology hydroprocessing + hydrocracking + aromatics maximization route. The LCO-X process combines the LCO-Unicracking process with the aromatics unit to convert the supersaturated alkyl cyclohexane and unconverted non-aromatic substances in the hydrocracking section into aromatics and LPG. The LCO-A technology developed by the Shanghai Institute processes 34.2% of the volume fraction of light aromatics in the LCO product.

[0007] Patent CN201810948782.X developed a catalyst using modified Y molecular sieve as a new active component. When used to process catalytic cracking diesel LCO, the BTX content in the gasoline product is 37.7%.

[0008] The above technologies all have the problem of over-hydrogenation or insufficient hydrogenation of the distillate oil in the hydrocracking unit, which makes the aromatics in LCO over-saturated to cycloalkanes, or only open the ring to dicyclic aromatics, causing the BTX yield in the product to reach a bottleneck value and cannot be further improved.

[0009] Patent CN201711470751.X first fractionates diesel raw materials into light fractions and heavy fractions, and then extracts the light fractions to obtain light aromatic fractions and light non-aromatic fractions; the heavy fractions and light non-aromatic fractions are successively subjected to hydrofining and light reforming reactors to finally obtain gasoline components with high aromatic content. This patent improves process flexibility and product quality by performing targeted hydrogenation on different components. However, the above process is relatively complicated, with high energy consumption and high investment. In addition, the hydro-reforming reactor and the hydro-refining reactor use traditional fixed-bed reactors, and there is a situation where the distillate oil is over-hydrogenated or cracked, which ultimately affects the yield of light aromatics.

[0010] In summary, although existing technicians have developed different processes and catalysts in an effort to achieve a higher BTX yield, they are limited by the inherent defects of fixed bed reactors, namely, since the cracking products may undergo cracking reactions again during the cracking process, the reaction depth is uncontrollable. In this reaction system, the selectivity of distillates such as light and heavy naphtha is limited, and the reactor will be coked and blocked. In addition, the hydrogenation reactor also has the risk of "heating up" or flooding of the catalyst bed, especially for catalysts with higher activity. The reactor temperature rises at an alarming rate, usually reaching 500°C to 600°C, and some exceed 800°C. In severe cases, the catalyst may fail and be scrapped. Summary of the invention

[0011] In view of the deficiencies in the prior art, the object of the present invention is to provide a method and device for catalytic diesel hydrogenation to produce more high-light aromatics. By adopting a multi-bed hydrogenation distillation reactor and a self-circulating system, a cascade reaction of the raw oil is achieved, and the target product is produced and shipped immediately, the reaction depth is effectively controlled, and the purpose of increasing the yield of the product high-light aromatic naphtha is achieved. At the same time, the risk of bed "flying temperature" or flooding can be eliminated, and the service life of the catalyst can be extended.

[0012] The first aspect of the present invention provides a hydroprocessing method for catalytic diesel to produce high-light aromatic naphtha. The hydroprocessing method comprises the following steps:

[0013] (1) After entering the hydrogenation pretreatment reactor, the raw diesel is directly separated into light components and heavy components. The light components are distilled and separated after passing through the refining bed layer upwards, and the naphtha fraction is extracted from the side line; the heavy components are downwardly sent into the multi-bed reaction chamber, and the upper part of the bed layer of the reaction chamber is placed with a hydrorefining catalyst, and the lower part of the bed layer is placed with a hydrocracking catalyst; the light components produced in each bed layer in the reaction chamber are eventually extracted upwards by the side line, and the unconverted oil of each bed layer is successively sent downwards into the next bed layer for hydrogenation treatment, and the obtained refined diesel is discharged from the hydrogenation pretreatment reactor;

[0014] (2) The refined diesel produced in step (1) enters the lower bed of the hydrocracking reactor, and after being dispersed by the liquid distribution component, the lighter part is carried to the middle of the reactor by the hydrogen flowing upward from the bottom, and the heavier part enters the hydrocracking reaction zone of the multiple beds downward to contact with the hydrogen countercurrent to carry out the hydrocracking reaction; the light products after hydrocracking in each bed of the lower hydrocracking reaction zone are separated from the reaction bed in time, and the unconverted oil enters the next bed for reaction, and the raw oil of each bed is the unconverted oil of the previous bed; the light products separated from each bed are gathered together and enter the middle of the reactor;

[0015] (3) The reactor described in step (1) and step (2) is provided with a stripping chamber in the middle and a distillation chamber in the upper part. The light components described in steps (1) and (2) first enter the stripping chamber in the middle for stripping, and then enter the distillation chamber in the upper part of the reactor for separation;

[0016] (4) A side product line is provided in the upper distillation chamber of the hydrocracking reactor in step (2), and a naphtha product containing high light aromatics is extracted through the side product line;

[0017] (5) A side return line is provided in the distillation chamber above the hydrocracking reactor bed in step (2) and step (3), and the heavy monocyclic aromatic hydrocarbon-rich fraction extracted from the side return line re-enters the upper reaction zone of the hydrocracking reactor in step (2) for hydrogenation reaction; the light components produced by the reaction enter the stripping chamber and the distillation chamber of the reactor, and the heavy components enter the lower hydrocracking reaction zone for further reaction, and the bottom of the reactor is the unconverted heavy components.

[0018] In step (5), at least part of the unconverted heavy components are circulated into the lower bed of the hydrocracking reaction zone through a pump for reaction.

[0019] Furthermore, the BTX-rich naphtha fraction product described in the above step (4) enters an aromatic extraction unit to separate light aromatic BTX products and light non-aromatic fraction oil.

[0020] Furthermore, the above light non-aromatic fraction oil enters into a supplementary hydrotreating reactor, and the refined oil is used as ethylene cracking feed.

[0021] Furthermore, in the above technical solution, the hydrotreating reactor and hydrocracking reactor described in step (1) and step (2) are "thin-sheet" multi-bed hydrogenation reaction and distillation integrated reactors. The reactor includes a reaction chamber, a stripping chamber and a distillation chamber. The reaction chamber is provided with a multi-layer catalyst bed, and the stripping chamber is a light component lifting channel generated by the bed reaction. The distillation chamber is a place for separating different fractions of the light component.

[0022] Furthermore, in the above technical solution, the hydrotreating catalyst bed of step (1) may be entirely composed of hydrofining catalysts, or may be a mixture of refining catalysts and other hydrotreating catalysts. Preferably, the upper part of the bed in the reaction chamber is filled with hydrofining catalysts, and the lower part of the bed is filled with hydrocracking catalysts. The volume ratio of hydrofining catalyst to hydrocracking catalyst is 1:1 to 10:1, preferably 3:1 to 5:1. The two different types of catalysts can desulfurize and denitrogenate the crude oil, and after appropriate cracking, the crude oil can be converted into a clean medium-molecular-sized distillate oil as a raw material for subsequent hydrocracking reactions.

[0023] Furthermore, in the above technical solution, the hydrorefining catalyst can be any suitable hydrorefining catalyst, preferably a hydrorefining catalyst with high desulfurization and denitrification activity and less aromatic saturation. The hydrogenation active component can be selected from one, two or more metals in Group VIB and Group VIII, such as W, Mo, Co, and Ni, and the total content of the metal oxide is 5% to 75% by weight. The carrier of the hydrogenation catalyst can be selected from alumina, amorphous silicon aluminum, silicon oxide, titanium oxide, etc. At the same time, a part of the additives, such as P, Ti, Zr, Si, B, etc., can also be added to the catalyst. The hydrorefining catalyst of the present invention can be a commercial catalyst, such as 3936, 3996, CH-20, FF-14, FF-16, FF-18, FF-24, FF-26, FF-34, FF-46, FF-56, FF-66, FH-UDS series, FZC-41, FZC-42 and other hydrogenation catalysts developed by Fushun Research Institute of Petrochemical Industry (FRIPP), HC-P, HC-T, HC-K developed by UOP, UF-210 / 220 and other hydrogenation catalysts, IFP's HR-406, HR-416, HR-448 and other hydrogenation catalysts, CLG's ICR154, ICR174, ICR178, ICR179 and other hydrogenation catalysts, Topsoe's TK-525, TK-555, TK-557 and other hydrogenation catalysts, AKZO's KF-752, KF-756, KF-757, KF-840, KF-848, KF-901, KF-907 and other hydrogenation catalysts. But not limited to the above catalysts.

[0024] Furthermore, in the above technical solution, the reaction chamber of the hydrogenation pretreatment reactor in step (1) contains at least two catalyst beds, and the ratio of the equivalent diameter of the cross-sectional area of ​​the catalyst bed to the total height of the catalyst bed (hereinafter referred to as the equivalent diameter-to-height ratio) is 0.2:1 to 20:1, preferably 3:1 to 10:1. Unless otherwise specified herein, the equivalent diameter-to-height ratio refers to the ratio of the equivalent diameter to the total height of the catalyst bed in the reactor. When there are multiple catalyst beds, the total height of the catalyst bed refers to the sum of the multiple catalyst beds.

[0025] Furthermore, in the above technical solution, the number of catalyst beds in the reaction chamber of the hydrogenation pretreatment reactor and the hydrocracking reactor in steps (1) and (2) is 2 to 30, preferably 2 to 15. The height of each bed is 50 to 500 mm, preferably 100 to 300 mm, and the liquid feedstock oil is in thin film contact with the catalyst. The catalyst height of different beds can be determined according to process requirements.

[0026] Furthermore, in the above technical solution, the liquid feed inlet of the hydrogenation pretreatment reactor is arranged at the top of the reaction chamber, and the hydrogen feed inlet is arranged at the bottom of each bed layer in each part of the reaction chamber.

[0027] Furthermore, in the above technical solution, the distillation chamber of the hydrogenation pretreatment reactor is provided with a reaction zone, and the number of beds is 1 to 3, and refining catalysts with different functions can be provided. After the light component crude oil passes through the refining catalyst, the sulfides and nitrides in the light component are effectively removed, and part of the dicyclic aromatics and tricyclic aromatics are saturated into monocyclic aromatics.

[0028] Furthermore, in the above technical solution, a light fraction extraction line is arranged at the upper part of the reaction zone of the distillation chamber of the hydrogenation pretreatment reactor, and the extraction line temperature is 170°C to 300°C, preferably 180°C to 280°C, and the distillation range is 60°C to 240°C.

[0029] Furthermore, in the above technical solution, a condensation reflux line is provided at the top of the distillation chamber of the hydrogenation pretreatment reactor, and the condensation temperature is 50°C to 70°C, preferably 55°C to 65°C. The condensate is fully refluxed.

[0030] In the above technical scheme, the heavy components separated by entering the hydrogenation pretreatment reactor are first desulfurized and denitrogenated by the refined catalyst in the upper bed of the reaction chamber, and some polycyclic aromatic hydrocarbons are hydrogenated and saturated. The refined distillate oil then enters the cracking catalyst bed in the lower part of the reaction chamber to further hydrogenate and saturate the distillate oil, open the chain and open the ring. The hydrogen sulfide, ammonia and monocyclic light hydrocarbon components generated during the reaction are quickly taken away from the bed by hydrogen. The heavy components that have not opened the ring or opened the long chain enter the next bed, so that the reaction process is carried out in stages, and the light aromatic components in the product are retained to the maximum extent.

[0031] Furthermore, the operating conditions of the hydrogenation pretreatment process described in step (1) are as follows: the average reaction temperature is 200-450°C, the reaction pressure is 2.5-20 MPa, the hydrogen-oil volume ratio is 100:1-3000:1. The feedstock oil flow rate and the reactor catalyst volume ratio are 0.1-10.0h -1 The preferred operating conditions are: average reaction temperature of 260°C to 420°C, reaction pressure of 4 to 12 MPa, hydrogen to oil volume ratio of 200:1 to 1500:1, and reflux ratio of 1.5 to 3.0.

[0032] Furthermore, in the above technical solution, the refined oil at the bottom of the hydrotreating pretreatment reactor enters the lower bed of the reaction chamber of the hydrocracking reactor. A cascade reaction is carried out in the multi-bed reaction chamber of the hydrocracking reactor, the light component enters the distillation chamber through the stripping chamber, and the heavy component finally circulates into the hydrocracking reaction chamber.

[0033] Furthermore, in the above technical solution, the distillation chamber of the hydrogenation pretreatment reactor is provided with a light fraction extracted from the upper part of the reaction zone and is drawn into the lower part of the distiller of the hydrocracking reactor, and is separated into light components and heavy components at the top of the hydrocracking reaction chamber. The light components enter the distillation chamber of the hydrocracking reactor upward, and the heavy components enter the reaction chamber of the hydrocracking reactor downward.

[0034] Furthermore, in the above technical solution, two extraction lines are arranged in the distillation chamber of the hydrocracking reactor. The first line is an extraction line for products containing high light aromatic fractions. The extraction temperature is 150°C to 185°C, preferably 160°C to 170°C. The second line is a side return line, and the extraction temperature is 170°C to 300°C, preferably 180°C to 280°C. The fraction in the side return line is a heavier fraction separated from the distillation chamber, containing long-chain monocyclic aromatic hydrocarbons such as alkylbenzenes, a small amount of bicyclic aromatic hydrocarbons and tricyclic aromatic hydrocarbons, which are returned to the catalyst bed of the reaction chamber for chain opening, saturation and ring opening again.

[0035] Furthermore, in the above technical solution, a reflux line is provided at the top of the distillation chamber of the hydrocracking reactor, the condensation temperature of the reflux line is 50°C to 70°C, preferably 55°C to 65°C, and the reflux ratio is 1.1 to 10.0, preferably 1.5 to 3.0.

[0036] Furthermore, in the above technical solution, the average reaction temperature of the upper reaction zone of the hydrocracking reactor described in step (5) is 210-430°C, the reaction pressure is 3-20 MPa, and the hydrogen-oil volume ratio is 100:1-3500:1. The ratio of the feedstock oil flow rate to the reactor catalyst volume is 0.1-10.0h -1 The preferred operating conditions are: average reaction temperature of 250°C to 390°C, reaction pressure of 4 to 18 MPa, and hydrogen to oil volume ratio of 200:1 to 1500:1.

[0037] Furthermore, in the above technical solution, the average reaction temperature of the lower reaction zone of the hydrocracking reactor described in step (2) is 230-430°C, the reaction pressure is 3-20MPa, and the hydrogen-oil volume ratio is 100:1-4000:1. The ratio of the feedstock oil flow rate to the reactor catalyst volume is 0.1-10.0h-1. The preferred operating conditions are: the average reaction temperature is 280-420°C, the reaction pressure is 4-18MPa, and the hydrogen-oil volume ratio is 200:1-1800:1.

[0038] The second aspect of the present invention also provides a hydroprocessing device for catalyzing diesel to produce high-light aromatic naphtha. The device comprises:

[0039] (1) A hydrogenation pretreatment reactor, which is used to hydrogenate the inferior diesel raw material. It is a multi-bed hydrogenation distillation reactor, in which reaction beds are arranged above and below the feed inlet;

[0040] (2) The hydrocracking reactor is a multi-bed hydrodistillation reactor with two liquid feed ports, an upper extraction line for extracting the product, and a lower reaction zone with two parts;

[0041] (3) The multi-bed hydrodistillation reactor of step (1) and step (2) comprises: a reaction chamber, in which a catalyst bed is arranged in an annular manner, and the liquid feedstock oil and the hydrogen feed are in countercurrent contact in the catalyst bed; a stripping chamber, which is located in the middle of the hydrodistillation reactor and is a gas phase channel that runs through the reactor from top to bottom, and the gas phase product after the hydrogenation reaction of each catalyst bed enters the gas phase channel under the action of stripping; and a distillation chamber, which is located in the upper part of the hydrodistillation reactor, and the gas phase product is distilled and separated in the distillation chamber, and an extraction line is provided, and the extraction line can return to the reaction bed.

[0042] Furthermore, in the above technical solution, the ratio of the equivalent diameter of the cross-sectional area of ​​the catalyst bed to the total height of the catalyst bed (hereinafter referred to as the equivalent diameter-to-height ratio) is 0.2:1 to 10:1, preferably 3:1 to 10:1. The reactor bed described in the present invention is preferably of equal diameter, that is, the cross-sectional area at different positions within the entire catalyst bed is the same. The cross-sectional area of ​​the catalyst bed is usually the same as the cross-sectional area of ​​the reaction chamber in the reactor, and the cross section refers to the top view cross section, that is, the cross section perpendicular to the vertical line in the reaction chamber. If there are differences in the cross-sectional area of ​​the reactor within the catalyst bed height range, the cross-sectional area here refers to the average value of the cross-sectional area of ​​the catalyst bed or the cross-sectional area of ​​the reaction chamber at various locations within the catalyst bed range.

[0043] Furthermore, in the above technical solution, the reaction chamber is preferably a cylindrical storage tank, the axial direction of which is arranged along the longitudinal direction, or a horizontal storage tank, the axial direction of which is arranged along the transverse direction, and the two ends of the horizontal storage tank are provided with heads.

[0044] The catalyst bed in the present invention can also greatly increase the material flux passing through the bed at a higher diameter-to-height ratio, while reducing the accumulation of heat in the catalyst bed. At the same time, the hydrogen sulfide and ammonia generated in each sub-bed are quickly taken out of the reactor, and the blockage phenomenon caused by the by-products staying in the bed for too long, which is the case in traditional reactors, will not occur.

[0045] Furthermore, in the above technical solution, a conical gas-liquid guide plate may be provided in the stripping chamber, and the number and position of the gas-liquid guide plate match the catalyst bed. A through hole is provided on the gas-liquid guide plate, a cofferdam is provided around the through hole, and a cap is preferably provided above the through hole. The opening rate of the gas-liquid guide plate is 5% to 90%, the diameter of the through hole is 3 to 100 mm, and the height of the cofferdam is 1 to 30 mm. The gas phase products after the hydrogenation reaction rise from the pores and converge into the distillation chamber, and part of the heavy component liquid phase products drip onto the gas-liquid guide plate and flow back to the reaction chamber.

[0046] Furthermore, in the above technical solution, 1 to 3 liquid feed distributors are arranged in the reaction chamber, which are arranged above different catalyst beds in the upper part of the reaction zone, preferably in the form of an annular tube or an annular belt; hydrogen feed distributors are arranged at the bottom of each catalyst bed, preferably in the form of an annular tube or an annular belt. However, the present invention is not limited to the preferred gas and liquid feed distributors.

[0047] Further, in the above technical solution, a gas guide unit may be provided between the hydrogen feed distributor and the bottom of the catalyst bed, and the gas guide unit includes: a guide plate, which is provided at the bottom of the catalyst bed and is used to receive the liquid phase product after the hydrogenation reaction of the catalyst bed; a guide pipe, which is evenly arranged on the guide plate and guides the hydrogen feed upward to the catalyst bed, and a cap is provided on the top of the guide pipe; an overflow weir, which is close to one side of the distillation chamber, is used to overflow the liquid phase product on the guide plate to the next catalyst bed. Preferably, but not restrictively, the opening rate of the guide plate is 5% to 90%, the diameter of the guide hole is 3mm to 100mm, the height of the guide pipe is 10mm to 300mm, and the height difference between the overflow weir and the guide pipe is 5 to 250mm. Keep the cap just above the guide hole.

[0048] Furthermore, in the above technical solution, a liquid redistribution plate with an annular shape and evenly provided through holes may be provided below the overflow weir. The liquid redistribution plate may include: an overflow ring, which is provided around the top of the through hole, and the height of the overflow ring is 1 mm to 50 mm, and is used to evenly guide the liquid phase product to the next catalyst bed; a baffle, which is provided on the side away from the distillation chamber, and the upper edge of the baffle is higher than the overflow ring.

[0049] Furthermore, in the above technical solution, the distillation chamber may include a mixing section and a separation section, and the separation section may be provided with fillers or tower plates.

[0050] Furthermore, in the above technical solution, the ratio of the diameter of the upper distillation chamber to the diameter (or maximum equivalent diameter) of the lower reaction chamber is 1:1.2 to 1:50, preferably 1:2 to 1:10. In the present invention, the diameter of the upper distillation chamber of the hydrodistillation reactor is reduced, so that the light fraction load under high pressure is fully matched with the tower plate, the tower plate separation efficiency is high, and it has full substitutability for the fractionation tower.

[0051] Furthermore, in the above technical solution, a tower plate or filler is placed in the distillation chamber, and the above filler or tower plate is a conventional form in the art, such as a filler can be selected from one or more random packings such as ball rings, Raschig rings, rectangular saddle rings, saddles, open-hole ring types, semi-rings, step rings, double arcs, Haier rings, conjugated rings, flat rings, and wreaths, and the filler can also be selected from metal or ceramic corrugated packing. The tower plate is one or more of the tower plates with downcomers such as bubble plates, sieve plates, float valve plates, mesh plates, tongue-shaped plates, guide sieve plates, and multi-downcomer tower plates, and can also be tower plates without downcomers such as flow-through sieve plates and flow-through corrugated plates. High-efficiency tower plates such as float valve plates and sieve plates are preferred.

[0052] Furthermore, in the above technical solution, the distillation chamber plates or packing areas can be set up as reaction zones according to process requirements, in which corresponding catalysts are placed.

[0053] Furthermore, in the above technical solution, the hydrodistillation reactor further comprises: a reboiler, one end of which is connected to the outlet of the heavy oil bunker and the other end is connected to the hydrogen distribution chamber. The reboiler is used to keep the temperature of the heavy oil bunker at the temperature required by the reaction bed.

[0054] Furthermore, in the above technical scheme, it is preferred to open 1 to 4 side lines in the separation section of the distillation chamber to extract the required components. At the first side line or the top of the distillation chamber, the extraction line condensation temperature is 50°C to 70°C, preferably 55°C to 65°C; the reflux ratio is 1.1 to 10.0, preferably 1.5 to 3.0, and all the liquid hydrocarbons are refluxed after condensation, and the hydrogen-rich non-condensable gas returns to the inlet of the device for recycling. The extraction temperature of the second side line or the side line product line is 150°C to 185°C, preferably 160°C to 170°C, and the component extracted from the side line is a heavy naphtha fraction. The extraction temperature of the third side line or the side line return line is 180 to 250°C, preferably 180 to 230°C, and the component extracted from the side line is a jet fuel fraction. The extraction temperature of the third side line or the side line return line is 200 to 370°C, preferably 210 to 350°C, and the component extracted from the side line is a diesel fraction. The present invention can extract distillate oils of different distillation ranges according to process requirements and is not limited to the temperature of the side line.

[0055] Furthermore, in the hydrodistillation reactor, the number of catalyst beds is 1 to 30, preferably 2 to 15. The diameter-to-height ratio may be 0.2:1 to 20:1, preferably 3:1 to 10:1. The liquid feedstock oil is in thin-film contact with the catalyst.

[0056] The hydrogenation reactor of the present invention adopts a multi-bed hydrogenation distillation reactor to split the light and heavy components and separate the components.

[0057] In the hydroprocessing reactor, the separated light components are carried by hydrogen and move upward into the distillation chamber for separation. A hydrorefining bed is set within the appropriate temperature fraction range. The light component fraction oil undergoes desulfurization, denitrogenation, olefin saturation and other reactions. The refined light oil aromatics contain more than 70% of monocyclic aromatics, which are all extracted from the top, and the rest are dicyclic aromatics. Heavy oil aromatics contain more than 90% of dicyclic aromatics and polycyclic aromatics.

[0058] In the hydrocracking reactor, the aromatic-rich light oil entering the distillation chamber is separated in the separation section of the distillation chamber. The monocyclic aromatic hydrocarbons with longer side chains, such as alkylbenzenes, are drawn out from the side line and returned to the catalyst bed on the upper layer of the reaction chamber to react with hydrogen, resulting in a chain breaking reaction. The generated light aromatic-rich hydrocarbons leave the bed and enter the distiller again through the stripping chamber for separation. After the heavy fraction oil enters the cracking reaction bed, the polycyclic aromatic hydrocarbons open the ring to produce monocyclic aromatic hydrocarbons, which leave the reaction bed in time and are separated through the distillation chamber. The ring-opened polycyclic aromatic hydrocarbons and heavy fraction oil then enter the next bed layer. Similarly, the light aromatic fraction generated by the reaction leaves the reaction system in time and is separated out, and the heavy components of the reaction react in turn in the multiple beds below. In this way, not much of the reaction product is removed, and the cracking reaction rate is accelerated. Since the distillation chamber only separates the light aromatic products, the unseparated fractions in the distillation chamber enter the catalyst bed layer again to react. Therefore, the present invention is more selective for light aromatics and can increase the yield of light aromatics. The load of the catalyst bed is also greatly reduced. According to the test results, the load of each catalyst bed is reduced by 5% to 20%.

[0059] The catalyst bed of the multi-bed hydrocracking reactor in the present invention can also greatly increase the material flux passing through the bed at a higher diameter-to-height ratio, while reducing the accumulation of heat in the catalyst bed. At the same time, the generated by-products such as hydrogen sulfide and ammonia are quickly taken out of the reactor, and the blockage phenomenon caused by the by-products staying in the bed for too long, which is the case in traditional reactors, will not occur.

[0060] In the present invention, the heat generated by the hydrogenation reaction can be fully taken away in time from the bed through the hydrogen stripping and light component gasification process, and can be provided for the heat used for the gasification and distillation of the light component, with obvious energy-saving effect.

[0061] The present invention can disperse the liquid after hydrogenation treatment into suitable small droplets through the design and arrangement of the liquid distribution component, and with the help of the stripping effect of hydrogen, the lighter part can be directly taken out and no longer enter the hydrocracking reaction zone to participate in the hydrocracking reaction. At the same time, the design of the liquid distribution component allows the heavy components entering the hydrocracking reaction zone to be distributed more evenly, solving the problem of poor contact of reactants on the catalyst bed under a large diameter-to-height ratio in traditional reactors. Using the hydrogenation reactor of the present invention, under the same process conditions and product index requirements, the temperature rise of the reaction bed is controlled to be smaller and the allowable feed load is larger. The improvement of the pore structure can improve the liquid flooding characteristics (no liquid flooding for a long time) and ensure good mass transfer performance. Finally, the product properties have good controllability.

[0062] Compared with the prior art, the present invention has the following beneficial effects:

[0063] 1. The octane number or light aromatic content of light naphtha in the present invention is higher. Taking inferior catalytic diesel as raw material, the naphtha yield is 3 to 10 percentage points higher than that of the conventional fixed bed two-stage process. When the aromatic content in the catalytic diesel raw material is more than 80%, the yield of light aromatics (C6 to C8) in the naphtha product reaches more than 40%.

[0064] 2. The present invention can improve the selectivity of light aromatics in naphtha. The present invention uses the separator or distillation chamber to rapidly remove the hydrogenation target product, light naphtha, from the reaction system by hydrogen, and the target product will not be oversaturated or cracked. The target product of the required fraction can be separated by optimizing the process parameters of the separation section, thereby increasing the reaction rate and greatly improving the selectivity of light naphtha with a high aromatic content.

[0065] 3. In the present invention, the catalyst life is longer. In the present invention, the hydrogenation reaction byproducts such as hydrogen sulfide, ammonia, and small molecule olefins have a very short residence time in the reaction bed, and the reaction heat is also removed from the bed in time, thus basically eliminating the coking and clogging problems of the bed.

[0066] 4. The present invention couples the reaction and separation twice. After the raw oil passes through the reactor of the present invention, light distillate oil and heavy distillate oil are directly obtained. There is no need to set up a special distillation tower for separation, which saves investment and energy consumption. At the same time, the heat generated by the hydrogenation reaction process is exactly what the distillation process needs, which also saves heat energy. The distillation process is mostly used at a relatively low pressure or vacuum. The distillation process under the conditions of medium and high pressure hydrogenation of 8MPaG to 16MPaG is generally considered to be unable to achieve the expected effect. However, through the design of the present invention, although the light hydrocarbons after cracking under high pressure are in liquid form, under the action of the separation section and the stripping chamber, these light hydrocarbons can still easily separate from the reaction system to achieve the expected separation effect. BRIEF DESCRIPTION OF THE DRAWINGS

[0067] Figure 1 The present invention is a schematic diagram of the process flow of catalytic diesel hydrogenation to produce high-light aromatic naphtha.

[0068] Figure 2 Schematic diagrams of two structures of a liquid distributor and a gas distributor according to one embodiment of the present invention.

[0069] Figure 3 It is a schematic diagram of a catalyst bed structure according to one embodiment of the present invention.

[0070] Figure 4 It is a schematic diagram of the structure of a gas-liquid redistributor according to one embodiment of the present invention.

[0071] Figure 5 It is a schematic diagram of the structure of a gas-liquid guide plate according to one embodiment of the present invention.

[0072] Among them: 1-raw oil inlet; 2-new hydrogen; 3-hot raw oil; 4-circulating hydrogen; 5-hot hydrogen; 6-hydrogenation reactor; 7-hydrogenation reactor upper reaction zone; 8-refined oil and unconverted oil; 9-hydrocracking reactor; 10, 35-reaction chamber; 11, 33-distillation chamber; 12-cracking reactor upper liquid distributor; 13-hydrogen; 14-cracking reactor upper catalyst bed; 15-cracking reactor lower liquid distributor; 16-side product; 17, 34-liquid separator; 18, 36-condenser; 19-refined light oil; 20-desulfurization and deamination equipment; 21-side return line; 22-tail oil; 23-circulating oil pump; 24-circulating hydrogen compressor; 25-side return pump; 26-heating furnace; 27-cracking reactor Catalyst bed at the bottom of the reactor; 28, 32 - stripping chamber; 29 - refining bed of hydroprocessing reactor; 30 - cracking bed of hydroprocessing reactor; 101 - feed pipe; 102 - material distribution pipe; 103 - eyelet; 104, 105 - material distributor; 201 - catalyst bed; 202 - gas guide unit; 203 - gas distributor; 204 - gas distribution chamber; 205 - liquid redistribution plate; 206 - guide hole; 207 - guide groove; 208 - overflow weir; 209 - outer wall of guide groove; 210 - inner wall of guide groove; 211 - gas-liquid guide plate; 301, 303 - eyelet; 302 - overflow ring; 304 - baffle; 401, 501 - cap; 402 - air hole; 403 - annular cofferdam; 502 - guide pipe. DETAILED DESCRIPTION

[0073] The specific implementation modes of the present invention are described in detail below in conjunction with the accompanying drawings, but it should be understood that the protection scope of the present invention is not limited by the specific implementation modes.

[0074] Unless explicitly stated otherwise, throughout the specification and claims, the term “comprise” or variations such as “include” or “comprising”, etc., will be understood to include the stated elements or components but not to exclude other elements or components.

[0075] In this document, for the convenience of description, spatial relative terms such as "below", "below", "down", "above", "above", "upper", etc. may be used to describe the relationship between one element or feature and another element or feature in the accompanying drawings. It should be understood that the spatial relative terms are intended to include different directions of objects in use or operation in addition to the directions drawn in the drawings. For example, if the object in the figure is turned over, the elements described as being "below" or "below" other elements or features will be oriented "above" the elements or features. Therefore, the exemplary term "below" can include both below and above. Objects may also have other orientations (rotated 90 degrees or other orientations) and the spatial relative terms used in this document should be interpreted accordingly.

[0076] In this document, the terms "first", "second", etc. are used to distinguish two different elements or parts, and are not used to limit a specific position or relative relationship. In other words, in some embodiments, the terms "first", "second", etc. can also be interchangeable.

[0077] like Figure 1The figure shows a process flow diagram of a method and device for catalytic diesel production of high-light aromatic naphtha of the present invention. The catalytic diesel is fed into the heating furnace 26 from the crude oil inlet 1, and after being heated to 200-400°C, it enters the liquid distributor of the reaction chamber 35 of the hydroprocessing reactor 6. The hydrogen 2 is heated by the heating furnace 26 and enters the bottom of each bed layer of the reaction chamber 35 of the hydroprocessing reactor 6. The catalytic diesel becomes fine droplets after passing through the liquid distributor of the reaction chamber 35. After mixing and stripping with the hydrogen from the bottom upward, the light components in the crude oil enter the distillation chamber 33 upward, and are desulfurized and denitrogenated under the action of the refining catalyst in the distillation chamber, and the olefins are saturated. The refined light components are separated by rectification in the separation section of the distillation chamber to obtain light refined oil 19 at 170-300°C, and the refined oil enters the distillation chamber 11 of the hydrocracking reactor 9. After the light components at the top of the distillation chamber are condensed at 50°C to 70°C, all liquid hydrocarbons are refluxed, and the non-condensable gas is desulfurized by the desulfurization facility 20 and returned to the inlet by the hydrogen compressor 24. The heavy components in the raw oil enter the refined catalyst bed 29 (schematic diagram, not limited to 3 layers) downward, and after reacting with the countercurrent hydrogen in the catalyst bed, the generated hydrogen sulfide, ammonia and small molecular hydrocarbons are separated from the bed with the hydrogen and enter the stripping chamber 32, and finally enter the distillation chamber 33 for reaction and separation. The refined diesel and the circulating oil transported by the circulating oil pump 23 are combined into the hydrocracking raw oil 8 and enter the liquid distributor 15 above the lower bed 27 of the hydrocracking reactor 9 reactor 10, the light components enter the stripping chamber 28, the heavy components enter the bed 27 downward for reaction, and the hydrogen 13 enters from the bottom of each bed and enters the catalyst bed 27 evenly under the action of the gas distribution chamber 204 and the gas distributor 203. In each bed, the heavy components in the raw diesel are saturated and cracked with olefins under the action of hydrogen and catalyst to generate small molecular hydrocarbons. Under the hydrogen, the small molecular hydrocarbons and part of the large molecular hydrocarbons are separated from the bed and enter the stripping chamber 28. Part of the heavy components carried by hydrogen enter the stripping chamber 28, fall into the gas-liquid guide plate 211 under the action of gravity, and then re-enter the corresponding catalyst bed through the guide groove 207 for reaction. The light components entering the stripping chamber 28 in each bed enter the distillation chamber 11 upward, and are separated into light fractions of different temperatures under the action of the high-efficiency trays or fillers in the distillation chamber. The side line 16 finally obtains the target product. The side line 21 enters the catalyst bed 14 at the top of the reaction chamber through the side line return pump 25, and the light components generated after hydrogenation reaction with hydrogen enter the stripping chamber 28, and then enter the distillation chamber upward to separate the distillate oil of 150°C to 185°C again to enter the side line 16 to obtain the product. The separated heavier fraction oil 21 with a temperature of 170°C to 300°C is sent to the liquid distributor 12 at the top of the reaction chamber 10 through the return line pump 25, and then the light and heavy fractions are separated, and the heavy fractions are subjected to ring opening and side chain opening reactions. The non-condensable gas at the top of the distillation chamber is further cooled by the condenser 18, and then the liquid hydrocarbons are refluxed after gas-liquid separation in the liquid separator 17. The gas enters the desulfurization and deamination equipment 20 for purification and then is recycled by the circulating hydrogen compressor 24.The heavy diesel fraction at the bottom of the reaction chamber 10 is sent back to the inlet of the multi-bed hydrodistillation reactor 9 by the circulating oil pump 23 or extracted as a chemical raw material 22.

[0078] The following examples further illustrate a method for hydrogenating low-quality diesel oil that can improve the selectivity and yield of naphtha and the effects of the invention. In the examples, detailed implementation methods and specific operation processes are given based on the technical solution of the present invention, but the protection scope of the present invention is not limited to the following specific examples.

[0079] The experimental methods in the following examples, unless otherwise specified, are conventional methods in the art.

[0080] Example 1

[0081] Using the present invention Figure 1 In the flow chart shown, the feedstock oil is catalytic diesel, and its properties are shown in Table 2. After the feedstock oil and hydrogen are heated to 360°C in the heating furnace, they enter the hydroprocessing reactor together. The number of beds is 10, and the upper 7 layers are filled with refined catalyst A, and the lower 3 layers are filled with cracking catalyst B. The specific catalyst properties are shown in Table 1. The extraction temperature of refined light oil 19 is 280°C, the condensation temperature of condenser 29 is 60°C, and the upper reaction zone 7 is filled with refined catalyst A. The feedstock oil 8 mixed with refined oil and circulating oil at the bottom of the reaction chamber 35 enters the lower bed of the hydrocracking reactor 9, the side line 19. The reaction chamber 10 of the hydrocracking reactor has a total of 15 beds, 10 lower beds and 5 upper beds. The extraction temperature of the side line return line 21 is 260°C. The bed of the hydrocracking reactor is filled with catalyst C, and the catalyst bed equivalent diameter-to-height ratio is 4:1. The opening rate of the gas guide unit below the catalyst layer is 70%, and the opening rate of the gas-liquid guide plate in the stripping chamber is 80%. The liquid raw material of the multi-bed hydrodistillation reactor is first separated into light and heavy components in the reaction chamber through the liquid distribution component. The light component enters the distillation chamber upward and is finely separated into different fractions according to different temperature distributions. The heavier fraction after fine separation enters the catalyst bed again under the action of the gas-liquid guide plate in the stripping chamber to react until the target product is generated. The heavy fraction that is not sufficiently cracked flows out from the bottom of the reaction chamber and passes through the circulation pump and all enters the inlet of the multi-bed hydrodistillation reactor as circulating oil.

[0082] Example 2

[0083] The difference between this embodiment and embodiment 1 is that the catalyst bed equivalent diameter to height ratio in the hydrocracking reactor is 5:1, and the number of catalyst beds is 20. The other conditions are the same as those in embodiment 1.

[0084] Example 3

[0085] The present embodiment is different from the embodiment 1 in that the extraction temperature of the refined light oil 19 is 240°C, the extraction temperature of the side return line 21 is 220°C, and the catalyst bed equivalent diameter-to-height ratio is 5:1. The other conditions are the same as those of the embodiment 1.

[0086] The catalyst properties of the above-mentioned embodiments and comparative examples are shown in Table 1, the raw oil properties are shown in Table 2, the process conditions are shown in Table 3, and the product distribution of the embodiments and comparative examples is shown in Table 4. The temperature rise of the simulated hydrocracking bed of the embodiments and comparative examples is shown in Table 5.

[0087] Comparative Example 1

[0088] A conventional two-stage hydrogenation process, i.e., refining + cracking process, is adopted. Both the refining and cracking reactors adopt a reaction process in which the raw material and hydrogen flow in parallel from top to bottom. The cracked distillate oil is separated in a separation tower and then sent out as a product. In the comparative example, the refining reactor is filled with catalyst A and catalyst B in a volume ratio of 3:1, and the cracking reactor is filled with catalyst D. The tail oil full circulation process is adopted, and the tail oil is intermittently thrown out. The remaining process conditions are the same as those in Example 1.

[0089] Table 1 Catalyst properties

[0090]

[0091] Table 2 Properties of crude oil

[0092]

[0093] Table 3 Process conditions of the embodiments and comparative examples

[0094]

[0095] Table 4 Product distribution of embodiments and comparative examples

[0096]

[0097] From the results in Table 4, it can be seen that the yield and selectivity data of C6-C8 are significantly higher than those of Comparative Example 1 by using the method of the present invention.

[0098] The laboratory simulated and calculated the bed reaction temperature distribution of the embodiment and the comparative example by using ansys 19.0 version software. The simulation conditions were input according to the actual data of the embodiment and the comparative example. The simulation results show that the center temperature of the traditional fixed bed is the highest, and the temperature change is normally distributed from the inlet end to the outlet end. The bed temperature of the reactor of the present invention is relatively uniform. The simulated temperature rise change of the bed is shown in Table 5.

[0099] Table 5 Simulated bed temperature rise changes

[0100]

[0101] It can be seen from the results in Table 5 that the temperature difference of the catalyst bed in Examples 1-2 of the present invention is significantly lower than that in Comparative Example 1, from 20.1°C of the traditional fixed bed to 0.4°C. The difference between the average temperature of the examples and the controlled temperature is the smallest, indicating that the reactor of the present invention has eliminated the overheating phenomenon of the hydrocracking reaction.

Claims

1. A hydroprocessing device for catalyzing diesel to produce high-light aromatic naphtha, the device comprising: The hydrogenation pretreatment reactor is used to hydrogenate the inferior diesel fuel as a raw material. It is a multi-bed hydrogenation distillation reactor, and reaction beds are arranged at the upper and lower parts of the feed inlet. The refined diesel fuel from the hydrogenation pretreatment reactor enters the hydrocracking reactor. The hydrocracking reactor is a multi-bed hydrodistillation reactor with two liquid feed ports, an upper extraction line to extract the product, and a lower reaction zone with two parts; The multi-bed hydrodistillation reactor of the hydropretreatment reactor and the hydrocracking reactor comprises: a reaction chamber, in which a catalyst bed arranged in an annular manner is arranged, and the liquid-phase raw oil and the hydrogen feed are in reverse contact in the catalyst bed; a stripping chamber, which is located in the middle of the multi-bed hydrodistillation reactor and is a gas phase channel that runs through the reactor from top to bottom, and the gas phase product after the hydrogenation reaction of each catalyst bed enters the gas phase channel under the action of stripping; a distillation chamber, which is arranged at the upper part of the multi-bed hydrodistillation reactor, and the gas phase product is distilled and separated in the distillation chamber, and an extraction line is arranged, and the extraction line can return to the reaction bed; the liquid-phase raw oil is low-quality diesel, or a mixed component of diesel refined in the hydropretreatment reactor and a heavy monocyclic aromatic hydrocarbon-rich fraction extracted from the side return line of the distillation chamber of the hydrocracking reactor; the heavy monocyclic aromatic hydrocarbons are monocyclic aromatic hydrocarbons with long chains; The reaction chamber of the hydrogenation pretreatment reactor contains at least two catalyst beds, and the ratio of the cross-sectional area equivalent diameter of the catalyst bed of the multi-bed hydrodistillation reactor to the total height of the catalyst bed is 3 to 10; The reaction chamber is provided with 1 to 3 liquid phase feed distributors, which are arranged above different catalyst beds at the upper part of the reaction zone; and a hydrogen feed distributor, which is arranged at the lower part of each catalyst bed.

2. The hydroprocessing device according to claim 1, characterized in that: The reaction chamber is a cylindrical storage tank, and its axial direction is arranged along the longitudinal direction; or, the reaction chamber is a horizontal storage tank, and its axial direction is arranged along the transverse direction, and heads are provided at both ends of the horizontal storage tank.

3. The hydroprocessing device according to claim 1, characterized in that: The stripping chamber is provided with a conical gas-liquid guide plate, the number and position of which match the catalyst bed; the gas-liquid guide plate is provided with a through hole, a cofferdam is provided around the through hole, and a cap is provided above the through hole.

4. The hydroprocessing device according to claim 1, characterized in that: A gas guide unit is provided between the hydrogen feed distributor and the bottom of the catalyst bed, and the gas guide unit comprises: a guide plate, which is provided at the bottom of the catalyst bed and is used to receive liquid products after hydrogenation reaction from the catalyst bed; a guide pipe, which is evenly arranged on the guide plate and guides the hydrogen feed upward to the catalyst bed, and a cap is provided on the top of the guide pipe; and an overflow weir, which is close to one side of the stripping chamber and is used to overflow the liquid products on the guide plate to the next catalyst bed.

5. The hydroprocessing device according to claim 4, characterized in that: A liquid redistribution plate having an annular shape and evenly provided through holes is provided below the overflow weir; the liquid redistribution plate comprises: an overflow ring, which is provided around the top of the through hole and is used to evenly guide the liquid phase product to the next catalyst bed layer; a baffle, which is provided on the side away from the distillation chamber, and the upper edge of the baffle is higher than the overflow ring.

6. The hydroprocessing device according to claim 1, characterized in that: The distillation chamber comprises a mixing section and a separation section, and packing or a tower plate is arranged in the separation section.

7. The hydroprocessing device according to claim 1, characterized in that: The ratio of the diameter of the distillation chamber to the diameter of the reaction chamber is 1:1.2 to 1:

50.

8. The hydroprocessing device according to claim 1, characterized in that: A tower plate or filler is placed in the distillation chamber.

9. The hydroprocessing device according to claim 8, characterized in that: The trays or packing areas are provided with corresponding catalysts.

10. The hydroprocessing device according to claim 6, characterized in that: The distillation chamber separation section has 1 to 4 side lines.

11. The hydroprocessing device according to claim 10, characterized in that: The distillation chamber separation section is provided with four side lines: at the first side line or the top of the distillation chamber, the extraction line condensation temperature is 50°C to 70°C, and the reflux ratio is 1.1 to 10.0; the extraction temperature of the second side line or the side line product line is 150°C to 185°C; the extraction temperature of the third side line or the side line return line is 180 to 250°C; and the extraction temperature of the fourth side line or the side line return line is 200 to 370°C.

12. The hydroprocessing device according to claim 1, characterized in that: The number of catalyst beds in the multi-bed hydrodistillation reactor is 2 to 30.

13. A hydroprocessing method for catalytically producing diesel and high-light aromatic naphtha, using the hydroprocessing device according to any one of claims 1 to 12, characterized in that: The steps include: (1) After entering the hydrotreating pretreatment reactor, the low-quality diesel fuel is directly separated into light components and heavy components. The light components are distilled and separated after passing through the refining bed layer upwards, and the naphtha fraction is extracted from the side line; the heavy components are downwardly fed into the multi-bed reaction chamber of the hydrotreating pretreatment reactor, where the upper bed layer is placed with a hydrotreating catalyst, and the lower bed layer is placed with a hydrocracking catalyst; the light components produced in the bed layer of the hydrotreating pretreatment reactor are eventually extracted from the side line upwards, and the unconverted oil in each bed layer is successively fed downwards into the next bed layer for hydrotreating, and the refined diesel fuel enters the hydrocracking reactor; (2) The refined diesel produced in step (1) enters the lower bed of the hydrocracking reactor, and after being dispersed by the liquid distribution component, the lighter part is carried to the middle of the hydrocracking reactor by the hydrogen flowing upward from the bottom, and the heavier part enters the hydrocracking reaction zone of multiple beds downward to contact with the hydrogen countercurrent to carry out the hydrocracking reaction; the light products after hydrocracking in each bed of the lower hydrocracking reaction zone are separated from the reaction bed in time, and the unconverted oil enters the next bed for reaction, and the raw oil of each bed of the hydrocracking reactor is the unconverted oil of the previous bed; the light products separated from each bed are gathered together and enter the middle of the reactor; (3) The light component in step (1) and the lighter part in step (2) first enter the stripping chamber in the middle for stripping, and then enter the distillation chamber in the upper part of the reactor for separation; (4) A side product line is provided in the upper distillation chamber of the hydrocracking reactor, and a naphtha product containing high light aromatics is extracted through the side product line; (5) A side return line is provided in the upper distillation chamber of the hydrocracking reactor in step (2) and step (3), and the heavy monocyclic aromatic hydrocarbon-rich fraction extracted from the side return line re-enters the upper reaction zone of the hydrocracking reactor in step (2) for hydrogenation reaction; the light components produced by the reaction enter the stripping chamber and the distillation chamber of the hydrocracking reactor, and the heavy components enter the lower hydrocracking reaction zone for further reaction, and the bottom of the hydrocracking reactor is the unconverted heavy components.

14. The hydroprocessing method according to claim 13, characterized in that: At least part of the unconverted heavy components in step (5) are circulated into the lower bed of the hydrocracking reaction zone for reaction.

15. The hydroprocessing method according to claim 13, characterized in that: In the hydrogenation pretreatment reactor, the volume ratio of the hydrorefining catalyst to the hydrocracking catalyst is 1:1 to 10:

1.

16. The hydroprocessing method according to claim 13, characterized in that: The number of catalyst beds in the reaction chamber of the hydrogenation pretreatment reactor and the hydrocracking reactor is 2 to 30, and the height of each bed is 50 to 500 mm.

17. The hydroprocessing method according to claim 13, characterized in that: The liquid feed inlet of the hydrogenation pretreatment reactor is arranged at the top of the reaction chamber, and a hydrogen feed inlet is arranged at the bottom of each bed layer in the reaction chamber.

18. The hydroprocessing method according to claim 13, characterized in that: A reaction zone is arranged in the distillation chamber of the hydrogenation pretreatment reactor. The number of catalyst beds in the reaction zone is 1 to 3 layers, and refined catalysts with different functions are arranged.

19. The hydroprocessing method according to claim 13, characterized in that The operating conditions of the hydrogenation pretreatment reactor in step (1) are as follows: the average reaction temperature is 200-450°C, the reaction pressure is 2.5-20 MPa, the hydrogen-to-oil volume ratio is 100:1-3000:1, and the raw material inferior diesel flow rate and the hydrogenation pretreatment reactor catalyst volume ratio is 0.1-10.0h -1 , the reflux ratio is 1.1~10.

0.

20. The hydroprocessing method according to claim 13, characterized in that In step (5), the average reaction temperature of the upper reaction zone of the hydrocracking reactor is 210-430° C., the reaction pressure is 3-20 MPa, and the hydrogen-to-oil volume ratio is 100:1-3500:

1.

21. The hydroprocessing method according to claim 13, characterized in that The average reaction temperature of the lower reaction zone of the hydrocracking reactor is 230-430° C., the reaction pressure is 3-20 MPa, and the volume ratio of hydrogen to oil is 100:1-4000:1.

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