CN115785995A

CN115785995A - Inferior diesel oil hydrotreating method and device

Info

Publication number: CN115785995A
Application number: CN202111059102.7A
Authority: CN
Inventors: 郭兵兵; 赵玉琢; 刘涛; 宣根海; 王晶晶
Original assignee: China Petroleum and Chemical Corp; Sinopec Dalian Research Institute of Petroleum and Petrochemicals
Current assignee: Sinopec Dalian Petrochemical Research Institute Co ltd; China Petroleum and Chemical Corp
Priority date: 2021-09-10
Filing date: 2021-09-10
Publication date: 2023-03-14

Abstract

The invention discloses a method and a device for hydrotreating inferior diesel oil, which adopt a treatment mode of combining hydrotreating and hydrocracking. Wherein the heavier part of the liquid product after hydrotreating enters a hydrocracking section to be in countercurrent contact with hydrogen for hydrocracking reaction, and the lighter part is carried out by a hydrogen-carrying reactor and a light product obtained from the hydrocracking section to enter a separation zone at the upper part of the hydrocracking section together for separation. The design of the hydrogenation process of the invention ensures that the light component can be quickly separated from the system and can not excessively participate in the cracking reaction, and because the product can be quickly separated from the reaction system, the hidden trouble that the byproduct blocks the catalyst is eliminated while the positive reaction speed is increased, and the yield of the target product is improved.

Description

Inferior diesel oil hydrotreating method and device

Technical Field

The invention belongs to the field of hydrogenation reaction, and particularly relates to a method and a device for hydrotreating poor-quality diesel oil.

Background

With the improvement of life quality and the strict global environmental protection requirements of people in the future, the proportion of new energy vehicles is higher and higher, and instead, the demand proportion of gasoline and diesel is lower and lower, and the market competition of the part is more and more harsh. The diesel hydrogenation technology is further improved to improve the yield of the target products except the diesel and improve the market competitiveness.

Hydrocracking technology is one of important heavy oil lightening technologies, but the prior art mainly produces clean fuels, namely, heavy distillate oil is subjected to hydrocracking to produce high-quality clean fuels such as liquefied gas, gasoline, kerosene, jet fuel, diesel oil and the like. Therefore, the development of a new hydrocracking technology for converting the poor-quality diesel raw material into chemical light oil such as naphtha and hydrocracking tail oil as much as possible is one of effective ways for solving the shortage of the chemical light oil.

Straight run naphtha is the primary source of feed to the reformer. For a long time, the yield of crude oil light oil is low in China, the straight-run naphtha is one of the raw materials of an ethylene unit, and the shortage of reforming raw materials becomes one of the main factors limiting the development of the reforming unit. The hydrocracking process is an important means for converting heavy oil into light oil, and the obtained heavy naphtha has the characteristics of high aromatic hydrocarbon content and low sulfur and nitrogen impurity content, and can be directly used as a high-quality reforming device for feeding to make up for the defects of straight-run naphtha.

Depending on the market demand for naphtha distillates, there is a need for hydroprocessing techniques that further increase the yield and selectivity of heavy naphtha.

Chinese researchers obtained methods for improving selectivity and production of chemical materials by various methods, for example, patent CN 101173189A adopts two-stage hydrocracking method with different process parameters to improve the yield of chemical materials. CN103773469A adopts two-stage reactor of hydrofining and hydrocracking, and recycles different fractions and improves octane gasoline with different catalysts. CN103100400A gives full play to the hydrogenation and cracking functions of the catalyst through the optimized matching of the hydrogenation component and the cracking component, and achieves the purpose of increasing the yield of high-quality heavy naphtha. CN104611049A and CN103805256A combine reactive distillation and hydrocracking together, and produce high octane naphtha and high quality diesel oil by using catalytic diesel oil. CN104611062A adopts a rectification reactor and an isomerization catalyst to produce high-octane gasoline.

The height-diameter ratio (the ratio of the total height to the diameter of a reactor bed) of a traditional fixed bed for hydrogenation reaction is generally selected to be 2-10 so as to ensure that a gas-liquid material is fully contacted with a solid catalyst and achieve the required reaction depth and efficiency. Dongfangliang et al, 1998.1 (Yidui technology) and "determination of main structural parameters of hydrogenation reactor" mention that "the ratio of height to diameter of conventional fixed bed reactor is more than 4-9" in order to avoid small ratio of height to diameter "poor catalyst contact efficiency due to uneven fluid distribution". Patent CN109679689A also mentions that the aspect ratio of the existing liquid phase hydrogenation reactor is generally 2.5-12. The design of the aspect ratio of the hydrogenation reactor bed layer becomes the solidification cognition of the technicians in the field, the application of a large amount of industrial practice also proves that the design has rationality and relatively universal adaptability, and the wide industrial success can also lead the technicians to not more comprehensively and more deeply study whether other better choices exist in different types of reactions, and no relevant study report exists for a long time or only a study report which proves that the existing aspect ratio is suitable for the design exists.

Meanwhile, the cracking reaction of the cracked product may occur again in the cracking process, so that the reaction depth is uncontrollable, and under the reaction system, the selectivity of distillate oil such as light naphtha and heavy naphtha is limited, and the reactor is coked and blocked. In addition, the hydrogenation reactor also has the risk of catalyst bed temperature runaway or flooding, particularly for the catalyst with higher activity, the reactor temperature runaway speed is surprising and can usually reach 500-600 ℃, and individually exceeds 800 ℃, and the catalyst can be disabled under severe conditions.

Disclosure of Invention

Aiming at the defects of the prior art, the invention aims to provide a method and a device for hydrotreating poor-quality diesel oil, so as to effectively control the reaction depth, thereby achieving the purpose of improving the yield of target products, particularly improving the selectivity and the yield of naphtha, and simultaneously eliminating the risk of bed layer temperature runaway or flooding.

The first aspect of the invention provides a poor diesel oil hydrotreating method, which comprises the following steps:

(1) Raw materials of poor diesel and hydrogen enter a hydrotreating reaction zone to be subjected to hydrotreating and then are separated into gas and liquid;

(2) After the liquid generated in the step (1) is dispersed by a liquid distribution component, the lighter part of the liquid is carried upwards by hydrogen flowing upwards from the bottom, the heavier part of the liquid enters a hydrocracking reaction zone downwards and is in countercurrent contact with the hydrogen to carry out hydrocracking reaction, a light product generated by the hydrocracking reaction upwards leaves the hydrocracking reaction zone, and a heavy product generated by the hydrocracking reaction downwards leaves the hydrocracking reaction zone;

(3) A separation zone is arranged at the upper part of the hydrocracking reaction zone, the lighter part of liquid carried away by the hydrogen in the step (2) and the light product generated by the hydrocracking reaction upwards enter the separation zone and are separated to obtain a naphtha raw material, and the separated uncracked heavy component enters the hydrocracking reaction zone again for the hydrocracking reaction;

(4) At least part of heavy products generated by the hydrocracking reaction is recycled to the hydrocracking reaction zone to be used as hydrocracking feed.

Further, in the above technical scheme, the hydrocracking reaction zone in step (2) includes at least one catalyst bed, 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 equivalent diameter height ratio) is 2 to 1, and preferably 3. Unless otherwise specified herein, the ratio of equivalent diameter to height refers to the ratio of equivalent diameter to the total height of the catalyst bed in the reactor, and when there are multiple catalyst beds, the total height of the catalyst bed refers to the sum of the multiple catalyst beds.

A great deal of research shows that in the gas-liquid-solid three-phase reaction process with the liquid phase quantity rapidly reduced and the gas phase quantity rapidly increased in the reaction, the gas phase quantity rapidly increases, and occupies a great deal of bed gaps, so that the liquid phase flow velocity is greatly increased. According to the traditional design, although the gas-liquid-solid three-phase contact can be ensured to be sufficient, the effective reaction time of the liquid phase needing further conversion is reduced, the contact probability of the gas phase (such as the gas phase obtained by liquid phase conversion under the reaction condition) which does not need to be reacted again and the catalyst is increased, the overall reaction effect is limited to a certain extent for a system which needs more liquid phase conversion and gas phase control secondary reaction, and the reaction conversion rate, the selectivity and the like are generally expressed to be difficult to further promote.

Research shows that when the total airspeed is close, aiming at the gas-liquid-solid three-phase hydrogenation reaction with the liquid phase quantity rapidly reduced and the gas phase quantity rapidly increased in the reaction process, when hydrogen is contacted with raw oil in a gas-liquid countercurrent mode, the diameter-height ratio of a catalyst bed layer in a reactor is obviously higher than that of the conventional technology, the generated gas phase is enabled to rapidly leave the catalyst bed layer, the accumulation effect of adverse effects of the generated gas phase is small, the liquid phase can have more sufficient probability of reaction on the catalyst, the traditional recognition that the small diameter-height ratio can bring about the adverse effects such as poor contact effect is overcome, the effect of obviously improving the yield of a target product (the target product heavy naphtha in the inferior diesel hydrogenation technology) is obtained, and the problems of easy flooding of a countercurrent reactor, limited hydrogen-oil ratio and the like are solved.

Further, in the above technical scheme, the liquid in step (2) is dispersed into small droplets through the liquid distribution assembly, the lighter part is carried upwards under the action of hydrogen stripping, and the heavier part enters the hydrocracking reaction zone downwards under the action of gravity. The lighter fraction is generally referred to as naphtha fraction and the heavier fraction is generally referred to as diesel fraction.

Further, in the above technical scheme, the hydrotreating reaction zone in step (1) is a one-stage fixed bed hydrotreating reactor, which contains a hydrotreating catalyst bed. Preferably, a hydrorefining catalyst is placed at the upper part of the hydrotreating catalyst bed, a hydrocracking catalyst is placed at the lower part of the hydrotreating catalyst bed, and the volume ratio of the hydrorefining catalyst to the hydrocracking catalyst is 1-10, preferably 3. The raw oil can be desulfurized and denitrified by two different types of catalysts, and after proper cracking, the raw oil is changed into clean distillate oil with medium molecular size to be used as a raw material for subsequent hydrocracking reaction.

Furthermore, in the above technical solution, the hydrorefining catalyst may be any suitable hydrorefining catalyst, and is preferably a hydrorefining catalyst with high desulfurization and denitrification activity and less aromatic hydrocarbon saturation. The hydrogenation active component can be one, two or more selected from metals in VIB group and VIII group, such as W, mo, co and Ni, the total content of the metal oxides is 5-75% by weight of the metal oxides, and the carrier of the hydrogenation catalyst can be selected from alumina, amorphous silicon-aluminum, silica, titanium oxide and the like. Meanwhile, a part of auxiliary agents such as P, ti, zr, si, B and the like can also be added into the catalyst. The hydrofining catalyst 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 the Fuop company, HC-P, HC-T, HC-K UF-210/220 and other hydrogenation catalysts developed by the Uop company, HR-406, HR-416, HR-448 and other hydrogenation catalysts developed by the IFP company, ICR154, ICR174, ICR178, ICR179 and other hydrogenation catalysts developed by the CLG company, TK-525, TK-555, TK-557 and other hydrogenation catalysts developed by the Topsoe company, KF-752, KF-756, KF-757, KF-840, KF-848, KF-907, KF-901 and other hydrogenation catalysts developed by the AKZO company. But is not limited to the above-mentioned catalyst.

Further, in the above technical solution, the hydrocracking catalyst may be any suitable hydrocracking catalyst, and preferably, the catalyst active component is one or more of metals in group vib and/or group viii. The group VIB metal is typically Mo and/or W and the group VIII metal is typically Co and/or Ni. The carrier component of the catalyst comprises one or more of alumina, siliceous alumina and molecular sieve, preferably contains molecular sieve, and the molecular sieve can be Y-type molecular sieve. In a preferred embodiment, the catalyst comprises 10-35 wt% of group VIB metal calculated by oxide, 3-15 wt% of group VIII metal calculated by oxide, 5-40 wt% of molecular sieve, 15-72 wt% of amorphous silica-alumina and 10-67 wt% of alumina; the specific surface area of the catalyst is 100-650 m ² The pore volume is 0.15-0.50 mL/g. The hydrocracking catalyst can be a commercial catalyst, and the main commercial catalyst has smoothness3963, FC-18 and FC-32 catalysts developed by petrochemical research institute.

Further, in the above technical scheme, the raw oil poor diesel oil of the present invention is one or more of naphthenic base straight-run diesel oil, coking diesel oil, catalytic cracking diesel oil, coal tar diesel oil fraction, and the like. The density of diesel oil (20 ℃) is 0.85g/cm ³ As described above.

Further, the operation conditions of the hydrotreating process in step (1) are as follows: the average reaction temperature is 200-450 ℃, the reaction pressure is 3-20 MPa, the reflux ratio at the top of the tower is 1.2-4.5, and the volume ratio of hydrogen to oil is 100. The ratio of the flow of the raw oil to the volume of the catalyst in the reactor is 0.1-10.0 h ^-1 And the circulation ratio of the heavy oil at the bottom of the tower to the raw oil is 1. The preferred operating conditions are: the average reaction temperature is 260-420 ℃, the reaction pressure is 4-18 MPa, the reflux ratio at the top of the tower is 1.5-3.0, and the volume ratio of hydrogen to oil is 200. The circulation ratio of the heavy oil at the bottom of the tower to the raw oil is 1-8.

Further, the operating conditions of the hydrocracking reaction zone in the step (2) and the step (3) are as follows: the average operating temperature is 230-450 ℃, the operating pressure is 3-20 MPa, the volume ratio of hydrogen to oil is 200 ^-1 ～10.0h ^-1 The ratio of inlet circulating heavy oil to raw oil is 1. Preferred operating conditions are: the average operating temperature is 280-420 ℃, the operating pressure is 4-18 MPa, the volume ratio of hydrogen to oil is 300-1800.

Further, in the above technical scheme, the hydrocracking reaction zone in the step (2) is divided into a plurality of sub-reaction zones along the longitudinal direction, so that when hydrogen rises to the bottom of the bed layer through the distributor, the diffusion area of the hydrogen can cover the cross-sectional area of the bed layer at the bottom of the sub-reaction zone, and each sub-reaction zone is provided with an independent hydrogen inlet.

Further, in the above technical solution, the hydrocracking reaction zone in the step (2) is filled with a hydrocracking catalyst, and the shape of the hydrocracking catalyst may be any conventional existing shape of hydrocracking catalyst, and is preferably a porous catalyst, a shaped catalyst and/or a honeycomb catalyst. The aperture of the porous catalyst is 1-50 mm, preferably 4-20 mm; the average particle diameter of the special-shaped catalyst is 2-50 mm, preferably 4-30 mm; the diameter or side length of the honeycomb catalyst pores is 1 to 50mm, preferably 3 to 15mm; the voidage of the catalyst bed is recommended to be 15-85%, and the preferred range is 20-75%.

Further, in the technical scheme, the porous catalyst contains 10% -35% of VIB group metal calculated by oxide and 3% -15% of VIII group metal calculated by oxide.

Further, in the above technical solution, the hydrocracking catalyst carrier in the hydrocracking reaction zone is a ceramic honeycomb, pall ring, raschig ring, rectangular saddle ring, saddle, open-pore ring type, semi-ring, stepped ring, double arc, hel ring, conjugated ring, flat ring, flower ring, hollow ball, or other porous carrier.

Further, in the above technical scheme, the preparation method of the hydrocracking catalyst in the hydrocracking reaction zone in the step (2) comprises: the Y-type molecular sieve and a bonding agent prepared from small-pore alumina, large-pore alumina, dilute nitric acid solution and other additives are kneaded and extruded to prepare shapes of a honeycomb body, a pall ring, a Raschig ring, a hollow sphere and the like, and the porous catalyst carrier is prepared after drying and roasting. Then the VIB group and/or VIII group metal is used as active component solution to be dipped on the porous carrier, and after drying and roasting, the catalyst used in the invention is prepared. Or mixing a silicon source, an aluminum source, an active component metal salt solution and an auxiliary agent metal salt solution, aging and drying at proper pH and temperature, kneading and extruding together to prepare porous shapes such as a honeycomb body, a pall ring, a Raschig ring, a hollow sphere and the like, and drying and roasting to prepare the porous catalyst.

Further, in the above technical scheme, the raw poor diesel oil and hydrogen in step (1) are heated by a heating furnace before entering the hydrotreating reaction zone, preferably to 200-350 ℃. The separation in the step (1) is preferably carried out in a gas-liquid separation tank, and the separated gas containing hydrogen and non-condensable gas is pressurized from the top of the separation tank through a hydrogen compressor and enters a hydrotreating reaction zone as circulating hydrogen.

The second aspect of the present invention provides an inferior diesel hydrotreating apparatus, comprising:

(1) The hydrotreating reactor is used for hydrotreating the raw material poor-quality diesel oil;

(2) Heterotypic hydrogenation ware, heterotypic hydrogenation ware include by lower supreme heavy oil storehouse, reaction chamber and the separator that communicates in proper order, and the reaction chamber is provided with from top to bottom in proper order: liquid distribution subassembly, catalyst bed and hydrogen distribution chamber.

Further, in the above technical scheme, the ratio of the equivalent diameter of the cross-sectional area of the catalyst bed in the special-shaped hydrogenation reactor to the total height of the catalyst bed in the reactor is 2-10, preferably 3. The reactor bed according to the invention is preferably isodiametric, i.e. the cross-sectional area is the same at different locations throughout the catalyst bed. The cross-sectional area of the catalyst bed is generally the same as the cross-sectional area of the reaction chamber in the reactor, with cross-section being taken as the top cross-section, i.e. the section perpendicular to the vertical within the reaction chamber. If there is a difference in the cross-sectional area of the reactor over the height of the catalyst bed, the cross-sectional area here means the average of the cross-sectional area of the catalyst bed or the cross-sectional area of the reaction chamber throughout the catalyst bed.

Furthermore, in the above technical scheme, the reaction chamber is a horizontal storage tank, the axial direction of the reaction chamber is transversely arranged, and end enclosures are arranged at two ends of the horizontal storage tank. Or the reaction cavity is an oblate cylindrical tank, and the axial direction of the tank is arranged along the longitudinal direction.

Further, in the above technical scheme, the cross section of the catalyst bed in the reaction chamber is rectangular or circular, and the height of the bed is 100-5000 mm, preferably 200-1000 mm.

The catalyst bed layer in the invention can also greatly increase the material flux passing through the bed layer under a higher diameter-height ratio, and simultaneously reduce the accumulation of heat in the catalyst bed layer, and simultaneously, the generated hydrogen sulfide and ammonia gas are rapidly taken out of the reactor, thereby avoiding the blockage phenomenon of the traditional reactor caused by overlong time of the by-product staying in the bed layer.

Furthermore, in the above technical scheme, a plurality of clapboards are arranged in parallel along the vertical direction in the hydrogen distribution chamber and the catalyst bed layer in the reaction chamber, the plurality of clapboards divide the hydrogen distribution chamber into a plurality of air inlet units, and the bottom of each air inlet unit is provided with at least one hydrogen inlet. A plurality of holes are distributed on each clapboard; the partition plate extends upwards to the catalyst layer, the opening rate of the partition plate below the catalyst bed layer is less than 70%, and the opening rate of the partition plate in the catalyst layer is more than 50%.

Furthermore, in the above technical scheme, the bottom of the catalyst bed between every two adjacent partition plates corresponds to 1-3 hydrogen distributors, and the distribution area of hydrogen from all the hydrogen distributors in the partition plate area when reaching the bottom of the regional bed should cover the bottom of the whole regional bed. Further, the partition plate is annular or segmental.

Further, in the above technical scheme, a gas distributor is arranged at the hydrogen inlet. In the invention, the gas distributor is preferably a tangential circulation type distributor or a rotating blade distributor, and the gas distributor can ensure that the flow velocity of gas entering the whole catalyst bed layer interface is uniform, thereby avoiding the situations of bias flow, channeling and the like.

Further, in the above technical scheme, the heavy oil bin is arranged in the center of the bottom of the reaction chamber, and the heavy oil bin is communicated with the plurality of air inlet units.

Further, in the above technical scheme, the ratio of the diameter (or the maximum equivalent diameter) of the separator at the upper end of the special-shaped hydrogenation reactor to the diameter (or the maximum equivalent diameter) of the reaction chamber at the lower part is 1.2-1, preferably 1. In the invention, the diameter of the separator at the upper part of the special-shaped hydrogenation reactor is reduced, so that the load of light fraction under high pressure is completely matched with the tower plate, the separation efficiency of the tower plate is high, and the special-shaped hydrogenation reactor has the complete substitution of a fractionating tower.

Further, in the above technical scheme, the separator includes a mixing section, a separation section and a stabilizing section from bottom to top. The height of the mixing section is 25-40% of the total height of the separator, the height of the separating section is 50-65% of the total height of the separator, and the height of the stabilizing section is 5-10% of the total height of the separator.

Further, in the above technical scheme, a tray or a filler is placed in the separation section, and whether the filler is placed or not is not limited by the mixing zone and the stabilizing zone. The above-mentioned fillers or column plates are all conventional in the art, for example, the fillers can be selected from one or several kinds of random fillers such as pall rings, raschig rings, rectangular saddle rings, saddle shapes, open-pore ring types, semi-rings, stepped rings, double arcs, halter rings, conjugated rings, flat rings, flower rings, etc., and the fillers can also be selected from metal or ceramic corrugated fillers. The tower plate is one or more of bubble cap plate, sieve plate, float valve plate, mesh plate, tongue plate, guide sieve plate, multi-downcomer tower plate, etc. and may be also through-flow sieve plate, through-flow corrugated plate, etc. without downcomer. High-efficiency trays such as a float valve plate, a sieve plate and the like are preferred.

Further, in the above technical scheme, the separation section is provided with 1-3 product side lines; the mixing section is provided with 1-3 light raw material lateral lines, and the mixing section is provided with 1-3 reaction zones.

Further, in the above technical solution, the liquid distribution assembly in the reaction chamber includes a liquid distributor, and a liquid distribution tray and a distribution cone disposed above the liquid distributor. For dispersing the liquid produced in the hydroprocessing reactor into small droplets, the lighter part is carried upwards under the action of hydrogen stripping and the heavier part enters the hydrocracking reaction zone. The lighter fraction is typically a naphtha fraction and the heavier fraction is typically a diesel fraction.

Further, in the above technical solution, the liquid distributor is a conventional distributor in the art, such as a shower head distributor, a coil pipe distributor, a porous straight pipe distributor, a straight pipe baffle distributor, a baffle plate distributor, a tangential circulation distributor, a rotating vane distributor, a double-row vane distributor, and the like. The liquid distributor is preferably a porous tubular distributor or a straight tube baffle distributor, and the diameter of a pore passage of the tubular distributor is 0.5-20 mm, preferably 2-10 mm. The farther from the feedstock oil inlet end, the larger the pore diameter. The height of the liquid distributor from the top of the reactor bed is 1-1000 mm, preferably 50-500 mm. The height is related to the nature, temperature and pressure of the raw oil. Generally, the higher the temperature, the higher the height of the liquid distributor from the bed, so that the distributor can make the raw material fall on the bed surface more uniformly in a higher space. Also, the higher the pressure, the larger the spray angle of the liquid distributor, the lower the height from the top of the reactor bed can be, and the more space-saving.

Further, in the technical scheme, the shape of the liquid distribution disc of the special-shaped hydrogenation reactor is the same as the cross section of the catalyst bed layer of the main hydrogenation reactor, and the area of the liquid distribution disc is 10-100%, preferably 60-100% of the cross section of the catalyst bed layer.

Furthermore, in the above technical scheme, a plurality of first through holes are uniformly formed in the distribution disc, a first overflow ring is arranged around the first through holes, and an overflow part is arranged at the outer edge of the distribution disc; the aperture ratio of the distribution disc is 5-90%, the diameter of the first through hole is 5-100 mm, and the height of the first overflow ring is 1-30 mm.

Furthermore, in the above technical scheme, the inner side of the first overflow ring is provided with a sawtooth part, the sawtooth part bends downwards, and the sawtooth part is provided with a diversion trench.

Further, in the above technical solution, the distribution cone is disposed in the center of the upper portion of the liquid distribution tray, the distribution cone is provided with a plurality of second through holes, and a second overflow ring is disposed around the second through holes. The vertex angle of the distribution cone is more than 90 degrees, the aperture ratio of the distribution cone is 5-80 percent, and the height of the second overflow ring is 1-30 mm; the bottom area of the distribution cone is 2-15% of the area of the liquid distribution disc.

Further, in the above technical scheme, the catalyst bed layer in the step (2) is a porous catalyst layer, a shaped catalyst layer or a honeycomb layer. The porosity of the catalyst layer is 15-85%, preferably 20-75%; the diameter of the porous catalyst is 1-50 mm, preferably 4-20 mm; the average particle diameter of the special-shaped catalyst is 2-50 mm, preferably 4-30 mm; the honeycomb catalyst has a cell diameter or a cell edge length of 1 to 50mm, preferably 3 to 15mm.

Further, in the above technical scheme, the heterotypic hydrogenation ware still includes: and one end of the reboiler is connected with the outlet of the heavy oil bin, and the other end of the reboiler is connected with the hydrogen distribution cavity. The temperature of the heavy oil storage bin is kept at the temperature required by the reaction bed layer through a reboiler.

Further, in the above technical scheme, the special-shaped hydrogenation reactor further includes: the multistage auxiliary reaction chamber, each level auxiliary reaction chamber advance hydrogen alone, the bottom center sets up heavy oil storehouse alone, and the liquid raw materials import of each level auxiliary reaction chamber is connected with the heavy oil storehouse of last one-level, and the top of multistage auxiliary reaction chamber all is connected to the separator.

The hydrogen feeding pipe at the bottom of the reaction section at the lower part of the reaction cavity of the special-shaped hydrogenation reactor is provided with a plurality of inlets, and each hydrogen feeding pipe corresponds to the catalyst bed layer area between the two partition plates, so that hydrogen can upwards pass through the reaction area at the top after coming out of each distributor. The connection between the baffle with holes and the bottom of the main reactor has at least one pore channel.

Further, in the above technical solution, it is preferable to open 1 to 3 side lines at the separation section of the separator for withdrawing the desired product. The extraction temperature is 60-80 ℃ at the 1 st side line or the top of the separator, and the extracted components are gas and light naphtha fraction. The 2 nd side draw temperature is from 150 to 190 c, preferably from 160 to 180 c, the components withdrawn in this side draw being a heavy naphtha fraction. The 3 rd side draw temperature is from 180 to 250 c, preferably from 180 to 230 c, the components being the jet fuel fraction.

Further, in the above technical solution, the side draw-off line may be provided with a reflux.

The invention carries out deep research on the hydrogenation reaction with rapidly reduced liquid phase amount and rapidly increased gas phase amount under the reaction condition, breaks through the long-term design scheme in the field, and provides the technical means with completely different design parameters from the prior art. The method is characterized in that the method carries out comprehensive and deep research on multidimensional influence factors such as gas-phase reaction time, liquid-phase reaction time, gas-phase product secondary reaction control, reaction heat discharge control, flooding control, gas-liquid-solid contact effect and the like, and provides a proper technical scheme such as a process flow, a catalyst bed layer structure and the like on the basis. The method has the advantages that the method obtains outstanding technical effects on the aspect of maximum production of high-added-value heavy naphtha products by hydrogenation of poor diesel, the yield of the heavy naphtha can be improved by 20 percent under the same conditions, the expectation of technicians in the field is exceeded, and the conventional solidification cognition in the field is broken through.

The liquid distribution component is arranged, so that the liquid after hydrotreating can be dispersed into proper small droplets, and the lighter part can be directly carried out by virtue of the stripping action of hydrogen without entering a hydrocracking reaction zone to participate in hydrocracking reaction. Meanwhile, through the design of the liquid distribution component, heavy components entering the hydrocracking reaction zone can be distributed more uniformly, and the problem that reactants on a catalyst bed layer of a traditional reactor are poor in contact under the condition of large diameter-height ratio is solved. By adopting the reactor, the void ratio of the reactor can be smaller under the same process condition and product index requirements. Meanwhile, when the special-shaped catalyst is adopted, the temperature rise of a reaction bed layer is controlled to be smaller, and the allowable feeding load is larger. The improvement of the pore structure can improve the flooding characteristic (no flooding for a long time) and simultaneously can ensure good mass transfer performance. Finally, the product properties have good controllability.

In addition, the invention can make the light component of distillate oil after the second hydrogenation quickly separate from the system and not excessively participate in the cracking reaction by reasonably setting the reaction process and controlling the catalyst bed structure, and because the product can quickly leave the reaction system, the hidden trouble that the byproduct blocks the catalyst is eliminated while the positive reaction speed is increased, the yield of the target product is improved, and the service life of the catalyst is prolonged.

The invention can realize timely extraction of light intermediate products through flash evaporation and steam stripping by arranging the separator, thereby effectively controlling the reaction degree, furthest retaining aromatic hydrocarbon components and becoming good chemical raw materials. Meanwhile, the partial pressure of the product is kept in a low state all the time, so that the reaction speed is accelerated, the reaction efficiency is favorably improved, and harmful components such as hydrogen sulfide, ammonia and the like which are easy to coke can be taken away.

The preferable catalyst is loaded on the porous material, so that the void ratio is increased, the flux of the catalyst bed is increased, and the flooding cannot be formed.

Compared with the prior hydrocracking technology, the selectivity of the heavy naphtha fraction can be improved by 5-20 percent, and the yield of the light naphtha and the heavy naphtha can reach more than 95 percent. The naphtha fraction obtained by the invention is rich in aromatic hydrocarbon and naphthenic hydrocarbon, and is a high-quality raw material for a reforming device.

Drawings

FIG. 1 is a schematic diagram of a process flow of hydrogenation of inferior diesel and a schematic diagram of a special-shaped hydrogenation reactor;

wherein: 1-raw oil inlet; 2-new hydrogen; 3-hot raw oil; 4-hot hydrogen; 5-hydrogen oil mixed raw material; 6-a hydrotreating reactor; 7-hydrotreating reactor bottom pipe; 8-a gas-liquid separation tank; 9-special-shaped hydrogenation reactor inlet; 10-a special-shaped hydrogenation reactor reaction cavity; 11-special-shaped hydrogenation reactor catalyst bed layer; 12-hydrogen; 13-heavy oil sump; 14-a liquid distribution assembly; 15-catalyst bed layer partition; 16-a mixing section; 17-a separation section; 18-a stabilizing section; 19-a separator; 20-special-shaped hydrogenation reactor top pipeline; 21-a condenser; 22-a liquid separation tank; 23-hydrogen rich and non-condensable gases; 24-a light naphtha fraction; 25-heavy naphtha fraction; 26-heavy diesel fraction; 27-a circulating oil pump; 28-diesel feedstock; 29-a desulfurization deamination device; 30-a cyclic hydrogen press; 31-heating furnace.

Fig. 2 is a side view of a segmental baffle according to an embodiment of the invention.

Fig. 3 is a schematic diagram of a segmental baffle in side view in a configuration according to another embodiment of the invention, which is only located in the hydrogen distribution chamber and which has no segmental baffle inside the catalyst bed.

Fig. 4 is a schematic top view of an annular partition according to an embodiment of the present invention.

Fig. 5 is a schematic side view of an annular partition according to two embodiments of the present invention, wherein the partition of fig. 5-1 is located in both the hydrogen distribution chamber and the catalyst bed, and the partition of fig. 5-2 is located only in the hydrogen distribution chamber, and the annular partition is not located inside the catalyst bed.

Fig. 6 is a schematic top view of a liquid distribution assembly according to an embodiment of the present invention.

Fig. 7 is a schematic top and side view of a liquid distribution assembly according to another embodiment of the invention.

FIG. 8 is a schematic top view of a first overflow ring according to an embodiment of the invention.

Fig. 9 is a perspective view of the first overflow ring of fig. 8.

Detailed Description

Specific embodiments of the present invention will be described in detail below with reference to the accompanying drawings, but it should be understood that the scope of the present invention is not limited to the specific embodiments.

Throughout the specification and claims, unless explicitly stated otherwise, the word "comprise", or variations such as "comprises" or "comprising", will be understood to imply the inclusion of a stated element or component but not the exclusion of any other element or component.

Spatially relative terms, such as "below," "lower," "upper," "above," "upper," and the like, may be used herein for ease of description to describe one element or feature's relationship to another element or feature in the figures. It will be understood that the spatially relative terms are intended to encompass different orientations of the object in use or operation in addition to the orientation depicted in the figures. For example, if the items in the figures are turned over, elements described as "below" or "beneath" other elements or features would then be oriented "above" the elements or features. Thus, the exemplary term "below" can encompass both an orientation of below and above. The articles may be otherwise oriented (rotated 90 degrees or at other orientations) and the spatially relative terms used herein should be interpreted accordingly.

In this document, the terms "first", "second", etc. are used to distinguish two different elements or portions, and are not used to define a particular position or relative relationship. In other words, the terms "first," "second," etc. may also be interchanged with one another in some embodiments.

FIG. 1 shows a schematic flow chart of the hydrogenation process of poor diesel oil of the present invention. The inferior diesel oil is sent into a heating furnace 31 from a raw oil inlet 1, heated to 200-350 ℃, then converged with heated hot hydrogen 4 to form a hydrogen-oil mixed raw material 5, and enters a hydrotreating reactor 6, wherein a hydrotreating catalyst and a hydrocracking catalyst are sequentially placed in the hydrotreating reactor from top to bottom, distillate oil after bottom hydrogenation enters a gas-liquid separation tank 8 from a hydrotreating reactor bottom pipeline 7 for gas-liquid separation, and non-condensable gas enters a condenser 21. The liquid diesel oil fraction enters a reaction cavity 10 of the special-shaped hydrogenation reactor through an inlet 9 of the special-shaped hydrogenation reactor, and after passing through a liquid distribution assembly, small molecular component hydrocarbons dissolved in the diesel oil are further released into a separator 19 as light components. The rest heavy component diesel oil uniformly enters different sub catalyst bed regions in the catalyst bed layer 11 of the special-shaped hydrogenation reactor, and hydrogen 12 enters from the bottom and uniformly enters the catalyst bed layer 11 of the special-shaped hydrogenation reactor under the action of a gas distributor and a catalyst bed layer partition 15. Heavy components in the raw material diesel oil are subjected to olefin saturation and cracking under the action of hydrogen and a catalyst to generate small molecular hydrocarbons, and the small molecular hydrocarbons and part of macromolecular hydrocarbons are separated from the bed layer and enter a separator 19 under the carrying of the hydrogen. Under the action of a separation tower plate of the separation section 17, macromolecular hydrocarbons carried into the separator are separated, the macromolecular hydrocarbons downwards enter the catalyst bed layer 11 of the special-shaped hydrogenation reactor again for cracking, the micromolecular hydrocarbons upwards are extracted as a heavy naphtha fraction 25 at the lateral line and sent out as a product, non-condensable gas is continuously cooled by a condenser 21, liquid hydrocarbons are refluxed or extracted as a light naphtha fraction 24 after gas-liquid separation in a liquid separating tank 22, and the gas enters desulfurization and deamination equipment 29 for purification and then is recycled. The heavy diesel fraction 26 at the bottom of the reaction chamber is returned to the inlet 9 of the profiled hydrogenation reactor by a recycle oil pump 27 or withdrawn as a diesel feedstock 28.

Further, in one or more exemplary embodiments of the present invention, the reaction chamber 10 may be a horizontal tank, as shown in fig. 1, which is disposed axially and transversely, and has a sealing head at both ends. Further, in one or more exemplary embodiments of the present invention, the reaction chamber 10 may also be a flat cylindrical tank, which is axially disposed in the longitudinal direction.

Further, in one or more exemplary embodiments of the present invention, the shape of the partition is matched with the bottom of the reaction chamber 10, and when the reaction chamber 10 is a horizontal tank, the partition is a segmental partition, as shown in fig. 2 and 3; when the reaction chamber 10 is a flat cylindrical tank, the plurality of partition plates are coaxial annular partition plates, as shown in fig. 4 to 6. Further, in one or more exemplary embodiments of the present invention, a plurality of circular holes are distributed on each of the separators. Further, in one or more exemplary embodiments of the present invention, as shown in fig. 1, a plurality of partition plates 15 may extend upward to the catalyst bed 11, and the partition plates not in contact with the catalyst bed 11 at the bottom have an opening rate of less than 70%, and a lower opening rate is beneficial to increase the resistance, so that the hydrogen gas enters the catalyst bed 11 as upward as possible, further functioning as a gas distributor. The aperture ratio of the partition plates in the catalyst bed layer 11 is more than 50 percent, which is beneficial to more fully utilizing the catalyst.

Further, in one or more exemplary embodiments of the present invention, the ratio of the cross-sectional area of the separator 19 to the reaction chamber 10 is 1.

Further, in one or more exemplary embodiments of the present invention, the separator 19 includes, from bottom to top, a mixing section 16, a separation section 17, and a stabilization section 18.

Further, in one or more exemplary embodiments of the present invention, the liquid distribution assembly 14 between the separator 19 and the reaction chamber 10 includes a liquid distribution tray and a distribution cone. The liquid distribution plate is arranged above the liquid distributor, the shape of the liquid distribution plate is the same as that of the top surface of the porous catalyst layer, a plurality of first through holes are uniformly formed in the liquid distribution plate, a first overflow ring is arranged around the first through holes, and an overflow part (not shown in the figure) is arranged at the outer edge of the distribution plate. The distribution cone is arranged at the center of the upper part of the liquid distribution disc, the distribution cone is provided with a plurality of second through holes, and a second overflow ring (not shown in the figure) is arranged around the second through holes.

Further, in one or more exemplary embodiments of the invention, the liquid distribution tray area is 10% to 100%, preferably 60% to 100% of the cross-section of the catalyst bed. The opening ratio of the liquid distribution disc is 5-90%, the diameter of the first through hole is 5-100 mm, the height of the first overflow ring is 1-30 mm, and the height of the overflow part at the edge of the liquid distribution disc can be the same as the height of the first overflow ring, which is not limited by the invention. The vertex angle of the distribution cone is larger than 90 degrees, the vertex angle of the distribution cone is more gentle and better, and the flow speed of the liquid drops in the liquid distribution disc can be more gentle. The aperture ratio of the distribution cone is 5-80%, and the height of the second overflow ring is 1-30 mm; the bottom area of the distribution cone is 2-15% of the area of the liquid distribution disc. The distribution cone prevents the liquid returned by the separator from flowing down all the way through the first through hole in the center of the liquid distribution tray, so that the liquid is distributed more evenly.

Preferably, but not limitatively, in one or more exemplary embodiments of the present invention, as shown in fig. 8 and 9 in combination, the inner side of the first overflow ring is provided with serrations 14-1 bent downward and provided with guide grooves 14-2. Illustratively, the flow guide grooves are opened along the center of the sawtooth part.

In one or more embodiments of the present invention, the profiled hydrogenation reactor further comprises an auxiliary reaction chamber. It should be appreciated that the auxiliary reaction chamber may be multi-staged. Each level of auxiliary reaction cavity independently enters hydrogen, the bottom center independently sets up heavy oil bin, the liquid raw material inlet of each level of auxiliary reaction cavity is connected with the heavy oil bin of last level, and the top of multistage auxiliary reaction cavity all is connected to separator 19.

Further, in one or more exemplary embodiments of the invention, the catalyst bed has a void volume of 15% to 85%, preferably 20% to 75%.

Further, in one or more exemplary embodiments of the present invention, the catalyst bed layer is a porous catalyst layer, a shaped catalyst layer, or a honeycomb layer. The diameter of the porous catalyst is 1-50 mm, preferably 4-20 mm; the average particle diameter of the special-shaped catalyst is 2-50 mm, preferably 4-30 mm; the honeycomb catalyst has a cell diameter or a cell edge length of 1 to 50mm, preferably 3 to 15mm.

The following examples further illustrate a hydrogenation method of poor diesel oil capable of improving naphtha selectivity and yield and the effects of the invention. In the examples, detailed embodiments and specific operation procedures are given on the premise of the technical scheme of the present invention, but the scope of the present invention is not limited to the following specific examples.

The experimental procedures in the following examples are, unless otherwise specified, conventional in the art.

Example 1

The raw oil is catalytic diesel oil, and the properties are shown in Table 2. Heating raw oil and hydrogen to 300 ℃ by a heating furnace, and then feeding the raw oil and the hydrogen into a hydrotreating reactor, wherein the upper part of a reaction zone in the hydrotreating reactor is a catalyst A, the lower part of the reaction zone is filled with a catalyst C, and the volume ratio of the two catalysts is V _A ：V _C 1, and specific catalyst properties are shown in table 1. The raw oil is hydrotreated in a hydrotreating reactor, the obtained product is separated by a gas-liquid separation tank, and gas flows out from the top of the separation tank and is refluxed to the hydrotreating reactor. The liquid is used as the raw material of the special-shaped hydrogenation reactor. And (3) filling a catalyst D in the special-shaped hydrogenation reactor, wherein the equivalent diameter-height ratio of a catalyst bed layer is 4:1. a catalyst bed layer is arranged in the reaction cavity of the special-shaped hydrogenation reactor, the cross section of the catalyst bed layer is circular, and the height of the bed layer is 800mm. An annular partition plate is arranged in the catalyst bed layer, and the number of the partition plates is 4. A plurality of holes are distributed on the partition plate; the partition plates extend upward to the catalyst layer, and the partition plates below the catalyst layer have an aperture ratio of 40% and the partition plates in the catalyst layer have an aperture ratio of 70%. The liquid used as the raw material of the special-shaped hydrogenation reactor is firstly subjected to light and heavy component separation through a liquid distribution assembly. And the light component upwards enters a separator at the upper part of the special-shaped hydrogenation reactor to carry out fine separation of different fractions according to different temperature distributions. The heavy fraction after the fine separation downwards passes through the liquid distribution component, and then is subjected to a cracking reaction together with heavy components after the heavy fraction and the light component are separated from each other through the liquid distribution component at the raw material inlet downwards and with hydrogen ascending from the bottom under the action of a catalyst, the generated light fraction quickly and upwards breaks away from a reaction system and enters a mixing section of a separator, the separated light fraction upwards enters a separation section of the separator, and is subjected to desulfurization and denitrification through a hydrofining bed layer in the separation section, and olefins are saturated. The heavy fraction which is not cracked sufficiently flows out from the bottom of the heavy oil bin, passes through a circulating pump and then enters the inlet of the special-shaped hydrogenation reactor as circulating oil, and the specific operating process conditions are shown in the specificationTable 3. The material extracted from the upper side line of the special-shaped reactor is used as a heavy naphtha raw material, and can be fed into a stripping tower for stripping, and then a naphtha fraction product is produced from the bottom of the stripping tower. The reaction conditions and the product distribution are shown in Table 4, and the simulated bed temperature rise changes are shown in Table 5.

The liquid distribution assembly comprises a liquid distributor, a liquid distribution disc and a distribution cone, wherein the liquid distribution disc and the distribution cone are arranged above the liquid distributor. The liquid distribution plate has the same shape as the top surface of the catalyst bed, and the area of the liquid distribution plate is 70% of the cross section of the catalyst bed. A plurality of first through holes are uniformly formed in the liquid distribution disc, a first overflow ring is arranged around the first through holes, an overflow part is arranged at the outer edge of the distribution disc, the aperture ratio of the liquid distribution disc is 50%, the diameter of each first through hole is 10mm, and the height of each first overflow ring is 10mm. The distribution cone is arranged in the center of the upper part of the liquid distribution disc, the distribution cone is provided with a plurality of second through holes, and second overflow rings are arranged around the second through holes; the vertex angle of the distribution cone is 120 degrees, the opening rate of the distribution cone is 50 percent, and the height of the second overflow ring is 10mm; the base area of the distribution cone is 10% of the area of the liquid distribution tray.

Example 2

The difference between the embodiment and the embodiment 1 is that the raw oil is coker diesel, and the equivalent diameter-height ratio of the catalyst bed in the heterotype hydrogenation reactor is 5:1, 6 baffles in the catalyst bed. A plurality of holes are distributed on the partition plate; the partition plate extends upward to the catalyst layer, and the partition plate below the catalyst layer has an aperture ratio of 30% and the partition plate in the catalyst layer has an aperture ratio of 80%. The area of the liquid distribution tray was 90% of the cross-section of the catalyst bed. The opening rate of the liquid distribution disc is 80%, the diameter of the first through hole is 20mm, and the height of the first overflow ring is 20mm. The distribution cone is arranged in the center of the upper part of the liquid distribution disc, the distribution cone is provided with a plurality of second through holes, and second overflow rings are arranged around the second through holes; the vertex angle of the distribution cone is 150 degrees, the opening rate of the distribution cone is 70 percent, and the height of the second overflow ring is 20mm; the base area of the distribution cone is 15% of the area of the liquid distribution tray. The remaining conditions were the same as in example 1.

Example 3

The difference between the embodiment and the embodiment 1 is that the adopted raw oil is naphthenic base straight-run diesel oil, the catalyst C is selected as the cracking catalyst of the heterotype hydrogenation reactor, and the equivalent diameter-height ratio of the catalyst bed is 5:1. the remaining conditions were the same as in example 1.

Example 4

This example differs from example 1 in that the hydrorefining catalyst A in the upper part of the hydrotreatment reactor was replaced by catalyst B. The remaining conditions were the same as in example 1.

Example 5

The difference between the embodiment and the embodiment 1 is that the catalyst C is selected as the heterotype hydrogenation reactor cracking catalyst, and the equivalent diameter-height ratio of the catalyst bed is 4:1. the remaining conditions were the same as in example 1.

Example 6

This example differs from example 1 in the different proportions of refining catalyst and cracking catalyst in the hydroprocessing reactor, V _A ：V _C 1, = 5. The remaining conditions were the same as in example 1.

Example 7

The difference from the example 1 is that the diameter-height ratio of the catalyst bed in the heterotype hydrogenation reactor is 1:3.

example 8

The difference from the example 1 is that the catalyst filled in the special-shaped hydrogenation reactor is prepared by adopting a conventional carrier, and the prepared catalyst E has the diameter of 1.8mm, the length of 5-8 mm and the bed layer void ratio of 30%.

Example 9

The difference is that no partition plate is arranged at the bottom of the reaction chamber of the special-shaped hydrogenation reactor as in example 1.

The properties of the catalysts of the above examples and comparative examples are shown in Table 1, the properties of the feedstock are shown in Table 2, the process conditions are shown in Table 3, and the product distributions of the examples and comparative examples are shown in Table 4. The simulated temperature rises of the hydrocracking beds of the examples and comparative examples are shown in Table 5.

Comparative example 1

Adopts a conventional two-stage hydrogenation method, namely a refining and cracking method process. The refining reactor and the cracking reactor adopt the reaction process that the raw material and the hydrogen flow from top to bottom in cocurrent flow. The cracked distillate oil is separated by a separation tower and then is sent out as a product. In the comparative example, a refining reactor is filled with a catalyst A and a catalyst C, the volume ratio of the catalyst A to the catalyst C is 3. The remaining process conditions were the same as in example 1.

TABLE 1 catalyst Properties

TABLE 2 Properties of the feed oil

Raw oil	1#	2#	3#
				Name(s)	Catalytic diesel fuel	Coking diesel oil	Naphthenic straight-run diesel oil
Density, g/cm ³	0.930	0.861	0.866
				Range of distillation range, deg.C	191.4～338.6	182～376	200～364
Sulfur content, vol%	1.0	1.2	0.2
				Nitrogen content,. Mu.g/g	1500	2600	340

TABLE 3 Process conditions of examples and comparative examples

One-stage hydrogenation reactor
		Reaction temperature of	320
Reaction pressure, MPaG	8
		Volume hydrogen to oil ratio	300：1
Volumetric space velocity h ^-1	1.2
		Two-section special-shaped hydrogenation reactor
Reaction temperature, deg.C	365
		Reaction pressure, MPaG	8
Volume hydrogen to oil ratio	1200:1
		Volumetric space velocity h ^-1	1.5

TABLE 4 product distribution of examples and comparative examples

From the results in table 4, the yield and selectivity data for heavy naphtha using the process of the present invention are significantly higher than the data associated with comparative example 1.

The bed reaction temperature profiles of the examples and comparative examples were calculated by simulation in the laboratory using ansys version 19.0 software. The simulation conditions were input according to the actual data of examples and comparative examples. The simulation result shows that the traditional fixed bed has the highest central temperature, the temperature change is normally distributed from the inlet end to the outlet end, and the bed temperature of the reactor is relatively uniform. The bed simulated temperature rise changes are shown in Table 5.

TABLE 5 bed simulated temperature rise Change

Temperature point of bed layer	Example 1	Example 2	Comparative example 1
				Maximum radial temperature difference, ° c	1.1	0.5	18.4
Average temperature of	371.7	371.4	382.3

As can be seen from the results in Table 5, the temperature difference of the catalyst beds in examples 1-2 using the present invention is significantly lower than that in comparative example 1, and the temperature difference from the conventional fixed bed of 18.4 ℃ is reduced to 0.5 ℃, and the difference between the average temperature and the control temperature in the examples is minimal, which indicates that the reactor of the present invention has eliminated the overheating phenomenon of the hydrocracking reaction.

Claims

1. The hydrotreating method of inferior diesel oil is characterized by comprising the following steps:

2. The method for hydrotreating poor quality diesel oil according to claim 1, characterized in that in step (2), the hydrocracking reaction zone comprises at least one catalyst bed, and the ratio of the cross-sectional area equivalent diameter of the catalyst bed to the total height of the catalyst bed is (2) - (10), preferably (3).

3. The method for hydrotreating poor quality diesel oil according to claim 1, wherein in step (2), the hydrocracking reaction zone is divided into a plurality of sub-reaction zones along the longitudinal direction, so that when the hydrogen gas rises to the bottom of the bed through the distributor, the diffusion area of the hydrogen gas can cover the cross-sectional area of the bed at the bottom of the sub-reaction zones, and each sub-reaction zone is provided with a separate hydrogen gas inlet.

4. The method for hydrotreating poor diesel fuel according to claim 1, characterized in that in step (2), the hydrocracking reaction zone is filled with a hydrocracking catalyst, and the hydrocracking catalyst is a porous catalyst, a shaped catalyst and/or a honeycomb catalyst; the pore diameter of the porous catalyst is 1-50 mm, preferably 4-20 mm; the average particle diameter of the special-shaped catalyst is 2-50 mm, preferably 4-30 mm; the honeycomb catalyst has a cell diameter or a cell side length of 1 to 50mm, preferably 3 to 15mm.

5. An inferior diesel oil hydrotreater characterized in that it comprises:

(2) Heterotypic hydrogenation ware, heterotypic hydrogenation ware include by lower supreme heavy oil bunker, reaction chamber and the separator that communicates in proper order, the reaction intracavity has set gradually from top to bottom: liquid distribution subassembly, catalyst bed and hydrogen distribution chamber.

6. The poor quality diesel oil hydrotreater according to claim 5, wherein the ratio of the equivalent diameter of the catalyst bed cross-sectional area of the profiled hydrogenation reactor to the total height of the catalyst bed in the reactor is (2).

7. The hydrotreater for poor quality diesel oil according to claim 5, wherein the cross section of the catalyst bed in the reaction chamber is rectangular or circular, and the height of the bed is 100-5000 mm, preferably 200-1000 mm.

8. The poor quality diesel oil hydrotreater according to claim 5, wherein a plurality of baffles are arranged in parallel in the vertical direction in the hydrogen distribution chamber, the plurality of baffles divide the hydrogen distribution chamber into a plurality of inlet units, and the bottom of each inlet unit is provided with at least one hydrogen inlet; a plurality of holes are distributed on each clapboard; the partition plate extends upwards to the catalyst layer, the opening rate of the partition plate below the catalyst layer is less than 70%, and the opening rate of the partition plate in the catalyst layer is more than 50%.

9. The poor quality diesel hydrotreater according to claim 5, wherein said liquid distribution assembly comprises a liquid distributor and a liquid distribution tray and a distribution cone disposed above said liquid distributor.

10. The poor quality diesel hydrotreater according to claim 9, wherein said liquid distribution tray is the same shape as the top surface of the catalyst bed, and the area of the liquid distribution tray is 10% to 100%, preferably 60% to 100%, of the cross section of the catalyst bed.

11. The hydrotreater for poor quality diesel oil according to claim 10, wherein said liquid distribution tray is uniformly provided with a plurality of first through holes, a first overflow ring is provided around the first through holes, an overflow part is provided at the outer edge of the liquid distribution tray, the opening ratio of the liquid distribution tray is 5-90%, the diameter of the first through holes is 5-100 mm, and the height of the first overflow ring is 1-30 mm.

12. The poor quality diesel hydrotreater as set forth in claim 10, wherein said first overflow ring has a serrated portion on its inner side, said serrated portion being curved downward and said serrated portion having a flow guide groove.

13. The poor quality diesel oil hydrotreater according to claim 9, wherein said distribution cone is disposed at the center of the upper portion of said liquid distribution tray, and is provided with a plurality of second through holes, and a second overflow ring is disposed around said second through holes; the vertex angle of the distribution cone is more than 90 degrees, the aperture ratio of the distribution cone is 5-80 percent, and the height of the second overflow ring is 1-30 mm; the bottom area of the distribution cone is 2-15% of the area of the liquid distribution disc.

14. The poor quality diesel oil hydrotreater of claim 5, wherein in step (2), the catalyst bed is a porous catalyst layer, a profiled catalyst layer and/or a honeycomb catalyst layer; the aperture of the porous catalyst is 1-50 mm, preferably 4-20 mm; the average particle diameter of the special-shaped catalyst is 2-50 mm, preferably 4-30 mm; the honeycomb catalyst has a cell diameter or a cell side length of 1 to 50mm, preferably 3 to 15mm.

15. The poor quality diesel oil hydrotreater of claim 5, wherein the heterogeneous hydrogenation reactor further comprises: multistage auxiliary reaction chamber, each level auxiliary reaction chamber advance hydrogen alone, the bottom center sets up heavy oil storehouse alone, and the liquid raw material import of each level auxiliary reaction chamber is connected with the heavy oil storehouse of last one-level, and the top of multistage auxiliary reaction chamber all is connected to the separator.