CN115785994A

CN115785994A - Wax oil hydrogenation method and device

Info

Publication number: CN115785994A
Application number: CN202111059067.9A
Authority: CN
Inventors: 郭兵兵; 赵玉琢; 宣根海; 徐彤; 刘涛
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 wax oil hydrogenation method and a wax oil hydrogenation device. The method comprises the following steps: the wax oil raw material is firstly separated into gas and liquid through hydrotreating, the liquid enters a hydrocracking reactor and is separated into a lighter part and a heavier part after passing through a liquid distribution assembly. Wherein the lighter part is not subjected to hydrocracking, and the heavier part is subjected to hydrocracking reaction, and then the obtained product is separated to obtain naphtha fraction and diesel fraction. The method controls the reaction depth in a more balanced way, finally achieves the aim of improving the yield of the target product, particularly improves the selectivity and the yield of naphtha, and can eliminate the risk of bed temperature runaway or flooding.

Description

Wax oil hydrogenation method and device

Technical Field

The invention belongs to the technical field of hydrogenation, and particularly relates to a wax oil hydrogenation method and device.

Background

Hydrocracking technology is one of the important heavy oil lightening technologies, and has become the standard configuration of most large oil refining chemical enterprises. Wax oil hydrogenation plays a more important role in Chinese oil refining enterprises, and a wax oil hydrogenation device mainly provides qualified raw material products for a catalytic cracking device. With the transformation from more Chinese oil refineries to chemical industry enterprises, a new hydrotreating technology will convert the inferior raw oil feedstock into light oil as chemical raw material such as naphtha and chain hydrocarbon distillate oil as much as possible, which is also one of the main technologies for solving the transformation in chemical industry refineries.

Two ways are mainly used for increasing the yield of naphtha fraction by wax oil hydrogenation, one way is to research and develop a hydrocracking catalyst with higher activity or achieve the purpose by a catalyst grading mode, for example, CN201110446028.4 sequentially comprises a protective agent, a demetallizing agent, a desulfurizing agent and a denitriding agent from top to bottom by optimizing the grading filling of the catalyst, a magnesia alumina spinel-macroporous alumina carrier is subjected to organosilicon surface modification, and an active component is loaded to improve the utilization efficiency of the active component so as to improve the activity of the catalyst and improve the adaptability of raw materials; CN201811113894.X is filled with a hydrocracking catalyst containing a modified Y molecular sieve in a hydrocracking reactor along the 1 st to the n-1 st catalyst bed layers in the material direction; the nth catalyst bed layer is filled with a hydroisomerization catalyst containing beta and/or ZSM series molecular sieves with stronger isomerization performance; by setting a reasonable catalyst grading mode and regulating the composition and structure of the catalyst in the grading bed layer, the yield of the heavy naphtha with high aromatic hydrocarbon and high latent aromatic hydrocarbon is increased, and the quality of the aviation kerosene and diesel oil products is improved.

The other way is to achieve the target of different target product yields by optimizing the process and adopting flexible raw material proportion. As in CN201711468997.3, the hydrotreated material of wax oil raw material is divided into two parts; one stream of materials is pumped out of the first hydrotreating reactor through the middle of the bed layer and is mixed with LCO to enter a second hydrotreating reactor for hydrogenation reaction; and the other part of the material continuously flows downwards through a hydrotreating catalyst bed layer at the lower part, and the obtained first hydrotreating reaction material and the second hydrotreating reaction material are subjected to gas-liquid separation and fractionation to obtain naphtha, diesel oil, hydrogenation heavy fraction and the like. Finally, different wax oil raw materials are flexibly processed, and the adaptability of the wax oil raw materials is improved; CN200910207389.6 proposes a wax oil hydrotreating and catalytic cracking combined method, the liquid phase after hydrotreating directly enters a catalytic cracking device without fractionation, the method can produce gasoline to the maximum, simultaneously, save equipment investment, fully utilize reaction heat energy and reduce production energy consumption. However, no matter what way, the traditional fixed bed hydrogenation mode is still adopted, and because of the overhigh bed layer total height-diameter ratio and the parallel flow mode of the traditional fixed bed, the wax oil raw material is unevenly reacted in the reactor, part of the raw material oil is excessively cracked, and part of the raw material oil is not properly hydrogenated and directly flows out of the reactor, which is not beneficial to improving the yield of the target product.

The traditional fixed bed reactor has the advantages of simple operation and mature technology; the main problem is that the reaction process is longer, namely the height and diameter are larger, part of raw oil is over-cracked, and the other part of raw oil does not reach the cracking depth because the catalytic site is occupied. The height-diameter ratio (the ratio of the total height to the diameter of a reactor bed) of a fixed bed in the traditional 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" mentioned that "ratio of height to diameter" of conventional fixed bed reactor is more than 4-9 "in order to avoid poor catalyst contact efficiency caused by uneven fluid distribution. CN201710974167.1 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 may 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 and heavy naphtha is limited, and the reactor is coked and blocked. In addition, the hydrogenation reactor has the risk of catalyst bed temperature runaway or flooding, and particularly for the catalyst with higher activity, the reactor temperature runaway speed is remarkable and can reach 500-600 ℃, and individually exceeds 800 ℃, so that the catalyst can be failed and scrapped under severe conditions.

The disadvantages of the fixed bed reactor not only inhibit the yield of the wax oil hydrogenation target product, but also seriously shorten the service life of the catalyst.

Disclosure of Invention

Aiming at the defects of the prior art, the invention aims to provide a wax oil hydrogenation method and a wax oil hydrogenation device, which control the reaction depth in a more balanced manner, finally achieve the aim of improving the yield of target products, particularly improve the selectivity and the yield of naphtha, and simultaneously eliminate the risk of bed layer temperature runaway or flooding.

The first aspect of the invention provides a wax oil hydrogenation method, which comprises the following steps:

(1) The raw material wax oil is separated into gas and liquid after being subjected to hydrotreating;

(2) The liquid generated in the step (1) enters a hydrocracking reactor, the upper part of the hydrocracking reactor is a separation zone, the lower part of the hydrocracking reactor is a hydrocracking reaction zone, the liquid is dispersed by a liquid distribution component, wherein the lighter part of the liquid directly enters the separation zone upwards without hydrocracking, the heavier part of the liquid enters the hydrocracking reaction zone downwards for hydrocracking reaction, the light product generated by hydrocracking reaction upwards leaves the hydrocracking reaction zone, and the heavy product generated by hydrocracking reaction downwards leaves the hydrocracking reaction zone; the gas generated in the step (1) enters the upper end of a separation zone of a hydrocracking reactor or is directly condensed and separated;

(3) And separating the material entering the separation zone to obtain gas, naphtha fraction, jet fuel fraction, diesel fraction and uncracked heavy components, wherein the uncracked heavy components enter the hydrocracking reaction zone again for hydrocracking reaction.

Further, in the above technical scheme, the hydrocracking reaction zone in step (2) includes 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 (hereinafter simply 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 increased occupies a great deal of bed gaps, so that the liquid phase flow rate 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, and for a system which needs more liquid phase conversion and controls the secondary reaction by the gas phase, the overall reaction effect is limited to a certain extent, and generally, the reaction conversion rate, the selectivity and the like are 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 amount rapidly reduced and the gas phase amount rapidly increased in the reaction process, when the hydrogen is contacted with the 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 probability of the liquid phase on the catalyst for reaction can be more sufficient, the traditional recognition that the small diameter-height ratio can bring the adverse effects such as poor contact effect is further overcome, the effect of obviously improving the yield of a target product (heavy naphtha or tail oil which is a target product in the inferior wax oil 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 heavy product generated by the hydrocracking reaction in the step (2) flows downwards from the bottom of the hydrocracking reactor, wherein at least part of the heavy product is recycled to the hydrocracking reaction zone as the hydrocracking feed.

Further, in the above technical solution, the hydrotreating in step (1) is performed in a one-stage fixed bed hydrotreating reactor, which contains a hydrotreating catalyst bed, and preferably, the hydrotreating catalyst bed has a hydrotreating catalyst bed in which a hydrotreating catalyst is placed at an upper portion and a hydrocracking catalyst is placed at a lower portion, and a volume ratio of the hydrotreating catalyst to the hydrocracking catalyst is 1-10, preferably 3. The raw oil can be desulfurized and denitrified by two different types of catalysts, aromatic hydrocarbon is saturated, and the polycyclic aromatic hydrocarbon is subjected to ring opening to generate monocyclic aromatic hydrocarbon with side chains. After the macromolecule chain hydrocarbon is properly cracked, the raw oil is changed into clean distillate oil with medium molecular size to be used as the raw material of the subsequent hydrocracking reaction.

Further, in the above technical solution, the hydrorefining catalyst may be any suitable hydrorefining catalyst, and the hydrorefining active component may be one, two or more selected from metals in group VIB and group viii, such as W, mo, co, and Ni.

Furthermore, in the above technical scheme, the hydrocracking catalyst can be any suitable hydrocracking catalyst, and preferably, the catalyst active component is one or more of metals in a VIB group and/or a VIII group. 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 molecular sieve, which can be Y-type molecular sieve.

Further, in the above technical solution, the wax oil raw material of the present invention is generally Vacuum Gas Oil (VGO), and may also include one or more of straight run wax oil (AGO), coker wax oil (CGO), catalytic cracking Heavy Cycle Oil (HCO), deasphalted oil (DAO), coal synthetic oil, and coal tar.

Further, in the above technical solution, the operation conditions of the hydrotreating in the step (1) are as follows: the average reaction temperature is 200-450 ℃, the reaction pressure is 3-20 MPa, and the volume ratio of hydrogen to oil is 200. The ratio of the flow of the raw oil to the volume of the catalyst in the reactor is 0.1h ^-1 ～10.0h ^-1 . Preferred operating conditions are: the average reaction temperature is 230-420 ℃, the reaction pressure is 4-18 MPa, and the volume ratio of hydrogen to oil is 300.

Further, the operation conditions of the hydrocracking reaction zone and the separation zone in the step (2) and the step (3) are as follows: the average reaction temperature is 200-450 ℃, the reaction pressure is 3-20 MPa, the volume ratio of hydrogen to oil is 200-3000 ^-1 ～10.0h ^-1 The ratio of inlet circulating heavy oil to raw oil is 1. The preferred operating conditions are: the average reaction temperature is 230-420 ℃, the reaction pressure is 4-18 MPa, the volume ratio of hydrogen to oil is 300.

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

Further, in the above technical solution, in the step (2), the hydrocracking reaction zone is filled with the hydrocracking catalyst, and the shape of the hydrocracking catalyst can be any conventional existing hydrocracking catalyst shape, and is preferably 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 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 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 wax oil in step (1) is heated by a heating furnace before entering the hydrotreating reaction zone, preferably heated to 200 ℃ to 450 ℃.

In a second aspect, the present invention provides a wax oil hydrotreating apparatus, including:

(1) The hydrotreating reactor is used for hydrotreating the wax oil;

(2) Heterotypic hydrogenation ware, heterotypic hydrogenation ware include by the supreme heavy oil storehouse, the hydrocracking reaction zone and the disengagement zone that communicate in proper order down, have set gradually in the hydrocracking reaction zone from top to bottom: 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 heterotype hydrogenation reactor to the total height of the catalyst bed in the reactor is (2-1), preferably (3-6). 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 zone in the reactor, with cross-section being taken as a top view cross-section, i.e. a section perpendicular to the vertical within the reaction zone. 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 zone throughout the catalyst bed.

Further, in the above technical scheme, the hydrocracking reaction zone is a horizontal storage tank, the axial direction of the hydrocracking reaction zone is transversely arranged, and two ends of the horizontal storage tank are provided with end sockets. Or the hydrocracking reaction zone is an oblate cylindrical tank which is axially and longitudinally arranged.

Further, in the above technical scheme, the cross section of the catalyst bed in the hydrocracking reaction zone is rectangular or circular, and the bed height 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 carried out of the reactor, thereby avoiding the blockage phenomenon of the traditional reactor caused by the overlong retention time of byproducts in the bed layer.

Further, in the above technical scheme, a plurality of clapboards are arranged in parallel along the vertical direction in the hydrogen distribution cavity and the catalyst bed layer in the hydrocracking reaction zone, the plurality of clapboards separate the hydrogen distribution cavity 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 at the center of the bottom of the reaction zone, 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 upper separation zone and the diameter (or the maximum equivalent diameter) of the lower hydrocracking reaction zone of the special-shaped hydrogenation reactor is 1.2-1, preferably 1. In the invention, the diameter of the upper separation area of the special-shaped hydrogenation reactor is reduced, so that the light fraction load under high pressure is completely matched with the tower plate, the tower plate separation efficiency is high, and the special-shaped hydrogenation reactor has the complete substitution of a fractionating tower.

Further, in the above technical scheme, the separation zone comprises 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 separation section, the height of the separation section is 50-65% of the total height of the separation section, and the height of the stabilizing section is 5-10% of the total height of the separation section.

Further, in the above technical scheme, a column plate 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 column plate can be one or more of bubble cap plate, sieve plate, floating valve plate, mesh plate, tongue plate, guide sieve plate, multi-downcomer column plate, etc., or can be a through-flow sieve plate, a 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 1-4 product lateral lines; the mixing section is provided with 1-3 side lines of light raw materials, and the mixing section is provided with 1-3 reaction zones.

Further, in the above technical solution, the liquid distribution assembly in the reaction area 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 in the invention is preferably a porous tubular distributor or a straight tube baffle distributor, and the diameter of the pore passage of the tubular distributor is 0.5-20 mm, preferably 2-10 mm. The farther from the raw 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 surface of the bed 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 above 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 in the hydrocracking reaction zone, and the area of the liquid distribution disc is 10% -100%, preferably 60% -100% of the cross section of the catalyst bed.

Further, in the above technical solution, a plurality of first through holes are uniformly formed on the distribution plate, a first overflow ring is arranged around the first through holes, and an overflow part is arranged at the outer edge of the distribution plate; 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 at 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 special-shaped hydrogenation reactor is a porous catalyst layer, a special-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 side 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 an 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 solution, the heterotypic hydrogenation reactor still includes: the multistage auxiliary reaction area, each level auxiliary reaction area advance hydrogen alone, the bottom center sets up heavy oil storehouse alone, and the liquid raw materials import of each level auxiliary reaction area is connected with the heavy oil storehouse of last one-level, and the top in multistage auxiliary reaction area all is connected to the disengagement zone.

The hydrogen feeding pipe at the bottom of the reaction section at the lower part of the reaction zone of the special-shaped hydrogenation reactor is provided with a plurality of inlets, and each hydrogen feeding pipe corresponds to the catalyst bed area between the two clapboards, so that hydrogen can upwards pass through the reaction zone 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 that 1 to 4 side lines are opened in the separation section of the separation zone for withdrawing the desired product. At the top of the 1 st side line or the separation zone, the extraction temperature is 60-95 ℃, and the extracted components are gas and light naphtha fraction. The 2 nd side draw temperature is from 150 ℃ to 190 ℃, preferably from 160 ℃ to 180 ℃, and the components drawn off in this side draw are heavy naphtha fractions. The 3 rd side draw temperature is from 180 ℃ to 250 ℃, preferably from 180 ℃ to 230 ℃, and the components drawn off at the side draw are jet fuel fractions. The 4 th side draw temperature is from 330 ℃ to 390 ℃, preferably from 350 ℃ to 380 ℃, and the components drawn off in this side draw are diesel fractions. In practical application, the number of the side lines required to be led out can be determined according to practical requirements, and is not limited to 4 side lines.

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 the liquid phase quantity rapidly reduced and the gas phase quantity rapidly increased under the reaction condition, breaks through the long-term design scheme in the field, and provides the technical means with design parameters completely different from the prior art. The method is characterized by comprehensively and deeply researching 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 providing 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 has outstanding technical effects on the aspect of producing heavy naphtha products with high added values in the maximum amount by hydrogenating wax oil, 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 curing cognition in the field is broken through.

According to the invention, through the arrangement of the design of the liquid distribution assembly, 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, so that the lighter part does not enter a hydrocracking reaction zone to participate in a hydrocracking reaction. Meanwhile, through the design of the liquid distribution component, heavy components entering a hydrocracking reaction zone can be distributed more uniformly, and the problem that reactants on a catalyst bed layer are poor in contact when a traditional reactor is in a 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 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 secondary hydrogenation quickly separate from the system and not excessively participate in cracking reaction by reasonably setting the reaction process and controlling the catalyst bed structure, and the product can quickly leave the reaction system, thereby increasing the positive reaction speed, eliminating the hidden trouble of catalyst blockage caused by-products, improving the yield of the target product and prolonging the service life of the catalyst.

The invention can realize timely extraction of light intermediate products through flash evaporation and steam stripping by arranging the separation zone, 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 is not formed.

The bottom tail oil of the hydrocracking reactor can be fully circulated to produce naphtha fraction in a large amount, can also be partially circulated, greatly improves the paraffin content of the tail oil product, reduces the aromatic hydrocarbon content, and is a high-quality ethylene material prepared by steam cracking.

Drawings

FIG. 1 is a schematic diagram of a wax oil hydrogenation process flow and a schematic diagram of a special-shaped hydrogenation reactor structure according to the invention;

wherein: 1-raw oil inlet; 2-new hydrogen; 3-hot raw oil; 4-hot hydrogen; 5-hydrogen oil mixed raw material; 6-a hydroprocessing reactor; 7-hydrotreating reactor bottom pipe; 8-a gas-liquid separation tank; 9-inlet of the special-shaped hydrogenation reactor; 10-special-shaped reaction zone of hydrogenation reactor; 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 separation zone; 20-special-shaped hydrogenation reactor top pipeline; 21-a condenser; 22-liquid separation tank; 23-hydrogen rich and non-condensable gases; 24-light naphtha fraction; 25-heavy naphtha fraction; 26-heavy oil fraction; 27-a circulating oil pump; 28-hydrogenated tail oil; 29-desulfurization deamination equipment; 30-a cyclic hydrogen press; 31-a heating furnace; 32-jet fuel fraction; 33-diesel fraction.

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

Fig. 3 is a schematic side view of a segmental baffle plate in accordance with another embodiment of the invention, the baffle plate being located only within the hydrogen distribution chamber and there being no segmental baffle plate 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 absent from the interior of 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 top and side view structural schematic of a liquid distribution assembly according to another embodiment of the present invention.

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

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

Detailed Description

The following detailed description of the present invention is provided in conjunction with 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 term "comprise" or variations such as "comprises" or "comprising", etc., 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 article 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 have other orientations (rotated 90 degrees or otherwise) and the spatially relative terms used herein should be interpreted accordingly.

In this document, the terms "first", "second", and the like 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 diagram of a wax oil hydrogenation process flow of the present invention. The wax oil is sent into a heating furnace 31 from a raw oil inlet 1, heated to 200-450 ℃, then converged with heated hot hydrogen 4 to form a hydrogen-oil mixed raw material 5, and then enters a hydrotreating reactor 6, a hydrofining 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 mixing section 16. The liquid heavy oil fraction enters a reaction zone 10 of the special-shaped hydrogenation reactor through an inlet 9 of the special-shaped hydrogenation reactor, and after passing through a liquid distribution component, small molecular component hydrocarbons dissolved in the heavy oil are further released into a separation zone 19 as light components. The rest heavy component distillate 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 separation zone 19 under the carrying of the hydrogen. Under the action of a separation tower plate of a separation section 17, macromolecular hydrocarbons carried into a separation zone are separated, the macromolecular hydrocarbons downwards enter a catalyst bed layer 11 of the heterotype hydrogenation reactor again to be cracked, micromolecular hydrocarbons upwards face upwards and can be sequentially used as a diesel fraction 33, a jet fuel fraction 32 and a heavy naphtha fraction 25 at a lateral line to be extracted and sent out as products, non-condensable gas is continuously cooled by a condenser 21, liquid hydrocarbons are refluxed or extracted to be used as a light naphtha fraction 24 after gas-liquid separation by a liquid separation tank 22, and the gas enters desulfurization and deamination equipment 29 to be purified and recycled. The heavy oil fraction 26 at the bottom of the reaction zone is returned to the inlet 9 of the special-shaped hydrogenation reactor by a circulating oil pump 27 or is extracted as hydrogenation tail oil 28. As an extension, the process flow of the present invention is also shown in FIGS. 1-2, but is not limited to these two.

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

Further, in one or more exemplary embodiments of the present invention, the shape of the partition matches the bottom of the reaction zone 10, and when the reaction zone 10 is a horizontal tank, the partition is a segmental partition, as shown in fig. 2 and 3; when the reaction zone 10 is a flat cylindrical tank, the plurality of baffles are coaxial annular baffles 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 advantageous for increasing the resistance, so that 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 separation zone 19 to reaction zone 10 is from 1.2 to 1.

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

Further, in one or more exemplary embodiments of the present invention, the liquid distribution assembly 14 between the separation zone 19 and the reaction zone 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 on 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 in the present invention. The vertex angle of the distribution cone is more than 90 degrees, the more gentle the vertex angle of the distribution cone is, the better, so that 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 separation area from flowing down from the first through hole at the center of the liquid distribution disc, so that the liquid is distributed more uniformly.

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 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 zone. It is to be understood that the auxiliary reaction zone may be multi-staged. Each level of auxiliary reaction zone independently enters hydrogen, the bottom center is independently provided with a heavy oil bin, the liquid raw material inlet of each level of auxiliary reaction zone is connected with the heavy oil bin of the upper level, and the top of each level of auxiliary reaction zone is connected to the separation zone 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 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.

Further, in one or more exemplary embodiments of the present invention, the non-condensable gas from the gas-liquid separation tank 8 enters the condenser 21, is subjected to gas-liquid separation in the knockout drum 22, and then enters the desulfurization and deamination device 29 for purification and recycling.

The following examples further illustrate a wax oil hydrogenation process that can improve 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 flow chart shown in figure 1 of the invention is adopted, the raw oil is atmospheric and vacuum wax oil, and the properties are shown in table 2. Heating the raw oil and hydrogen to 375 ℃ 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. 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 enters a separation zone mixing section of the special-shaped reactor for further separation. The liquid is used as the raw material of the special-shaped hydrogenation reactor. The special-shaped hydrogenation reactor is filled with a catalyst D, and a catalyst bed layer is used asThe gauge height ratio is 4:1. a catalyst bed layer is arranged in the reaction area 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. The catalyst bed layer is internally provided with 4 annular clapboards. 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 by a liquid distribution assembly. The light component enters the upper separation area of the special-shaped hydrogenation reactor upwards 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 cracking reaction with the heavy component which is separated from the light and heavy components of the raw material inlet through the liquid distribution component under the action of a catalyst together with the hydrogen which goes upwards from the bottom, the produced light fraction is quickly separated from the reaction system upwards and enters the mixing section of the separation zone, and the separated light fraction enters the separation section of the separation zone upwards to be subjected to naphtha fraction separation. Light components at the top of the special-shaped reactor are condensed by a condenser, pass through desulfurization and deamination equipment and then are recycled by a hydrogen compressor. The material extracted from the upper side line 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. And (3) 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 operation process conditions are shown in Table 3. The reaction conditions and product distribution are shown in Table 4, and the simulated bed temperature rise change is 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 tray is the same shape as the top surface of the catalyst bed, and the area of the liquid distribution tray 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 on the outer edge of the distribution disc, the opening rate 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 oil, and the equivalent diameter height ratio of the catalyst bed layer in the special-shaped hydrogenation reactor is 5:1, the number of the baffles in the catalyst bed is 6. 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 liquid distribution plate has an opening rate of 80%, the first through holes have a diameter of 20mm, and the first overflow ring has a height of 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 hydrofinishing catalyst a in the upper part of the hydrotreating 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 that the hydrogenation is carried outDifferent proportions of refined catalyst and cracking catalyst in the treatment 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 method is the same as example 1 except that the catalyst filled in the heterotype 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 zone of the special-shaped hydrogenation reactor like the 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

The conventional two-stage hydrogenation process, namely the refining and cracking process, is adopted. The refining reactor and the cracking reactor adopt a reaction process that the raw materials 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 was filled with a catalyst a and a catalyst C at a volume ratio of 3. The remaining process conditions were the same as in example 1.

TABLE 1 catalyst Properties

TABLE 2 Properties of the feed oils

Raw oil	Pressure-reducing wax oil	Coker gas oil	Straight-run wax oil and 27% diesel oil catalyst
				Density (20 ℃ C.), g/cm ³	0.918	0.914	0.934
Range of distillation range, deg.C	292～515	213～547	203～520
				w(S),％	2.6	1.1	1.412
w(N/μg·g ^-1	850	6100	2456
				Residual carbon content%	0.08	0.52	0.50

TABLE 3 Process conditions of examples and comparative examples

One-stage hydrogenation reactor
		Reaction temperature, deg.C	375
Reaction pressure, MPaG	14.7
		Volume hydrogen to oil ratio	850：1
Volumetric space velocity h ^-1	1.2
		Two-section special-shaped hydrogenation reactor
Reaction temperature of	370
		Reaction pressure, MPaG	14.2
Volume hydrogen to oil ratio	1000：1
		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 distribution of the examples and comparative examples was calculated in a laboratory by simulation using ansys version 19.0 software. The simulation conditions were input according to the actual data of examples and comparative examples. Simulation results show 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 temperature of the bed layer of the reactor is relatively uniform. The bed simulated temperature rise changes are shown in Table 5.

TABLE 5 simulated temperature rise change of special-shaped hydrogenation reactor bed

Temperature point of bed layer	Example 1	Example 2	Comparative example 1
				Maximum radial temperature difference, ° c	1.2	0.6	19.2
Maximum longitudinal temperature difference, ° c	0.8	1.1	17.2
				Average temperature of	371.9	371.3	388.7

As can be seen from the results in Table 5, the temperature difference of the catalyst bed layers in examples 1-2 using the present invention is significantly lower than that in comparative example 1, the maximum radial temperature difference is reduced from the 19.2 ℃ temperature difference of the conventional fixed bed to 0.6 ℃, the maximum longitudinal temperature difference is reduced from 17.2 ℃ to 0.8 ℃ of the conventional reactor, and the difference between the average temperature and the control temperature in the examples is the smallest, which indicates that the reactor of the present invention has eliminated the overheating phenomenon of the hydrocracking reaction.

Claims

1. A wax oil hydrogenation method is characterized by comprising the following steps:

2. The method of claim 1, wherein: in the step (2), the hydrocracking reaction zone comprises 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 is (2) - (1-10), preferably (3).

3. The method of claim 1, wherein: in the step (2), the liquid is dispersed into small liquid drops after passing through the liquid distribution component, 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.

4. The method of claim 1, wherein: the hydrotreating in the step (1) is carried out in a first-stage fixed bed hydrotreating reactor, wherein the hydrotreating reactor contains a hydrotreating catalyst bed, a hydrofining catalyst is placed on the upper part of the hydrotreating catalyst bed, a hydrocracking catalyst is placed on the lower part of the hydrotreating catalyst bed, and the volume ratio of the hydrofining catalyst to the hydrocracking catalyst is 1-10, preferably 3.

5. The method of claim 1, wherein: the hydrotreating operation conditions in step (1) are as follows: the average reaction temperature is 200-450 ℃, the reaction pressure is 3-20 MPa, the volume ratio of hydrogen to oil is 200-3000 ^-1 ～10.0h ^-1 (ii) a Preferred operating conditions are: the average reaction temperature is 230-420 ℃, the reaction pressure is 4-18 MPa, and the hydrogen-oil volume ratio is 300.

6. The method of claim 1, wherein: the hydrocracking reaction zone and the separation zone are operated under the following conditions: the average reaction temperature is 200-450 ℃, the reaction pressure is 3-20 MPa, the volume ratio of hydrogen to oil is 200The volume ratio of the agent is 0.1h ^-1 ～10.0h ^-1 The ratio of inlet circulating heavy oil to raw oil is 1; preferred operating conditions are: the average reaction temperature is 230-420 ℃, the reaction pressure is 4-18 MPa, the volume ratio of hydrogen to oil is 300.

7. The method of claim 1, wherein: in the step (2), the hydrocracking reaction zone is longitudinally divided into a plurality of sub-reaction zones, so that when hydrogen rises to the bottom of the bed layer through the distributor, the diffusion area of the hydrogen can cover the sectional area of the bed layer at the bottom of the sub-reaction zones, and each sub-reaction zone is provided with an independent hydrogen inlet.

8. The method of claim 1, wherein: in the step (2), a hydrocracking reaction area is filled with a hydrocracking catalyst, and the hydrocracking catalyst is a porous catalyst, a special-shaped catalyst and/or a honeycomb catalyst according to the shape; 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 pore is 1-50 mm, preferably 3-15 mm; the voidage of the catalyst bed is 15-85%, preferably 20-75%.

9. A wax oil hydrotreating apparatus for carrying out the process of any one of claims 1 to 8, characterized in that: the device comprises:

(1) The hydrotreating reactor is used for hydrotreating the wax oil;

10. The apparatus of claim 9, wherein: 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).

11. The apparatus of claim 9, wherein: the cross section of the catalyst bed layer in the hydrocracking reaction zone is rectangular or circular, and the bed layer height is 100-5000 mm, preferably 200-1000 mm.

12. The apparatus of claim 9, wherein: a plurality of clapboards are arranged in parallel along the vertical direction in the hydrogen distribution cavity and the catalyst bed layer in the hydrocracking reaction zone, the clapboards divide the hydrogen distribution cavity 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%.

13. The apparatus of claim 12, wherein: the bottom of the catalyst bed layer 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 to the bottom of the regional bed layer covers the bottom of the whole regional bed layer.

14. The apparatus of claim 9, wherein: the diameter ratio of the upper end separation zone and the lower hydrocracking reaction zone of the special-shaped hydrogenation reactor is 1.2-1, preferably 1.

15. The apparatus of claim 9, wherein: the separation zone comprises a mixing section, a separation section and a stabilizing section from bottom to top, wherein the height of the mixing section is 25-40% of the total height of the separation zone, the height of the separation section is 50-65% of the total height of the separation zone, and the height of the stabilizing section is 5-10% of the total height of the separation zone.

16. The apparatus of claim 9, wherein: the liquid distribution component in the reaction zone 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.

17. The apparatus of claim 16, wherein: 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.

18. The apparatus of claim 17, wherein: the inner side of the first overflow ring is provided with a sawtooth part which bends downwards, and the sawtooth part is provided with a diversion trench.

19. The apparatus of claim 16, wherein: 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 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.