CN116948673B

CN116948673B - A device for producing low aromatic solvent oil and a hydrogenation method

Info

Publication number: CN116948673B
Application number: CN202210415867.8A
Authority: CN
Inventors: 刘涛; 马蕊英; 郭兵兵; 徐彤; 毕文卓
Original assignee: Sinopec Dalian Petrochemical Research Institute Co ltd; China Petroleum and Chemical Corp
Current assignee: Sinopec Dalian Petrochemical Research Institute Co ltd; China Petroleum and Chemical Corp
Priority date: 2022-04-20
Filing date: 2022-04-20
Publication date: 2024-12-06
Anticipated expiration: 2042-04-20
Also published as: CN116948673A

Abstract

The present invention discloses a device and method for producing low aromatic solvent oil by hydrogenating bio-oil. Bio-oil and Fischer-Tropsch oil are used as raw materials, and a treatment mode combining hydrodeoxygenation, hydrofining and hydrocracking is adopted; wherein, the bio-oil hydrodeoxygenation product and the Fischer-Tropsch oil hydrofining liquid phase enter the hydrocracking section and are countercurrently contacted with hydrogen to carry out a hydrocracking reaction, and the lighter part is carried away from the reactor by hydrogen and the light product obtained from the hydrocracking section leaves the hydrogenation reactor for separation. The method of the present invention allows the light component to quickly leave the system and no longer excessively participate in the cracking reaction. Since the product can quickly leave the reaction system, the positive reaction rate is increased while the yield of the target product, low aromatic solvent oil, is improved.

Description

Device for producing low aromatic hydrocarbon solvent oil and hydrogenation method

Technical Field

The invention belongs to a hydrogenation method in the technical field of oil refining, relates to a hydrogenation distillation method, and in particular relates to a device and a hydrogenation method for producing low aromatic hydrocarbon solvent oil from biological oil and Fischer-Tropsch synthetic oil through a hydrogenation technology.

Background

With worldwide carbon neutralization related policies falling to the ground, the economic status of green recycling is highlighted. Biological grease belongs to renewable resources, is one of important development directions, and is widely regarded in the world. The biological grease can remove oxygen in the biological grease by a hydrogenation method to obtain hydrocarbon products, such as motor fuels for producing aviation kerosene, diesel oil and the like, namely the biological grease is subjected to hydrodeoxygenation (all or part of removal) to produce products meeting the motor fuel standard, and the method can directly meet the requirements of the existing market, for example, low-carbon alkane is produced, namely the biological grease is subjected to hydrodeoxygenation and then hydrocracking, and macromolecules are cracked into small molecules to produce the low-carbon alkane meeting the product requirements.

The Fischer-Tropsch synthesis reaction is a technological process using hydrogen and carbon monoxide as raw materials and using the Fischer-Tropsch synthesis reaction to generate liquid fuel as main material under the action of a catalyst, and the raw materials are mainly coal and natural gas. The Fischer-Tropsch synthetic oil mainly comprises normal alkane, alkene and a certain amount of oxygen-containing compounds, and the content of non-ideal components such as sulfur, nitrogen, aromatic hydrocarbon and the like is extremely low. The Fischer-Tropsch synthetic oil has a wider range, and different boiling range fractions obtained by fractionation are basically free of sulfur and nitrogen, and have the highest olefin content and oxygen content, so that the Fischer-Tropsch synthetic oil is subjected to corresponding hydrotreatment to obtain petrochemical raw materials or transportation fuels meeting the use specifications.

The prior vegetable oil hydrogenation technology, US5705722 discloses a diesel blending component in the range of producing diesel fraction by mixing vegetable oil containing unsaturated fatty acid, fat and the like with animal oil and then hydrogenation. CN101462915 discloses a preparation method of C ₆-C₁₂ alkane, which mainly uses animal and vegetable oil as raw oil to directly carry out hydrocracking to produce alkane, and because pretreatment is not carried out, the produced water has adverse effect on the hydrocracking catalyst molecular sieve, and the operation period is obviously shortened. EP1741767 and EP1741768 disclose a method for producing a diesel fraction from animal and vegetable oils, mainly by hydrotreating animal and vegetable oils first and then passing through an isomerization catalyst bed to obtain a low-condensation-point diesel component, but since water is produced during hydrotreating, the isomerization catalyst is adversely affected, and the device cannot be stably operated for a long period.

The prior art ,CN103102922、CN103102898、CN103102899、CN103102900、CN103102901、CN103102902、CN103102903、CN103102907、CN103102908、CN103102911、CN103102918、CN103102919 、CN103102920 for producing solvent oil by hydrogenating vegetable oil discloses a method for producing low aromatic hydrocarbon solvent oil by using biological oil, which comprises the steps of firstly carrying out hydrodeoxygenation reaction on biological oil raw materials under hydrogenation conditions, then carrying out hydrocracking on the obtained product oil under hydrocracking conditions after dehydrating, and cracking macromolecules into micromolecules to produce the low aromatic hydrocarbon solvent oil.

The general Fischer-Tropsch synthetic oil hydrotreating process flow is that after Fischer-Tropsch synthetic oil (comprising naphtha fraction, diesel fraction and heavy oil fraction) is mixed with hydrogen, the mixture is first fed into a hydrofining reactor, olefin saturation is carried out on a catalyst bed, and hydrodeoxygenation reaction of an oxygen-containing compound is carried out, and then the mixture is fed into a fractionation system. In the fractionating tower, naphtha, diesel oil and heavy oil are separated from the distillation range of the hydrofining product. The heavy oil continues to enter the subsequent hydrocracking reactor, the cracked product enters the second set of fractionation system, and the tail oil is partially or fully recycled back to the hydrocracking reactor, so as to obtain more naphtha and diesel fractions through the cracking reaction. The processes all adopt a traditional fixed bed hydrogenation mode, the two parts of hydrotreating and hydrocracking are completely separated, and two sets of fractionation systems are arranged in total, so that the process is complex, the device investment is high and the energy consumption is high.

The conventional fixed bed hydrocracking mode is still adopted in the heavy oil hydrocracking reaction in the Fischer-Tropsch synthesis oil at present, the raw oil and the hydrogen flow into the reactor together in parallel, and the materials flow downwards while reacting in the bed propelling process. In the whole reaction process, gas-phase reaction products cannot leave the reactor in time, so that the gas-phase reaction products occupy the pore channels of the catalyst, and the cracking reaction occurs again, so that the selectivity of the reaction and the yield of target products are reduced.

The current general technological process of vegetable oil hydrogenation is that vegetable oil is mixed with hydrogen, then enters into a hydrodeoxygenation reactor, and is subjected to olefin hydrogenation saturation and oxygen-containing compound hydrodeoxygenation reaction on a catalyst bed layer, and then is separated and dehydrated. The dehydrated fraction enters a subsequent hydrocracking reactor, the cracked product enters a fractionating system to obtain light alkane, and unconverted tail oil can be partially recycled or fully recycled back to the hydrocracking reactor, so that more low aromatic solvent oil is obtained through the cracking reaction. The processes all adopt a traditional fixed bed hydrogenation mode, and particularly, a plurality of catalyst beds are arranged in a hydrocracking reactor, so that adverse effects are caused, mainly that target products obtained by the reaction of contacting with the hydrocracking catalyst at the beginning can continuously enter the subsequent hydrocracking catalyst beds, and the generated target products can continuously undergo the hydrocracking reaction to obtain smaller molecular alkanes. That is, the selectivity of the target product is poor when the same conversion rate is achieved, or the yield of the target product is low when the raw materials are all converted.

Disclosure of Invention

Aiming at the defects of the prior art, the invention aims to provide a device and a hydrogenation method for producing low aromatic hydrocarbon solvent oil by hydrocracking biological grease and Fischer-Tropsch synthetic oil. The method improves the selectivity of the target product of the low aromatic hydrocarbon solvent oil by using a method that the target product is continuously removed from a reaction zone by using a hydrogenation distillation reactor, and can reduce the chemical hydrogen consumption when the yield of the same target product is achieved.

The first aspect of the invention provides a hydrogenation device for producing low aromatic solvent oil from biological grease and Fischer-Tropsch synthetic oil.

The hydrogenation device comprises:

(1) A hydrodeoxygenation reactor for hydrodeoxygenation of a raw material bio-oil;

(2) A hydrofining reactor for carrying out hydrofining reaction on Fischer-Tropsch synthesis oil;

(3) The hydrogenation reaction distillation tower receives liquid phase products and Fischer-Tropsch synthetic oil hydrofining products which are obtained after biological oil hydrodeoxygenation as liquid phase feed, wherein the raw material feed flows through each catalyst bed layer at the middle upper part of the distillation tower from top to bottom, preheated hydrogen uniformly enters each catalyst bed layer from bottom to top, the liquid phase feed and the hydrogen are subjected to hydrocracking reaction in hydrocracking catalyst beds filled in the distillation tower, light components generated by the hydrocracking reaction and gas phase feed which do not participate in the reaction are discharged from the top of the hydrogenation reaction distillation tower from independent gas phase channels, and liquid products obtained by the hydrocracking are discharged from the bottom of the hydrogenation reaction distillation tower;

(4) And the separation unit is used for separating light components generated by hydrocracking and gas-phase feed which does not participate in the reaction, and obtaining hydrogen-rich gas, dry gas, liquefied gas and low aromatic solvent oil.

Further, in the above technical scheme, the device further comprises (5) a gas circulation unit for selectively concentrating the hydrogen-rich gas obtained by the separation unit and circulating the hydrogen-rich gas back to the hydrodeoxygenation reactor, the hydrofining reactor and the hydrogenation reaction distillation tower.

Further, in the above technical scheme, the apparatus further comprises (6) a liquid circulation unit for circulating at least a part of the liquid product obtained from the hydrogenation reaction distillation column back to the hydrogenation reaction distillation column.

In the technical scheme, the hydrogenation reaction distillation tower comprises at least one hydrogenation reaction zone, wherein the hydrogenation reaction zone is of a 2~n-layer tower plate structure, each layer of tower plate is filled with a hydrocracking catalyst bed, n is an integer greater than 2, an inclined baffle plate, preferably an umbrella-shaped baffle plate, is arranged above each catalyst bed, and each catalyst bed and each inclined baffle plate are provided with gas phase channels.

In the technical scheme, the hydrodeoxygenation reaction in the hydrodeoxygenation reactor comprises a hydrodesaturation reaction of olefins contained in the biological oil raw material oil, a hydrodeoxygenation reaction of carboxyl and carbonyl contained in the biological oil raw material oil, and the like.

In the technical scheme, the hydrofining reaction in the hydrofining reactor comprises a hydrogenation saturation reaction of olefin contained in the Fischer-Tropsch synthesis oil raw material oil, a hydrodeoxygenation reaction of hydroxyl carboxyl contained in the Fischer-Tropsch synthesis oil raw material oil, a hydrodeimpurity reaction and the like.

Further, in the above technical solution, the hydroreactive distillation column further includes a liquid phase feed subunit disposed above an inclined partition of the topmost catalyst bed, through which liquid phase feed is directed to the catalyst bed.

Further, in the above technical scheme, the end of the inclined partition plate is provided with an annular downcomer, and the bottom of the annular downcomer is spaced from the bottom of the catalyst bed by a distance, so that the liquid phase feed enters the catalyst bed along the radial direction.

Further, in the above technical solution, the hydrogenation reaction distillation tower further includes a gas phase feeding subunit, which is disposed between the catalyst bed layer of the upper layer and the inclined partition plate of the lower layer. The gas phase feed is directed upward into the catalyst bed of the upper layer.

Further, in the above technical scheme, the lower end of the gas phase passage passing through the previous catalyst bed is connected with the inclined partition plate on the next catalyst bed. The design ensures that the gas phase channel and the gas phase feeding subunit are in a relative isolation state, and gas phase products generated after the gas phase feeding and the liquid phase feeding react in the catalyst bed layer directly enter the gas phase channel.

In the technical scheme, the catalyst bed is provided with an overflow weir arranged at one side close to the gas phase channel, and a liquid sealing baffle plate arranged at the upper part of the overflow weir and used for isolating gas phase feeding from gas phase products.

Further, in the technical scheme, the liquid sealing baffle comprises a horizontal part and a vertical part, wherein the horizontal part is in an annular flat plate shape and is positioned above the overflow weir, the vertical part is in a cylinder shape, the vertical part and the horizontal part are integrally formed, and the lower end of the vertical part is separated from the bottom of the catalyst bed by a certain distance.

In the technical scheme, the liquid-phase feeding subunit further comprises a liquid-phase feeding pipe which extends along the radial direction of the catalyst bed layer, and a liquid-phase distributing pipe which is annular and is orthogonal or tangential to the liquid-phase feeding pipe, wherein the pipe wall of the liquid-phase distributing pipe is provided with a plurality of pore canals for uniformly distributing liquid-phase feeding materials to all directions of the annular downcomer.

In the technical scheme, the gas-phase feeding subunit further comprises a gas-phase feeding pipe which extends along the radial direction of the catalyst bed layer, and a gas-phase distributing pipe which is annular or multi-layer concentric annular, wherein the gas-phase distributing pipe is orthogonal or tangentially intersected with the gas-phase feeding pipe, and a plurality of pore channels are arranged on the wall surface of the gas-phase distributing pipe and are used for uniformly distributing gas-phase feeding to all directions of the bottom of the catalyst bed layer.

In the technical scheme, the gas-phase feeding subunit can further comprise a gas-phase distributing disc which is positioned at the bottom of the catalyst bed layer and is in a disc shape as a whole, and a plurality of holes are uniformly distributed on the gas-phase distributing disc.

Further, in the above technical scheme, the gas phase channel is located in the middle of the hydrogenation reaction unit and penetrates through all the catalyst beds from bottom to top.

In the technical scheme, the uppermost end of the gas phase channel is discharged out of the hydrogenation reactor, the lowermost end of the gas phase channel is close to the bottom of the hydrogenation reactor, the port is required to be immersed in the liquid phase product, and the bottom of the hydrogenation reactor is provided with the liquid level monitoring unit.

Further, in the above technical scheme, the height of the packing or the catalyst bed layer is set to be 10mm to 1000mm according to the different reaction systems.

Further, in the technical scheme, the upper edge of the overflow weir can be 10 to 100mm higher than the upper surface of the catalyst of the bed.

Further, in the above technical solution, the gas-phase distribution pipe may be disposed below the catalyst bed or within the catalyst bed.

Furthermore, in the above technical scheme, gas-liquid separation devices are required to be arranged between the hydrodeoxygenation reactor and the hydrogenation reaction distillation tower and between the hydrofining reactor and the hydrogenation reaction distillation tower so as to remove water generated by hydrodeoxygenation reaction. A gas-liquid separation device can also be arranged, and the hydrodeoxygenation material flow and the hydrofinishing material flow are combined and then enter the separation device for dehydration.

The second aspect of the invention is to provide a method for producing low aromatic hydrocarbon solvent oil by hydrogenating biological grease and Fischer-Tropsch synthetic oil.

The method comprises the following steps:

(1) The biological oil raw material is mixed with hydrogen and then enters a hydrodeoxygenation reactor, and hydrogenation reaction is carried out through a hydrodeoxygenation catalyst bed under hydrodeoxygenation conditions;

(2) The Fischer-Tropsch synthesis oil raw material is mixed with hydrogen and then enters a hydrofining reactor, and hydrogenation reaction is carried out through a hydrofining catalyst bed layer under the hydrofining condition;

(3) The mixed liquid products of the material flow obtained in the step (1) and the hydrofining material flow obtained in the step (2) after dehydration enter a hydrogenation reaction distillation tower, and enter a plurality of catalyst beds in sequence in the tower in a descending manner to contact hydrogen entering the catalyst beds upwards for hydrogenation cracking reaction;

(4) And (3) separating the gas-phase product obtained in the step (3) in a separation unit to obtain hydrogen-rich gas, dry gas, liquefied gas and low-aromatic solvent oil.

Further, in the technical scheme, the method further comprises a step (5) of returning the hydrogen-rich gas obtained in the step (4) to the step (1) and/or the step (2) after optional concentration.

Further, the technical scheme also comprises a step (6) that the liquid product obtained in the step (4) is recycled to the step (1) and/or the step (2).

In the above technical solution, at least one of the biological oil and fat raw material vegetable oil and animal oil and fat in the step (1). The vegetable oil comprises one or more of soybean oil, peanut oil, castor oil, rapeseed oil, corn oil, olive oil, palm oil, coconut oil, tung oil, linseed oil, sesame oil, cottonseed oil, sunflower seed oil, rice bran oil, etc. The animal fat comprises one or more of adeps bovis seu Bubali, adeps Sus Domestica, adeps Caprae Seu Ovis and fish oil. The bio-oil material may be selected from an acid oil, a waste kitchen oil, and the like.

In the technical scheme, the hydrodeoxygenation reactor in the step (1) and the hydrofining reactor in the step (2) are conventional fixed bed reactors. The hydrogenation pretreatment operation conditions are generally reaction pressure of 3.0-20.0 MPa, hydrogen oil volume ratio of 200:1-3000:1, volume space velocity of 0.1h ^-1~6.0h^-1, average reaction temperature of 180-450 ℃, and preferable operation conditions are reaction pressure of 3.0-18.0 MPa, hydrogen oil volume ratio of 300:1-2500:1, volume space velocity of 0.2h ^-1~4.0h^-1 and average reaction temperature of 200-440 ℃.

In the technical scheme, the hydrogenation protecting agent which is firstly passed through by the reaction materials can be selected from conventional protecting agents, only one protecting agent can be selected, the granularity of the particles can be changed from large to small according to the contact sequence with the raw oil, the content of hydrogenation metal is changed from small to large, two or more protecting agents are selected from weak to strong in hydrogenation performance, and a protecting agent system is formed by multiple protecting agents. For example, three protectant systems developed by FRIPP for FZC-100/FZC-105/FZC-106 (10/35/55 by volume) may be used herein. The total volume space velocity of the fresh feed of the protectant is generally 3.0h ^-1~30.0h^-1, preferably 4.0h ^-1~20.0h^-1.

Further, in the above-described embodiments, one, preferably two or more, hydrodeoxygenation catalysts may be used by continuing the flow of the hydroprotectant through the hydrodeoxygenation catalysts. The hydrodeoxygenation catalyst generally comprises Mo, W, ni, co of VIB group and VIII group, and the content of the hydrodeoxygenation catalyst calculated by the weight of oxide is 3% -20%, preferably 3% -15%, and more preferably 3% -10%. If two or more hydrodeoxygenation catalysts are used, the first hydrodeoxygenation catalyst through which the reaction mass first passes is 10% to 80%, preferably 20% to 70%, most preferably 30% to 60% by volume of all hydrodeoxygenation catalysts. In the two adjacent hydrodeoxygenation catalyst beds, the content of the hydrogenation active components in the upstream hydrodeoxygenation catalyst is 2-10 percent lower than that of the hydrogenation active components in the downstream hydrodeoxygenation catalyst in terms of weight of oxides, and is preferably 3-8 percent lower. The number of hydrodeoxygenation catalyst beds can be generally 2-5.

Furthermore, in the above technical scheme, in order to better play the protection function of the hydrogenation protecting agent and the hydrogenation function of the hydrodeoxygenation catalyst, the hydrogenation protecting agent and the hydrodeoxygenation agent need to be used together, and the volume ratio of the hydrogenation protecting agent to the hydrodeoxygenation agent is 15:85-70:30.

Further, in the above technical scheme, the carrier of the hydrodeoxygenation catalyst is generally alumina, amorphous silica-alumina, silica, titania and the like, and may contain other auxiliary agents such as P, si, B, ti, zr and the like. Commercially available catalysts may be used or may be prepared according to methods known in the art. The hydrogenation active component is in an oxidized state, and conventional vulcanization treatment is carried out before the catalyst is used, so that the hydrogenation active component is converted into a vulcanized state. As the commercial hydrogenation catalyst, there are mainly FHUDS series such as FF-24, FF-36, FF-46, FF-56, FF-66, FHUDS-6, FHUDS-7, FHUDS-8, etc., FZC series such as FZC-31, FZC-41, FZC-401, etc., hydrogenation catalysts such as HR-416, HR-448, etc. of IFP company, hydrogenation catalysts such as ICR174, ICR178, ICR 179, etc. of CLG company, hydrogenation catalysts such as HC-P, HC-KUF-210/220, TK-525, TK-555, TK-557, etc. of Topsor company, and hydrogenation catalysts such as KF-752, KF-840, KF-848, KF-901, KF-907, etc. of AKZO company.

Further, in the above technical scheme, the fischer-tropsch synthesis oil in step (2) refers to a fraction obtained by a fischer-tropsch synthesis technology, and may be a whole fraction, a light fraction, or a heavy fraction. The final distillation point of the Fischer-Tropsch oil is preferably not higher than 750 ℃, preferably not higher than 720 ℃.

In the technical scheme, a fixed bed hydrogenation reactor is generally used in the Fischer-Tropsch synthesis oil hydrofining process in the step (2), and hydrofining operation conditions are that the average reaction temperature is 150-450 ℃, the system reaction pressure is 3.0-20.0 MPa, the hydrogen-oil volume ratio is 100-1000, and the liquid hourly space velocity is 0.4h ^-1～10.0h^-1. The preferable operation condition is that the average reaction temperature is 180-430 ℃, the reaction pressure is 3.0-18.0 MPa, the hydrogen-oil volume ratio is 200-1000, and the liquid hourly space velocity is 0.5h ^-1～8.0h^-1.

Further, in the above technical scheme, the reaction pressure in the step (1) may be the same as or different from the reaction pressure in the step (2). If the reaction pressure in the step (1) can be the same as that in the step (2), the two parts can share one set of circulating hydrogen system, or one set of circulating hydrogen system can be used alone.

In the above technical solution, the number of the catalyst beds in the step (3) is n, and the upper portion of each catalyst bed is preferably filled with a solid filler, and the lower portion is preferably filled with a hydrocracking catalyst. The filler is in a conventional form in the field, for example, one or more random fillers such as pall rings, raschig rings, saddle-shaped, open pore ring types, semi-rings, ladder rings, double arcs, halfpace rings, conjugated rings, flat rings, flower rings and the like can be selected, and the filler can also be selected from metal or ceramic corrugated fillers. The hydrocracking catalyst generally comprises an active component and a carrier, wherein the carrier component comprises one or more of alumina, silicon-containing alumina and a molecular sieve, preferably contains the molecular sieve, the molecular sieve can be a Y-type molecular sieve agent, the active component comprises one or more of VIB group and VIII group metals, the VIB group metals are generally Mo and/or W, and the VIII group metals are generally Co and/or Ni. The hydrocracking catalyst shape may be any conventional existing hydrocracking catalyst shape, preferably porous, shaped and/or honeycomb catalysts. 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 pore diameter or pore side length of the honeycomb catalyst is 1-50 mm, preferably 3-15 mm, and the void ratio of the catalyst bed is recommended to be 15-85%, preferably 20-75%.

In the technical scheme, the operation condition of the hydrogenation reaction distillation tower in the step (3) is that the reaction temperature is 260-450 ℃, the reaction pressure is 3.0-20.0 MPa, the hydrogen-oil volume ratio is 100-2000, and the liquid hourly space velocity is 0.1h ^-1～10.0h^-1. The preferable operation condition is that the reaction temperature is 300-450 ℃, the reaction pressure is 3.0-18.0 MPa, the hydrogen-oil volume ratio is 100-1500, and the liquid hourly space velocity is 0.5h ^-1～10.0h^-1.

In the technical scheme, after hydrodeoxygenation in the step (1), hydrodeoxygenation products can be selectively separated, and liquid fractions obtained by separation enter a hydrogenation reaction distillation tower for further hydrogenation treatment.

In the technical scheme, after the hydrofining in the step (2), water generated by the hydrogenation reaction can be selectively separated, and the separated liquid fraction enters a hydrogenation reaction distillation tower for further hydrogenation treatment.

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

According to research, when the overall airspeed is similar, aiming at the gas-liquid-solid three-phase hydrogenation reaction with the rapid decrease of the liquid phase and the rapid increase of the gas phase in the reaction process, the generated gas phase rapidly leaves the catalyst bed, the adverse effect accumulation effect of the generated gas phase is small, the liquid phase can have more sufficient probability of reacting on the catalyst, the traditional recognition that the small height-diameter ratio can bring adverse effects such as poor contact effect is overcome, the effect of obviously improving the yield of the target product is obtained, and the problems that the countercurrent reactor is easy to be flooded, the hydrogen-oil ratio is limited are solved.

Compared with the prior art, the invention has the advantages that:

1. The gas phase product generated by the hydrocracking reaction can leave the cracking reaction zone in time, and can not enter the subsequent catalyst bed again, so that the gas phase product is prevented from occupying the channels of the cracking catalyst, the secondary cracking and gasification of the target product are effectively prevented, and the reaction selectivity and the yield of the target product are improved. Meanwhile, the partial pressure of the product is kept in a low state all the time, so that the driving force of the reaction is increased, and the equilibrium conversion rate is improved.

2. The components of the biological grease after hydrodeoxygenation contain more than 90 percent of normal paraffins, and the balance is isoparaffin. The components of the Fischer-Tropsch synthetic oil after hydrofining contain more than 90 percent of normal paraffins, and the balance of isoparaffins. The gas phase product generated after alkane cracking in the two hydrogenation products is rich in a large amount of low-carbon alkane, and the generated low-carbon alkane is carried out of the distillation tower through a gas phase channel and is hardly cracked into smaller molecules on a subsequent cracking catalyst, so that the load of a subsequent cracking reaction zone is increased.

3. The umbrella-shaped baffle plate arranged in the hydrogenation distillation reaction tower can separate gas phase feeding and product gas between adjacent bed layers on one hand, and plays a role in guiding liquid phase and gas phase on the other hand. The liquid sealing baffle can effectively isolate the gas phase feeding material from the gas phase product. The arrangement of the multi-layer concentric annular gas phase distributing pipe can keep the distribution of the gas phase feeding material uniform to the maximum extent.

Drawings

FIG. 1 is a schematic flow chart of one embodiment of the present invention for producing a low aromatic hydrocarbon solvent oil by hydro-distillation.

FIG. 2 is a schematic structural view of one embodiment of the hydrogenation reaction distillation column of the present invention.

FIG. 3 is a top view of the liquid distribution tube in the hydrogenation reaction distillation column of the present invention.

FIG. 4 is a top view of the gas phase feed line and gas phase distribution line in the hydroreactive distillation column of the present invention (showing the case where the gas phase feed line is disposed orthogonally to the annular gas phase distribution line).

FIG. 5 is another top view of the vapor feed line and vapor distribution line in the hydroreactive distillation column of the present invention (showing the vapor feed line tangentially intersecting the annular vapor distribution line).

FIG. 6 is a top view of a gas distribution tube of the present invention employing concentric dual ring distribution tubes.

FIG. 7 is a top view of a vapor distribution tray in a hydrogenation reaction distillation column according to the invention.

FIG. 8 is a top view of a catalyst support tray in a hydroreactive distillation column of the present invention.

The main reference numerals illustrate:

1-hydrogenation reaction distillation tower, 2-biological grease, 255-Fischer-Tropsch oil, 3-new hydrogen, 4-high-pressure hydrogen-rich gas, 5-hydrodeoxygenation reactor, 55-hydrofining reactor, 6-hydrodeoxygenation stream, 7-liquid phase stream, 8-gas phase stream, 9-high-pressure separator, 10-fractionating tower, 101-gas, 102-naphtha, 103-low aromatic hydrocarbon solvent oil and 104-unconverted oil.

11-Umbrella-shaped partition plates, 110-packing, 111-solid catalyst, 12-outside downcomers, 13-gas phase channels, 14-overflow weirs, 15-liquid falling folded plates, 16-liquid receiving plates, 17-liquid sealing baffles, 18-inside downcomers, 19-catalyst supporting plates and 191-grids;

21-a liquid phase feeding pipe, 22-a liquid phase distributing pipe, 220-a liquid phase distributing pipe body and 221-a liquid phase pore canal;

31-gas phase feeding pipe, 32-gas phase distributing pipe, 320-gas phase distributing pipe body, 321-gas phase channel, 33-gas phase distributing disk, 331-hole.

Detailed Description

The following detailed description of embodiments of the invention is, therefore, to be taken in conjunction with the accompanying drawings, and it is to be understood that the scope of the invention is not limited to the specific embodiments.

FIG. 1 is a schematic diagram of the process flow for producing low aromatic hydrocarbon solvent oil according to the present invention. As shown in fig. 1 and 2, the biological grease 2, the fresh hydrogen 3 and the circulating pressurized hydrogen-rich gas 4 are mixed and then enter a hydrodeoxygenation reactor 5, and a hydrogenation protecting agent and a hydrodeoxygenation catalyst are sequentially placed in the hydrodeoxygenation reactor from top to bottom. The Fischer-Tropsch oil 255, the fresh hydrogen 3 and the circulating pressurized hydrogen-rich gas 4 are mixed and then enter a hydrofining reactor 55, and a hydrogenation protecting agent and a hydrofining catalyst are sequentially placed in the hydrofining reactor from top to bottom. The hydrodeoxygenation stream 6 and the hydrofinishing stream are dehydrated and then enter a liquid phase feed pipe 21 of the hydrogenation reaction distillation column 1, are uniformly distributed through a liquid phase distribution pipe 22 and then flow downwards along a packing or catalyst bed. When the dehydrated hydrodeoxygenation material flow 6 passes through the packing of the separation zone, light components flow upwards through the gas phase channel 321, heavy components continue to flow downwards, and when entering the catalyst bed of the hydrogenation reaction zone, the heavy components react with hydrogen entering the catalyst bed upwards from the gas phase feeding pipe 31, gas phase products directly enter the gas phase channel 321, and the heavy components continue to flow downwards along the catalyst bed and react.

The liquid phase material flow 7 generated by the hydrocracking reaction flows out from the bottom of the hydrogenation reaction distillation column, the gas phase material flow 8 from the gas phase channel 321 flows out from the top of the hydrogenation reaction distillation column, the gas phase material flow 8 and the gas phase material flow are mixed and enter the high-pressure separator 9 to carry out gas-liquid separation, the separated high-pressure hydrogen-rich gas 4 is mixed with the new hydrogen 3 to be used as circulating hydrogen, the separated liquid enters the fractionating column 10 to be fractionated to obtain gas 101, naphtha 102, low aromatic hydrocarbon solvent oil 103 and unconverted oil 104, and the unconverted oil 104 is wholly or partially circulated back to the hydrogenation reaction distillation column 1 to continue the hydrocracking reaction, and the device can also be led out to be used as a raw material for producing lubricating oil base oil.

As shown in FIG. 2, the internals of the hydrogenation reaction distillation column 1 of the present invention include a catalyst bed, a liquid phase feed subunit, a gas phase feed subunit, and gas phase channels. The upper part of the hydrogenation reaction distillation tower is provided with a separation zone, and the lower part is provided with a hydrogenation reaction zone. Wherein the catalyst bed of the separation zone is used for placing the packing 110, and the catalyst bed of the hydrogenation reaction zone is used for packing the solid catalyst 111. The upper part of each catalyst bed is provided with an inclined baffle plate, the whole shape formed by the inclined baffle plates can be umbrella-shaped, and the inclined baffle plate plays a role of a baffle plate, on one hand, gas phase feeding and product gas between adjacent bed layers can be separated, and on the other hand, the baffle plate plays a role of guiding liquid phase and gas phase, and preferably, but not limited, the umbrella cover can be arc-shaped or folded umbrella-shaped. The liquid phase feed sub-unit is disposed above the inclined partition plate (i.e., umbrella-shaped partition plate 11) of the topmost catalyst bed, and the liquid phase feed passing through the umbrella-shaped partition plate 11 is guided to the catalyst bed to be in contact with the packing 110, specifically, the end of the umbrella-shaped partition plate 11 is provided with an annular outer downcomer 12, the bottom of which is spaced apart from the bottom of the catalyst bed by a distance such that the liquid phase feed enters the catalyst bed in the radial direction of the reactive distillation column 1.

Each catalyst bed in the hydrogenation reaction zone is provided with a gas phase feeding subunit, and the gas phase feeding subunit is specifically arranged between the catalyst bed in the upper layer and the umbrella-shaped baffle 11 in the lower layer, and the gas phase feeding of each layer upwards enters the catalyst bed. After the gas-liquid phase feed and the solid catalyst fully react in the catalyst bed, the gas phase product of each layer is guided to the gas phase channel 13 along the lower part of the umbrella-shaped baffle 11. The gas phase channel 13 is in a relatively isolated state from the gas phase feeding subunit, i.e. the gas phase product generated after the gas phase feeding and the liquid phase feeding react in the catalyst bed directly enters the gas phase channel 13. Preferably, but not by way of limitation, the gas phase channels are located in the middle of the reactive distillation column 1 and run through all catalyst beds from bottom to top.

As shown in fig. 2 and 3, the liquid phase feed subunit further comprises a liquid phase feed pipe 21 and a liquid phase distribution pipe 22. The liquid phase feeding pipe 21 extends along the radial direction of the hydrogenation reaction unit, the liquid phase distributing pipe 22 is annular, the liquid phase feeding pipe 21 is orthogonal or tangentially intersected with the pipe body 220 of the liquid phase distributing pipe 22, and the pipe wall of the liquid phase distributing pipe 22 is provided with a plurality of liquid phase pore canals 221 for uniformly distributing liquid phase feeding materials to all directions of the annular outer side downcomer 12. The openings of the liquid phase channels 221 may be in all directions on the upper, lower and side faces of the tube. Liquid phase feed enters the reactive distillation column 1 through a liquid phase feed pipe 21, is distributed into the column through an annular liquid phase distribution pipe 22, flows into an outer downcomer 12 from the periphery through an umbrella-shaped baffle plate 11, and transversely enters a catalyst bed layer to be contacted with packing and solid catalyst after passing through the outer downcomer 12. The feeding direction of the liquid phase feeding pipe 21 is the radial direction of the reactive distillation tower, and is intersected with the radial orthogonal or tangential direction of the annular liquid phase distributing pipe 22, the annular diameter of the annular liquid phase distributing pipe 22 is larger than the outer diameter of the gas phase channel 13 and smaller than the inner diameter of the reactive distillation tower 1, and a plurality of pore canals on the pipe wall of the annular liquid phase distributing pipe 22 are convenient for the liquid phase feeding to be uniformly distributed in all directions of the outer side downcomer 12. The height of the liquid-lowering folded plate 15 is generally smaller than the filling height of the filling material or the catalyst of the layer, and the distance between the liquid-lowering folded plate 15 and the inner wall of the reactive distillation column 1 is determined according to the flow rate of the liquid-phase reactant of the layer.

As further shown in FIG. 2, the height of each catalyst bed in the reactive distillation column 1 can be the same or different, and the upper part of the catalyst bed is fixed by a screen according to different chemical reaction systems, so that the bed is relatively stable, and the height of the bed is set to be 10mm to 1000mm. The catalyst bed is provided with an overflow weir 14 and a liquid seal baffle 17, and the overflow weir 14 is arranged at one side close to the gas phase channel 13. A liquid seal 17 is provided above weir 14 to isolate the gas phase feed from the gas phase product. Further, the liquid seal baffle 17 comprises a horizontal part and a vertical part, the horizontal part is in a ring-shaped flat plate shape and is positioned above the overflow weir 14, the vertical part is in a cylinder shape, the vertical part and the horizontal part are integrally formed, other seamless connection modes can be adopted, and the lower end of the vertical part is separated from the bottom of the catalyst bed by a certain distance, so that the outflow of liquid phase products can be ensured. Unreacted liquid feed and reacted but liquid-phase-maintaining material within the catalyst bed passes over weir 14, through inner downcomer 18 (i.e., the annular space between weir 14 and the outer wall of gas phase channel 13), along umbrella baffle 11, through outer downcomer 12 of the next layer into the next catalyst bed. The height of weir 14 is above the upper level of the catalyst in the bed, preferably 10 to 100mm. The size of the space between the overflow weir 14 and the annular inner downcomer 18 formed by the outer wall of the gas phase channel 13 depends on the size of the liquid phase load, and the size of each bed downcomer can be the same or different.

As further shown in fig. 2, 4 to 6, the gas phase feed subunit comprises a gas phase feed pipe 31 and a gas phase distribution pipe 32, the gas phase feed pipe 31 extending in the radial direction of the reactive distillation column 1. The gas distribution pipe 32 is annular (see fig. 4 and 5) or multi-layer concentric annular (see two-layer concentric rings of fig. 6), the gas distribution pipe 31 is orthogonal to the gas distribution pipe body 320 of the gas distribution pipe 32 (see fig. 4) or tangentially intersects with the gas distribution pipe body (see fig. 5), and a plurality of gas phase channels 321 are arranged on the wall surface of the gas distribution pipe 32 and are used for uniformly distributing gas phase feed materials to all directions at the bottom of the catalyst bed. Preferably, and not by way of limitation, the gas distribution tube 32 may be disposed below the catalyst bed or within the catalyst bed. As further shown in fig. 7, the gas phase feed subunit further includes a gas phase distribution plate 33, where the gas phase distribution plate 33 is located at the bottom of the catalyst bed and has a plate shape as a whole, and a plurality of holes 331 are uniformly distributed on the gas phase distribution plate. The gas phase feed enters the reactive distillation column 1 through a gas phase feed pipe 31 of each layer, is distributed into the reactive distillation column 1 through an annular gas phase distribution pipe 32, and enters the catalyst bed upward through a gas phase distribution plate 33 at the lower part of the catalyst support plate 19. The gas phase feeding pipe 31 enters the reactive distillation column 1 in a radial direction and is orthogonal or tangentially intersected with the annular gas phase distributing pipe 32, the annular gas phase distributing pipe 32 is positioned below the catalyst bed layer, the annular diameter of the annular gas phase distributing pipe 32 is smaller than the outer annular diameter of the catalyst bed layer, the inner diameter of the annular gas phase distributing pipe is larger than the inner annular diameter of the catalyst bed layer, and a plurality of gas phase channels 321 on the pipe wall of the annular gas phase distributing pipe 32 are convenient for gas to be uniformly distributed at all positions of the gas phase distributing plate 33. The function of the catalyst support plate 19 is mainly to support the catalyst bed, ensuring that the catalyst bed remains stable in the axial direction of the reactive distillation column. The purpose of the gas phase distributor plate 33 is to ensure uniform distribution of the gas phase feed while avoiding direct leakage of the liquid phase feed over the catalyst bed as much as possible (with the gas phase distributor plate 33 of the present invention, the liquid leakage is < 15%). On the same plane, when more than one concentric annular gas distribution tube 32 of different diameters is provided, the distribution of the gas phase feed may be made more uniform. The embodiment of fig. 2 provides for the annular gas distribution tube 32 to be positioned below the catalyst bed, and when the annular gas distribution tube 32 is installed within the catalyst bed, the catalyst support plate 19 may be modified from the grid 191 of fig. 8 to a support plate, while eliminating the gas distribution plate 33.

In the hydrogenation reaction unit, liquid-phase feeding and gas-phase feeding are subjected to hydrogenation reaction in the catalyst bed, gas-phase products and unreacted gas-phase feeding are lifted to be separated from a reaction system through the gas-phase channel 13, and gas-phase products generated after chemical reaction of reactants in the catalyst bed can timely leave the reaction zone and cannot enter the upper catalyst bed again (isolated by the umbrella-shaped partition plate), so that secondary reaction of target products is avoided, and the reaction selectivity is improved. Meanwhile, due to the fact that products in the reaction zone leave, the reaction driving force is increased, and the equilibrium conversion rate is improved.

The hydrogenation reaction distillation tower of the invention can be a multi-layer plate tower structure. The number of the catalyst beds in the hydrogenation reaction distillation tower is two or more than two. The hydrogenation reaction distillation column 1 of the present invention is suitable for a reaction system in which at least one liquid-phase feed and at least one gas-phase feed are chemically reacted on a solid catalyst and at least one gas-phase product is present in the reaction product. Such as hydrocracking of petroleum fractions and chemical synthesis oils, hydrodewaxing of diesel and lubricant oil fractions, hydrotreating of various petroleum fractions, and the like.

In the hydrogenation reaction distillation tower 1, each layer of tray comprises a downcomer, an overflow weir and a liquid receiving tray 16, a liquid sealing baffle plate is arranged on the tray, the liquid sealing baffle plate is connected with a gas phase channel, adjacent trays are separated by an umbrella-shaped baffle plate, each layer of tray is of an annular structure, the annular inner edge is connected with the gas phase channel, and the outer edge is connected with the inner wall of the hydrogenation reaction distillation tower. The gas phase channels are common channels for removing gas phase products generated by chemical reaction on each tray. The liquid feed locations of the embodiments of the present invention are all above one tray, or may have liquid feed on some trays or each tray, and gas phase feed has feed at the lower portion of each tray. The catalyst filling area is arranged above each layer of tray, the liquid phase feed radially flows through the catalyst bed, the gas phase feed enters from below the tray and reacts under the action of the catalyst, gas phase materials generated after the reaction are directly separated from the reaction system and enter a gas phase channel in the middle, and the liquid phase enters the next bed through a downcomer after leaving the bed. Because the reaction and the separation are carried out simultaneously, the reaction balance can be destroyed, and the conversion rate of reactants and the selectivity of target products can be effectively improved.

Example 1

By adopting the device and the flow, the biological grease and the Fischer-Tropsch oil are hydrotreated. The biological grease raw oil is soybean oil, and the Fischer-Tropsch oil is diesel oil part in the Fischer-Tropsch full fraction.

The upper part in the hydrodeoxygenation reactor is provided with a hydrogenation protecting agent, the lower part is filled with a hydrodeoxygenation catalyst, and the volume ratio of the hydrogenation protecting agent to the hydrodeoxygenation catalyst is 1:10. The biomass raw material is subjected to olefin hydrogenation saturation and hydrodeoxygenation reaction in a hydrodeoxygenation reactor. The hydrogenation protecting agent is FZC-105/FZC-106, and the hydrodeoxygenation catalyst is FZC-41. All the above catalysts are produced by China petrochemical industry institute of great company petrochemical industry. The specific operating process conditions are shown in Table 2.

The upper part in the hydrofining reactor is provided with a hydrogenation protecting agent, the lower part is filled with a hydrofining catalyst, and the volume ratio of the hydrogenation protecting agent to the hydrofining catalyst is 1:10. The Fischer-Tropsch oil raw material is subjected to olefin hydrogenation saturation and hydrodeoxygenation reaction in a hydrofining reactor. The hydrogenation protecting agent is FZC-105/FZC-106, and the hydrofining catalyst is FF-66. All the above catalysts are produced by China petrochemical industry institute of great company petrochemical industry. The specific operating process conditions are shown in Table 2.

Five layers of tower plates are arranged in the hydrogenation reaction distillation tower, the upper part of the first layer of tower plates is filled with filler, the lower part is filled with hydrocracking catalyst, and the rest four layers of tower plates are filled with the hydrocracking catalyst. The packing is 5mm corundum Raschig ring, the hydrocracking catalyst is FC-46 produced by China petrochemical industry institute of great company petrochemical industry, and the specific operation process conditions are shown in Table 2.

The light fraction of the hydrogenation reaction distillation tower is recycled by a high-pressure separator to obtain high-pressure hydrogen-rich gas 4, the liquid enters a fractionating tower for fractionation to obtain gas, naphtha and low aromatic solvent oil, and the heavy fraction at the bottom of the tower is directly recycled to the hydrogenation reaction distillation tower.

The product distribution and properties are shown in Table 3.

Examples 2 to 3

The reaction conditions were changed as shown in Table 2, and the other conditions were the same as in example 1. The product distribution and properties are shown in Table 3.

Example 4

The biomass feedstock was replaced with palm oil, see in particular table 2. The product distribution and properties are shown in Table 3.

TABLE 1 Main Properties of raw oil

Table 2 example process conditions

TABLE 3 product distribution and Properties

From the results in Table 3, animal and vegetable oils and Fischer-Tropsch oil can be hydrogenated to produce low aromatic hydrocarbon solvent oil with very low aromatic hydrocarbon content, which is a good quality chemical product.

Claims

1. A hydrogenation device for bio-oil and Fischer-Tropsch synthetic oil, comprising:

(1) a hydrodeoxygenation reactor, which is used to perform a hydrodeoxygenation reaction on the raw bio-oil;

(2) a hydrotreating reactor, which is used to perform a hydrotreating reaction on the Fischer-Tropsch synthetic oil;

(3) a hydrogenation reaction distillation tower, which receives a liquid feed, wherein the liquid feed is a liquid product of bio-oil hydrodeoxygenation and a Fischer-Tropsch synthetic oil hydrorefining product, and the liquid feed flows from top to bottom through each catalyst bed in the upper part of the hydrogenation reaction distillation tower; the preheated hydrogen enters each catalyst bed from the bottom to the top evenly from the bottom of each catalyst bed; wherein the liquid feed and hydrogen undergo a hydrocracking reaction in the hydrocracking catalyst bed loaded in the hydrogenation reaction distillation tower; the light components produced by the hydrocracking reaction and the gaseous feed that does not participate in the reaction are discharged from the top of the hydrogenation reaction distillation tower through an independent gas phase channel, and the liquid product obtained by the hydrocracking is discharged from the bottom of the hydrogenation reaction distillation tower;

(4) A separation unit, which is used to separate the light components produced by hydrocracking and the gaseous feed that does not participate in the reaction, and obtain hydrogen-rich gas, dry gas, liquefied gas, and low-aromatic solvent oil;

The hydrogenation reaction distillation tower comprises at least one hydrogenation reaction zone, the hydrogenation reaction zone is a 2-n-layer tower plate structure, each tower plate is provided with a hydrocracking catalyst bed, wherein n is an integer greater than 2; an inclined partition is provided above each catalyst bed, and each catalyst bed and inclined partition are provided with a gas phase channel;

The gas phase channel is located in the middle of the hydrogenation reaction distillation tower and passes through all catalyst beds from bottom to top; the lower end of the gas phase channel passing through the upper catalyst bed is connected to the inclined partition plate on the next catalyst bed;

The hydrogenation reaction distillation tower further comprises a liquid phase feed subunit, which is arranged above the inclined partition of the topmost catalyst bed layer, through which the liquid phase feed is guided to the catalyst bed layer; the liquid phase feed subunit comprises: a liquid phase feed pipe, which extends in the radial direction of the catalyst bed layer;

The hydrogenation reaction distillation tower also includes a gas phase feed subunit, which is arranged between the catalyst bed layer of the upper layer and the inclined partition plate of the lower layer;

Each catalyst bed is provided with: an overflow weir, which is arranged on the side close to the gas phase channel; and a liquid sealing baffle, which is arranged on the upper part of the overflow weir and is used to isolate the gas phase feed from the gas phase product.

2. The hydrogenation device according to claim 1 is characterized in that it also includes (5) a gas circulation unit, which is used to selectively concentrate the hydrogen-rich gas obtained by the separation unit and circulate it back to the hydrodeoxygenation reactor, the hydrotreating reactor and the hydrogenation reaction distillation tower.

3. The hydrogenation device according to claim 1 is characterized in that it also includes (6) a liquid circulation unit, which is used to recycle at least a portion of the liquid product obtained from the hydrogenation reaction distillation tower back to the hydrogenation reaction distillation tower.

The hydrogenation device according to claim 1 , wherein the inclined partition is an umbrella-shaped partition.

5. The device according to claim 1 is characterized in that an annular downcomer is provided at the end of the inclined partition, and the bottom of the annular downcomer is spaced a distance from the bottom of the catalyst bed.

6. The device according to claim 1 is characterized in that the liquid-sealed baffle comprises: a horizontal portion, which is in the shape of an annular flat plate and is located above the overflow weir; a vertical portion, which is in the shape of a cylinder, the vertical portion is integrally formed with the horizontal portion, and the lower end of the vertical portion is spaced a distance from the bottom of the catalyst bed.

7. The device according to claim 1 or 5 is characterized in that the liquid-phase feed subunit also includes: a liquid-phase distribution pipe, which is annular and intersects orthogonally or tangentially with the liquid-phase feed pipe, and the pipe wall of the liquid-phase distribution pipe is provided with a plurality of channels for evenly distributing the liquid-phase feed to all directions of the annular downcomer.

8. The device according to claim 1 is characterized in that the gas-phase feed subunit comprises: a gas-phase feed pipe, which extends in the radial direction of the catalyst bed; a gas-phase distribution pipe, which is annular or multi-layer concentric annular, and the gas-phase distribution pipe intersects the gas-phase feed pipe orthogonally or tangentially, and a plurality of holes are provided on the wall of the gas-phase distribution pipe for evenly distributing the gas-phase feed to all directions at the bottom of the catalyst bed.

9. The device according to claim 8, characterized in that the gas-phase feeding subunit further comprises: a gas-phase distribution plate, which is located at the bottom of the catalyst bed and is disk-shaped as a whole, and a plurality of holes are evenly and densely distributed on the gas-phase distribution plate.

10. A method for producing low aromatic solvent oil by hydrogenating bio-oil and Fischer-Tropsch synthetic oil, using the hydrogenation device according to claim 1, characterized in that it comprises the following steps:

(1) The bio-oil raw material is mixed with hydrogen and then enters the hydrodeoxygenation reactor, where it undergoes a hydrogenation reaction through a hydrodeoxygenation catalyst bed under hydrodeoxygenation conditions;

(2) The Fischer-Tropsch synthetic oil feedstock is mixed with hydrogen and then enters the hydrotreating reactor, where it undergoes a hydrogenation reaction through a hydrotreating catalyst bed under hydrotreating conditions;

(3) The liquid phase product of the bio-oil hydrodeoxygenation obtained in step (1) and the Fischer-Tropsch synthetic oil hydrorefining product obtained in step (2) are mixed and dehydrated to obtain a liquid product, the liquid product enters the liquid phase feed pipe of the hydrogenation reaction distillation tower, descends in the tower and sequentially enters multiple catalyst beds, and contacts with the hydrogen entering the catalyst beds upward to perform a hydrocracking reaction; the gas phase product obtained by the reaction of each catalyst bed enters the gas phase channel, and flows upward out of the hydrogenation reaction distillation tower in the gas phase channel; the liquid phase product obtained by the reaction of each catalyst bed sequentially enters the next catalyst bed to react, and the final liquid product flows downward out of the hydrogenation reaction distillation tower;

(4) The gas phase product obtained in step (3) is separated in a separation unit to obtain hydrogen-rich gas, dry gas, liquefied gas and low aromatic solvent oil.

11. The method according to claim 10, characterized in that it also includes step (5): the hydrogen-rich gas obtained in step (4) is selectively concentrated and then returned to step (1), step (2) and step (3).

12. The method according to claim 10, characterized in that it also includes step (6): at least a portion of the liquid product obtained in step (3) is recycled back to step (3).

13. The method according to claim 10, characterized in that the biological oil raw material includes at least one of soybean oil, peanut oil, castor oil, rapeseed oil, corn oil, olive oil, palm oil, coconut oil, tung oil, linseed oil, sesame oil, cottonseed oil, sunflower oil, rice bran oil, tallow, lard, mutton fat, fish oil, acidified oil, and waste cooking oil.

14. The method according to claim 10, characterized in that the hydrodeoxygenation conditions are: reaction pressure 3.0MPa-20.0MPa, hydrogen-oil volume ratio 200:1-3000:1, volume space velocity 0.1h ^-1 -6.0h ^-1 , average reaction temperature 180°C-450°C.

15. The method according to claim 10, characterized in that the final distillation point of the Fischer-Tropsch synthetic oil is not higher than 750°C.

16. The method according to claim 10, characterized in that the hydrotreating conditions in step (2) are as follows: average reaction temperature of 150°C to 450°C, system reaction pressure of 3.0MPa to 20.0MPa, hydrogen to oil volume ratio of 100 to 1000, and liquid hourly volume space velocity of 0.4h ^-1 to 10.0h ^-1 .

17. The method according to claim 10, characterized in that in step (1), a hydrogenation protective agent is loaded upstream of the hydrodeoxygenation catalyst, and the volume ratio of the hydrogenation protective agent to the hydrodeoxygenation agent is 15:85-70:30.

18. The method according to claim 10, characterized in that in step (3), a solid filler is placed on the upper part of each catalyst bed, and a hydrocracking catalyst is loaded on the lower part.

19. The method according to claim 10, characterized in that the operating conditions of the hydrogenation reaction distillation tower in step (3) are: reaction temperature of 260°C to 450°C, reaction pressure of 3.0MPa to 20.0MPa, hydrogen to oil volume ratio of 100 to 2000, and liquid hourly volume space velocity of 0.1h ^-1 to 10.0h ^-1 .