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

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

Technical Field

[0001] The invention belongs to the field of oil refining technology and relates to a hydrogenation distillation method, in particular to a device and a hydrogenation method for producing low aromatic solvent oil from bio-oil and Fischer-Tropsch synthetic oil by hydrogenation technology. Background Art

[0002] With the implementation of carbon neutrality-related policies around the world, the status of green and circular economy has become more prominent. Bio-oil is a renewable resource and one of the key development directions, which has received widespread attention from the world. Bio-oil can be deoxygenated by hydrogenation to obtain hydrocarbon products, such as the production of motor fuels such as jet fuel and diesel, that is, bio-oil is hydrodeoxygenated (completely or partially removed) to produce products that meet motor fuel standards. This method can directly meet the requirements of the existing market; for example, the production of low-carbon alkanes, that is, bio-oil is hydrodeoxygenated and then hydrocracking is performed to break large molecules into small molecules to produce low-carbon alkanes that meet product requirements.

[0003] The Fischer-Tropsch synthesis reaction is a process in which hydrogen and carbon monoxide are used as raw materials. The Fischer-Tropsch synthesis reaction occurs in the presence of a catalyst to produce liquid fuels. The raw materials are mainly coal and natural gas. Fischer-Tropsch oil is mainly composed of normal alkanes, olefins and a certain amount of oxygen-containing compounds, and its content of non-ideal components such as sulfur, nitrogen, and aromatics is extremely low. The Fischer-Tropsch oil fraction range is relatively wide, and the different boiling range fractions obtained by distillation basically do not contain sulfur and nitrogen, and have the highest olefin content and oxygen content. Therefore, the Fischer-Tropsch oil must undergo corresponding hydrogenation treatment to obtain petrochemical raw materials or transportation fuels that meet the use specifications.

[0004] Existing vegetable oil hydrogenation technology, US5705722 discloses a diesel blending component containing unsaturated fatty acids, fats, etc., and animal oils, which are mixed and hydrogenated to produce a diesel fraction range. CN101462915 discloses a C6-C 12 The preparation method of alkanes mainly uses animal and vegetable fats as raw materials to directly perform hydrocracking to produce alkanes. Because no pretreatment is performed, the generated water has a very adverse effect on the molecular sieve of the hydrocracking catalyst, and the operation cycle is significantly shortened. EP1741767 and EP1741768 disclose a method for producing diesel fractions from animal and vegetable fats. The animal and vegetable fats are first subjected to hydrotreatment and then passed through an isomerization catalyst bed to obtain a low-condensation point diesel component. However, since water is generated during the hydrotreatment process, it has a very adverse effect on the isomerization catalyst, and the device cannot be stably operated for a long period of time.

[0005] The existing technology of hydrogenating vegetable oil to produce solvent oil, CN103102922, CN103102898, CN103102899, CN103102900, CN103102901, CN103102902, CN103102903, CN103102907, CN103102908, CN103102911, CN103102918, CN103102919, CN103102920 disclose methods for producing low aromatic solvent oil from bio-oil, these methods are all under hydrogenation conditions under which the bio-oil raw material is first subjected to a hydrodeoxygenation reaction, and then the resulting oil is dehydrated and then hydrocracking is carried out under hydrocracking conditions to crack large molecules into small molecules to produce low aromatic solvent oil.

[0006] The general process flow of hydrotreating Fischer-Tropsch oil is that after the Fischer-Tropsch oil (including naphtha fraction, diesel fraction and heavy oil fraction) is mixed with hydrogen, it first enters the hydrotreating reactor, and then enters the fractionation system after olefin saturation and oxygen-containing compound hydrodeoxygenation reaction on the catalyst bed. In the fractionation tower, the hydrotreating product is separated into naphtha, diesel and heavy oil by distillation. The heavy oil continues to enter the subsequent hydrocracking reactor, and the cracking product enters the second fractionation system. The tail oil is partially or fully recycled back to the hydrocracking reactor, with the purpose of obtaining more naphtha and diesel fractions through cracking reactions. These processes all use the traditional fixed-bed hydrogenation mode, and completely separate the hydrotreating and hydrocracking parts, with two fractionation systems in total, which are complex processes, high equipment investment and high energy consumption.

[0007] The heavy oil hydrocracking reaction in the current Fischer-Tropsch synthesis oil still adopts the traditional fixed-bed hydrocracking mode, where the feedstock oil and hydrogen flow into the reactor together, and the materials react and flow downward during the bed advancement process. During the entire reaction process, the gas phase reaction products cannot leave the reactor in time, so they will occupy the catalyst pores and cracking reaction will occur again, resulting in reduced reaction selectivity and yield of the target product.

[0008] The current common process flow of vegetable oil hydrogenation is usually that after the vegetable oil is mixed with hydrogen, it first enters the hydrodeoxygenation reactor, and then undergoes olefin hydrogenation saturation and oxygen-containing compound hydrodeoxygenation reaction on the catalyst bed, followed by separation and dehydration. The dehydrated fraction enters the subsequent hydrocracking reactor, and the cracked product enters the fractionation system to obtain low-carbon alkanes. The unconverted tail oil can be partially or fully circulated back to the hydrocracking reactor, with the purpose of obtaining more low-aromatic solvent oil through cracking reaction. These processes all adopt the traditional fixed bed hydrogenation mode, especially multiple catalyst beds are set in the hydrocracking reactor, which will cause adverse effects, mainly because the target product obtained by the initial contact with the hydrocracking catalyst will continue to enter the subsequent hydrocracking catalyst bed, so that the generated target product will continue to undergo hydrocracking reaction to obtain smaller molecular alkanes. That is, when the same conversion rate is achieved, the selectivity of the target product is poor, or the yield of the target product is low when the raw materials are fully converted. Summary of the invention

[0009] In view of the shortcomings of the prior art, the present invention aims to provide a device and a hydrogenation method for producing low aromatic solvent oil by hydrocracking of bio-oil and Fischer-Tropsch synthetic oil. The method improves the selectivity of the low aromatic solvent oil target product by using a hydrodistillation reactor to continuously remove the generated target product from the reaction zone, and can reduce the chemical hydrogen consumption when achieving the same target product yield.

[0010] The first aspect of the present invention is to provide a hydrogenation device for producing low aromatic solvent oil from bio-oil and Fischer-Tropsch synthetic oil.

[0011] The hydrogenation device of the present invention comprises:

[0012] (1) a hydrodeoxygenation reactor, which is used to perform a hydrodeoxygenation reaction on the raw bio-oil;

[0013] (2) a hydrotreating reactor, which is used to perform a hydrotreating reaction on the Fischer-Tropsch synthetic oil;

[0014] (3) a hydrogenation reaction distillation tower, which receives the liquid phase product after hydrodeoxygenation of bio-oil and the hydrorefining product of Fischer-Tropsch synthetic oil as liquid phase feed, and the raw material feed flows from top to bottom through each catalyst bed in the upper part of the 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 phase feed and hydrogen undergo a hydrocracking reaction in the hydrocracking catalyst bed loaded in the distillation tower; the light components produced by the hydrocracking reaction and the gas phase 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 hydrocracking is discharged from the bottom of the hydrogenation reaction distillation tower;

[0015] (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.

[0016] Furthermore, the above technical solution 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 hydrofining reactor and the hydrogenation reaction distillation tower.

[0017] Furthermore, the above technical solution 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.

[0018] Furthermore, in the above technical solution, the hydrogenation reaction distillation tower comprises at least one hydrogenation reaction zone, and the hydrogenation reaction zone is a 2 to n-layer tower plate structure, each tower plate is filled with a hydrocracking catalyst bed, wherein n is an integer greater than 2; an inclined partition is arranged above each catalyst bed, preferably an umbrella-shaped partition; each catalyst bed and the inclined partition are provided with a gas phase channel.

[0019] Furthermore, in the above technical solution, the hydrodeoxygenation reaction carried out in the hydrodeoxygenation reactor includes a hydrogenation saturation reaction of olefins contained in the bio-oil feedstock, a hydrodeoxygenation reaction of carboxyl groups and carbonyl groups contained therein, and the like.

[0020] Furthermore, in the above technical solution, the hydrotreating reaction carried out in the hydrotreating reactor includes a hydrogenation saturation reaction of olefins contained in the Fischer-Tropsch synthetic oil feedstock, a hydrodeoxygenation reaction of hydroxyl carboxyl groups contained therein, and a hydrodecontamination reaction.

[0021] Furthermore, in the above technical solution, the hydrogenation reaction distillation tower further comprises a liquid-phase feed subunit, which is arranged above the inclined partition of the topmost catalyst bed, and the liquid-phase feed is guided to the catalyst bed through the inclined partition.

[0022] Furthermore, in the above technical solution, 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, so that the liquid feed enters the catalyst bed in a radial direction.

[0023] Furthermore, in the above technical solution, the hydrogenation reaction distillation tower further comprises 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. The gas phase feed enters the catalyst bed layer of the upper layer upward.

[0024] Furthermore, in the above technical solution, 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. This design makes the gas phase channel and the gas phase feed subunit relatively isolated, and the gas phase product generated after the gas phase feed and the liquid phase feed react in the catalyst bed directly enters the gas phase channel.

[0025] Furthermore, in the above technical solution, the 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 to isolate the gas phase feed from the gas phase product.

[0026] Furthermore, in the above technical solution, the liquid seal baffle includes: a horizontal portion, which is an annular flat plate and is located above the overflow weir; a vertical portion, which is cylindrical and 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.

[0027] Furthermore, in the above technical solution, the liquid-phase feed subunit further includes: a liquid-phase feed pipe, which extends along the radial direction of the catalyst bed; a liquid-phase distribution pipe, which is annular and intersects orthogonally or tangentially with the liquid-phase feed pipe, and the tube 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.

[0028] Furthermore, in the above technical solution, the gas-phase feed subunit further includes: a gas-phase feed pipe, which extends along the radial direction of the catalyst bed; a gas-phase distribution pipe, which is in a ring shape or a multi-layer concentric ring shape, 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 surface of the gas-phase distribution pipe for evenly distributing the gas-phase feed to all directions at the bottom of the catalyst bed.

[0029] Furthermore, in the above technical solution, the gas phase feeding subunit may also include: 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.

[0030] Furthermore, in the above technical solution, the gas phase channel is located in the middle of the hydrogenation reaction unit and passes through all catalyst beds from bottom to top.

[0031] In the above technical solution, the uppermost end of the gas phase channel exits the hydrogenation reactor, and the lowermost end is close to the bottom of the hydrogenation reactor. The port needs to be immersed in the liquid phase product, and a liquid level monitoring unit is provided at the bottom of the hydrogenation reactor.

[0032] Furthermore, in the above technical solution, the height of the filler or catalyst bed is set to 10 mm to 1000 mm according to different reaction systems.

[0033] Furthermore, in the above technical solution, the upper edge of the overflow weir may be 10 to 100 mm higher than the upper surface of the catalyst bed.

[0034] Furthermore, in the above technical solution, the gas phase distribution pipe can be arranged below the catalyst bed or inside the catalyst bed.

[0035] Furthermore, in the above technical solution, a gas-liquid separation device is required to be installed between the hydrodeoxygenation reactor and the hydrorefining reactor and the hydrorefining reactor to remove the water generated by the hydrodeoxygenation reaction. Alternatively, a gas-liquid separation device can be installed to combine the hydrodeoxygenation stream and the hydrorefining stream and then enter the separation device for dehydration.

[0036] The second aspect of the present invention is to provide a method for producing low aromatic solvent oil by hydrogenating bio-oil and Fischer-Tropsch synthetic oil.

[0037] The method comprises the following steps:

[0038] (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;

[0039] (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;

[0040] (3) The mixed liquid product of the stream obtained in step (1) and the hydrotreated stream obtained in step (2) is dehydrated and enters a hydrogenation reaction distillation tower, and then descends in the tower and enters a plurality of catalyst beds in turn, and contacts with the hydrogen entering the catalyst beds upward to carry out a hydrocracking reaction; the gaseous product obtained by the reaction of each catalyst bed enters a gas phase channel, and flows upward out of the hydrogenation reaction distillation tower in the gas phase channel; the liquid product obtained by the reaction of each catalyst bed enters the next catalyst bed in turn to react, and the liquid product finally flows downward out of the hydrogenation reaction distillation tower;

[0041] (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.

[0042] Furthermore, the above technical solution also includes step (5): the hydrogen-rich gas obtained in step (4) is returned to step (1) and / or step (2) after optional concentration.

[0043] Furthermore, the above technical solution also includes step (6): the liquid product obtained in step (4) is recycled back to step (1) and / or step (2).

[0044] Furthermore, in the above technical solution, the bio-oil raw material in step (1) is at least one of vegetable oil and animal oil. Vegetable oil includes 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 oil and rice bran oil. Animal oil includes one or more of beef tallow, lard, sheep oil and fish oil. The bio-oil raw material can also be selected from acidified oil, waste cooking oil and the like.

[0045] Furthermore, in the above technical solution, the hydrodeoxygenation reactor described in step (1) and the hydrofining reactor described in step (2) are both conventional fixed bed reactors. The operating conditions of the hydrogenation pretreatment are generally: 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℃~450℃; preferred operating conditions are reaction pressure 3.0MPa~18.0MPa, hydrogen oil volume ratio 300:1~2500:1, volume space velocity 0.2h -1 ~4.0h -1 , the average reaction temperature is 200℃~440℃.

[0046] Furthermore, in the above technical solution, the reaction material first passes through the hydrogenation protective agent. The protective agent is a conventional protective agent. Only one can be selected, or two or more protective agents can be selected according to the order of contact with the raw oil, according to the particle size from large to small, the content of hydrogenation metal from small to large, and the hydrogenation performance from weak to strong. Multiple protective agents form a protective agent system. For example, the three protective agent systems FZC-100 / FZC-105 / FZC-106 (volume ratio of 10 / 35 / 55) developed by FRIPP can be used here. Usually, the total volume air velocity of fresh raw materials of the protective agent is 3.0h -1 ~30.0h -1 , preferably 4.0h -1 ~20.0h -1 .

[0047] Further, in the above technical scheme, the logistics through the hydrogenation protection agent continues to pass through the hydrodeoxygenation catalyst, and one hydrodeoxygenation catalyst can be used, preferably two or more. The hydrogenation active components of the hydrodeoxygenation catalyst are usually Mo, W, Ni, Co, etc. of group VIB and group VIII, and the weight content of the oxide is 3% to 20%, preferably 3% to 15%, and more preferably 3% to 10%. If two or more hydrodeoxygenation catalysts are used, the first hydrogenation catalyst through which the reaction material first passes accounts for 10% to 80% of the volume of all hydrodeoxygenation catalysts, preferably 20% to 70%, and preferably 30% to 60%. In two adjacent hydrodeoxygenation catalyst beds, the content of the hydrogenation active component in the upstream hydrodeoxygenation catalyst by weight of the oxide is 2 to 10 percentage points lower than the content of the hydrogenation active component in the downstream hydrodeoxygenation catalyst by weight, preferably 3 to 8 percentage points lower. Generally, 2 to 5 hydrodeoxygenation catalyst beds can be set.

[0048] Furthermore, in the above technical solution, in order to better exert the protective effect of the hydrogenation protective agent and the hydrogenation effect of the hydrodeoxygenation catalyst, the hydrogenation protective agent and the hydrodeoxygenation agent need to be used together, and the volume ratio thereof is 15:85~70:30.

[0049] Furthermore, in the above technical solution, the carrier of the hydrodeoxygenation catalyst is generally alumina, amorphous silica-alumina, silicon oxide, titanium oxide, etc., and may contain other additives, such as P, Si, B, Ti, Zr, etc. A commercially available catalyst may be used, or it may be prepared according to existing methods in the art. The catalyst with the hydrogenation active component in an oxidized state is subjected to conventional sulfurization treatment before use to convert the hydrogenation active component into a sulfurized state. The main commercial hydrogenation catalysts include FHUDS series such as FF-24, FF-36, FF-46, FF-56, FF-66, FHUDS-6, FHUDS-7, FHUDS-8, FZC series such as FZC-31, FZC-41, FZC-401, IFP's HR-416, HR-448 and other hydrogenation catalysts, CLG's ICR174, ICR178, ICR 179 and other hydrogenation catalysts, UOP's newly developed HC-P, HC-K UF-210 / 220, Topsor's TK-525, TK-555, TK-557 and other hydrogenation catalysts, AKZO's KF-752, KF-840, KF-848, KF-901, KF-907 and other hydrogenation catalysts.

[0050] Furthermore, in the above technical solution, the Fischer-Tropsch synthetic oil in step (2) refers to a fraction obtained by the Fischer-Tropsch synthesis technology, which may be a full fraction, a light fraction, or a heavy fraction. The final distillation point of the Fischer-Tropsch synthetic oil is preferably not higher than 750°C, preferably not higher than 720°C.

[0051] Furthermore, in the above technical solution, the Fischer-Tropsch synthetic oil hydrotreating process described in step (2) generally uses a fixed bed hydrogenation reactor, and the hydrotreating operating conditions are as follows: an average reaction temperature of 150°C to 450°C, a system reaction pressure of 3.0MPa to 20.0MPa, a hydrogen-to-oil volume ratio of 100 to 1000, and a liquid hourly volume space velocity of 0.4h -1 ~10.0h -1 The preferred operating conditions are: average reaction temperature of 180°C to 430°C, reaction pressure of 3.0MPa to 18.0MPa, hydrogen to oil volume ratio of 200 to 1000, liquid hourly volume space velocity of 0.5h -1 ~8.0h -1 .

[0052] Furthermore, in the above technical solution, the reaction material first passes through the hydrogenation protective agent. The protective agent is a conventional protective agent. Only one can be selected, or two or more protective agents can be selected according to the order of contact with the raw oil, according to the particle size from large to small, the content of hydrogenation metal from small to large, and the hydrogenation performance from weak to strong. Multiple protective agents form a protective agent system. For example, the three protective agent systems FZC-100 / FZC-105 / FZC-106 (volume ratio of 10 / 35 / 55) developed by FRIPP can be used here. Usually, the total volume air velocity of fresh raw materials of the protective agent is 3.0h -1 ~30.0h -1 , preferably 4.0h -1 ~20.0h -1 .

[0053] Furthermore, in the above technical solution, in order to better exert the protective effect of the hydrogenation protective agent and the hydrogenation effect of the hydrodeoxygenation catalyst, the hydrogenation protective agent and the hydrodeoxygenation agent need to be used together, and the volume ratio thereof is 15:85~70:30.

[0054] Furthermore, in the above technical solution, the carrier of the hydrodeoxygenation catalyst is generally alumina, amorphous silica-alumina, silicon oxide, titanium oxide, etc., and may contain other additives, such as P, Si, B, Ti, Zr, etc. A commercially available catalyst may be used, or it may be prepared according to existing methods in the art. The catalyst with the hydrogenation active component in an oxidized state is subjected to conventional sulfurization treatment before use to convert the hydrogenation active component into a sulfurized state. The main commercial hydrogenation catalysts include FHUDS series such as FF-24, FF-36, FF-46, FF-56, FF-66, FHUDS-6, FHUDS-7, FHUDS-8, FZC series such as FZC-31, FZC-41, FZC-401, IFP's HR-416, HR-448 and other hydrogenation catalysts, CLG's ICR174, ICR178, ICR 179 and other hydrogenation catalysts, UOP's newly developed HC-P, HC-K UF-210 / 220, Topsor's TK-525, TK-555, TK-557 and other hydrogenation catalysts, AKZO's KF-752, KF-840, KF-848, KF-901, KF-907 and other hydrogenation catalysts.

[0055] Furthermore, in the above technical solution, the reaction pressure in step (1) may be the same as or different from the reaction pressure in step (2). If the reaction pressure in step (1) is the same as the reaction pressure in step (2), the two parts may share a circulating hydrogen system or use a separate circulating hydrogen system.

[0056] Furthermore, in the above technical solution, the number of catalyst beds described in step (3) is n layers, and solid fillers are preferably placed on the upper part of each catalyst bed, and hydrocracking catalysts are loaded on the lower part. The fillers are conventional in the art, such as ball rings, Raschig rings, rectangular saddle rings, saddles, open-hole ring types, half rings, step rings, double arcs, Haier rings, conjugated rings, flat rings, rosettes and other random packings, or metal or ceramic corrugated fillers. The hydrocracking catalyst generally includes an active component and a carrier, the carrier component includes one or more of alumina, silicon-containing alumina and molecular sieves, preferably containing molecular sieves, and the molecular sieves may be Y-type molecular sieve agents; the active component is one or more of Group VIB and Group VIII metals, Group VIB metals are generally Mo and / or W, and Group VIII metals are generally Co and / or Ni. The shape of the hydrocracking catalyst may be any conventional existing hydrocracking catalyst shape, preferably a porous catalyst, a shaped catalyst and / or a honeycomb catalyst. The pore size of the porous catalyst is 1 to 50 mm, preferably 4 to 20 mm; the average particle diameter of the shaped catalyst is 2 to 50 mm, preferably 4 to 30 mm; the pore diameter or pore side length of the honeycomb catalyst is 1 to 50 mm, preferably 3 to 15 mm; the recommended porosity of the catalyst bed is 15% to 85%, preferably 20% to 75%.

[0057] Furthermore, in the above technical solution, 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, liquid hourly volume space velocity of 0.1h -1 ~10.0h -1 The preferred operating conditions are: reaction temperature of 300°C to 450°C, reaction pressure of 3.0MPa to 18.0MPa, hydrogen to oil volume ratio of 100 to 1500, liquid hourly volume space velocity of 0.5h -1 ~10.0h -1 .

[0058] Furthermore, in the above technical solution, after the hydrodeoxygenation in step (1), the hydrodeoxygenation product can be selectively separated, and the separated liquid fraction enters the hydrogenation reaction distillation tower for further hydrogenation treatment.

[0059] Furthermore, in the above technical solution, after the hydrogenation refining in step (2), the water generated by the hydrogenation reaction can be selectively separated, and the separated liquid fraction enters the hydrogenation reaction distillation tower for further hydrogenation treatment.

[0060] After a lot of research, it was found that for the gas-liquid-solid three-phase reaction process in which the amount of liquid phase decreases rapidly and the amount of gas phase increases rapidly during the reaction, the rapid increase in the amount of gas phase occupies a large amount of bed space, which greatly increases the liquid phase flow rate. According to the traditional design, although the gas-liquid-solid three-phase contact can be guaranteed, the effective reaction time of the liquid phase that needs further conversion is reduced, and the gas phase that does not need to react again (such as the gas phase obtained by liquid phase conversion under reaction conditions) has an increased chance of contact with the catalyst. For systems that require more liquid phase conversion and gas phase control of secondary reactions, the overall reaction effect is subject to certain limitations, which is generally manifested in that the reaction conversion rate and selectivity are difficult to further improve.

[0061] Research has found that when the overall space velocity is similar, the gas-liquid-solid three-phase hydrogenation reaction in which the liquid phase amount decreases rapidly and the gas phase amount increases rapidly during the reaction process allows the generated gas phase to quickly leave the catalyst bed, and the adverse effects of the generated gas phase have little cumulative effect. The liquid phase has a more sufficient chance to react on the catalyst, thereby overcoming the traditional understanding that a small aspect ratio will lead to adverse effects such as poor contact effects, and achieving a significant increase in the yield of the target product. At the same time, it solves the problems of easy flooding of the countercurrent reactor and limited hydrogen-to-oil ratio.

[0062] Compared with the prior art, the advantages of the present invention are:

[0063] 1. The gaseous products produced by the hydrocracking reaction of the present invention can leave the cracking reaction zone in time and will not enter the subsequent catalyst bed again, so as to avoid the gaseous products occupying the cracking catalyst pores, effectively prevent the secondary cracking and gasification of the target product, and improve the selectivity of the reaction and the yield of the target product. At the same time, since the product partial pressure is always kept at a low state, the driving force of the reaction is increased and the equilibrium conversion rate is improved.

[0064] 2. The components of bio-oil after hydrodeoxygenation contain up to more than 90% normal alkanes, and the rest are isoalkanes. The components of Fischer-Tropsch synthetic oil after hydrorefining contain up to more than 90% normal alkanes, and the rest are isoalkanes. The gas phase products produced after the cracking of the alkanes in these two hydrogenation products are rich in a large number of low-carbon alkanes. These low-carbon alkanes are taken away from the distillation tower through the gas phase channel and are almost not cracked into smaller molecules on the subsequent cracking catalyst, which increases the load of the subsequent cracking reaction zone.

[0065] 3. The umbrella-shaped partitions provided in the hydrodistillation reaction tower of the present invention can separate the gaseous feed and product gas between adjacent beds on the one hand, and play a guiding role for the liquid phase and the gas phase on the other hand. The liquid seal baffle can effectively isolate the gaseous feed and the gaseous product. The arrangement of the multi-layer concentric annular gas phase distribution pipe can keep the distribution of the gaseous feed uniform to the maximum extent. BRIEF DESCRIPTION OF THE DRAWINGS

[0066] Figure 1The present invention is a schematic flow diagram of a specific embodiment of the production of low aromatic solvent oil by hydrogenation distillation.

[0067] Figure 2 It is a structural schematic diagram of a specific embodiment of the hydrogenation reaction distillation tower of the present invention.

[0068] Figure 3 It is a top view of the liquid phase distribution pipe in the hydrogenation reaction distillation tower of the present invention.

[0069] Figure 4 It is a top view of the gas phase feed pipe and the gas phase distribution pipe in the hydrogenation reaction distillation tower of the present invention (showing the situation where the gas phase feed pipe and the annular gas phase distribution pipe are orthogonally arranged).

[0070] Figure 5 It is another top view of the gas phase feed pipe and the gas phase distribution pipe in the hydrogenation reaction distillation tower of the present invention (showing the situation where the gas phase feed pipe and the annular gas phase distribution pipe are arranged to intersect tangentially).

[0071] Figure 6 It is a top view of the gas phase distribution pipe of the present invention which adopts a concentric double-ring distribution pipe.

[0072] Figure 7 It is a top view of the gas phase distribution plate in the hydrogenation reaction distillation tower of the present invention.

[0073] Figure 8 It is a top view of the catalyst support plate in the hydrogenation reaction distillation tower of the present invention.

[0074] Description of main reference numerals:

[0075] Description of main reference numerals:

[0076] 1-hydrogenation reaction distillation tower, 2-bio-oil, 255-Fischer-Tropsch oil, 3-new hydrogen, 4-high-pressure hydrogen-rich gas, 5-hydrogenation deoxygenation reactor, 55-hydrogenation refining reactor, 6-hydrogenation deoxygenation logistics, 7-liquid logistics, 8-gas logistics, 9-high-pressure separator, 10-fractionation tower, 101-gas, 102-naphtha, 103-low aromatic solvent oil, 104-unconverted oil.

[0077] 11- umbrella-shaped partition, 110- filler, 111- solid catalyst, 12- outer downcomer, 13- gas phase channel, 14- overflow weir, 15- downcomer folding plate, 16- liquid receiving tray, 17- liquid sealing baffle, 18- inner downcomer, 19- catalyst support plate, 191- grid;

[0078] 21-liquid phase feed pipe, 22-liquid phase distribution pipe, 220-liquid phase distribution pipe body, 221-liquid phase channel;

[0079] 31 - gas phase feed pipe, 32 - gas phase distribution pipe, 320 - gas phase distribution pipe body, 321 - gas phase channel, 33 - gas phase distribution plate, 331 - eyelet. DETAILED DESCRIPTION

[0080] The specific implementation modes of the present invention are described in detail below in conjunction with the accompanying drawings, but it should be understood that the protection scope of the present invention is not limited by the specific implementation modes.

[0081] Figure 1 The process flow diagram of producing low aromatic solvent oil of the present invention is shown in FIG. Figure 1 and Figure 2 As shown, the bio-oil 2 is mixed with new hydrogen 3 and circulating pressurized hydrogen-rich gas 4 and then enters the hydrodeoxygenation reactor 5, in which a hydrogenation protective agent and a hydrodeoxygenation catalyst are placed in order from top to bottom. The Fischer-Tropsch oil 255 is mixed with new hydrogen 3 and circulating pressurized hydrogen-rich gas 4 and then enters the hydrofining reactor 55, in which a hydrogenation protective agent and a hydrofining catalyst are placed in order from top to bottom. The hydrodeoxygenation stream 6 and the hydrofining stream enter the liquid phase feed pipe 21 of the hydrogenation reaction distillation tower 1 after being dehydrated, and flow downward along the filler or catalyst bed after being evenly distributed through the liquid phase distribution pipe 22. When the dehydrated hydrodeoxygenation flow 6 passes through the packing of the separation zone, the light component flows out upward through the gas phase channel 321, and the heavy component continues to flow downward. When entering the catalyst bed of the hydrogenation reaction zone, it reacts with the hydrogen that enters the catalyst bed from the gas phase feed pipe 31 upward. The gas phase product directly enters the gas phase channel 321, and the heavy component continues to flow downward along the catalyst bed and reacts.

[0082] The liquid phase stream 7 generated by the hydrocracking reaction flows out from the bottom of the hydrocracking distillation tower, and the gas phase stream 8 from the gas phase channel 321 flows out from the top of the hydrocracking distillation tower. The two are mixed and enter the high-pressure separator 9 for gas-liquid separation. The separated high-pressure hydrogen-rich gas 4 is mixed with the new hydrogen 3 and used as circulating hydrogen. The separated liquid enters the distillation tower 10 for fractionation to obtain gas 101, naphtha 102, low aromatic solvent oil 103 and unconverted oil 104. The unconverted oil 104 is fully or partially recycled to the hydrocracking distillation tower 1 to continue the hydrocracking reaction, and can also be drawn out of the device as a raw material for producing lubricating oil base oil.

[0083] like Figure 2As shown, the internal components of the hydrogenation reaction distillation tower 1 of the present invention include a catalyst bed, a liquid-phase feed subunit, a gas-phase feed subunit and a gas-phase channel. The upper part of the hydrogenation reaction distillation tower is set as a separation zone, and the lower part is a hydrogenation reaction zone. Among them, the catalyst bed in the separation zone is used to place the filler 110, and the catalyst bed in the hydrogenation reaction zone is used to fill the solid catalyst 111. There is an inclined partition on the top of each catalyst bed. The overall shape formed by the inclined partition can be umbrella-shaped, which plays the role of a partition. On the one hand, it can separate the gas-phase feed and product gas between adjacent beds, and on the other hand, it plays a role in guiding the liquid phase and the gas phase. Preferably, but not restrictively, the umbrella surface can be arc-shaped or folding umbrella-shaped. The liquid-phase feed subunit is arranged above the inclined partition (i.e., the umbrella-shaped partition 11) of the topmost catalyst bed. The liquid-phase feed is guided to the catalyst bed through the umbrella-shaped partition 11 to contact the filler 110. Specifically, an annular outer downcomer 12 is provided at the end of the umbrella-shaped partition 11. The bottom of the outer downcomer 12 is spaced a distance from the bottom of the catalyst bed, so that the liquid-phase feed enters the catalyst bed along the radial direction of the reaction distillation tower 1.

[0084] Each catalyst bed in the hydrogenation reaction zone is provided with a gas-phase feed subunit, which is specifically arranged between the catalyst bed of the upper layer and the umbrella-shaped partition 11 of the lower layer, and the gas-phase feed of each layer enters the catalyst bed upward. After the gas-liquid phase feed and the solid catalyst are fully reacted 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 partition 11. The gas-phase channel 13 is in a relatively isolated state from the gas-phase feed subunit, that is, the gas-phase product generated after the gas-phase feed and the liquid-phase feed react in the catalyst bed directly enters the gas-phase channel 13. Preferably, but not restrictively, the gas-phase channel is located in the middle of the reaction distillation tower 1, and runs through all the catalyst beds from bottom to top.

[0085] like Figure 2 , 3As shown, the liquid-phase feed subunit further includes a liquid-phase feed pipe 21 and a liquid-phase distribution pipe 22. The liquid-phase feed pipe 21 extends in the radial direction of the hydrogenation reaction unit, and the liquid-phase distribution pipe 22 is annular. The liquid-phase feed pipe 21 intersects orthogonally or tangentially with the pipe body 220 of the liquid-phase distribution pipe 22. The pipe wall of the liquid-phase distribution pipe 22 is provided with a plurality of liquid-phase channels 221 for evenly distributing the liquid-phase feed to all directions of the annular outer downcomer 12. The openings of the liquid-phase channels 221 can be in all directions of the upper, lower and side surfaces of the pipe body. The liquid-phase feed enters the reaction distillation tower 1 through the liquid-phase feed pipe 21, is distributed into the tower through the annular liquid-phase distribution pipe 22, flows into the outer downcomer 12 toward the periphery through the umbrella-shaped partition 11, and enters the catalyst bed laterally after passing through the outer downcomer 12 to contact with the filler and the solid catalyst. The feeding direction of the liquid phase feed pipe 21 is the radial direction of the reaction distillation tower, which is orthogonal to or tangential to the radial direction of the annular liquid phase distribution pipe 22. The annular diameter of the annular liquid phase distribution pipe 22 is larger than the outer diameter of the gas phase channel 13 and smaller than the inner diameter of the reaction distillation tower 1. The multiple holes on the wall of the annular liquid phase distribution pipe 22 facilitate the uniform distribution of the liquid phase feed in all directions of the outer downcomer 12. The height of the downcomer fold 15 is generally smaller than the height of the packing or catalyst loading of this layer, and the distance between the downcomer fold 15 and the inner wall of the reaction distillation tower 1 is determined according to the flow rate of the liquid phase reactants of this layer.

[0086] Further Figure 2 As shown, the height of each catalyst bed in the reaction distillation tower 1 can be the same or different, depending on the different chemical reaction systems. The catalyst bed is fixed with a screen to keep the bed relatively stable, and the bed height is set to 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 on the side close to the gas phase channel 13. The liquid seal baffle 17 is arranged on the upper part of the overflow weir 14 to isolate the gas phase feed from the gas phase product. Further, the liquid seal baffle 17 includes a horizontal part and a vertical part, the horizontal part is an annular flat plate and is located above the overflow weir 14; the vertical part is cylindrical, and the vertical part is integrally formed with the horizontal part, and other seamless connection methods can also be used. The lower end of the vertical part is spaced a distance from the bottom of the catalyst bed to ensure the outflow of the liquid phase product. The unreacted liquid feed and the reacted but liquid-phase material in the catalyst bed pass through the overflow weir 14, through the inner downcomer 18 (i.e., the annular space between the overflow weir 14 and the outer wall of the gas phase channel 13), along the umbrella-shaped partition 11, and through the outer downcomer 12 of the next layer into the next catalyst bed. The height of the overflow weir 14 is higher than the upper plane of the catalyst in the bed, preferably 10 to 100 mm higher. The annular inner downcomer 18 formed between the overflow weir 14 and the outer wall of the gas phase channel 13 has a spacing size determined by the size of the liquid phase load, and the size of the downcomer of each bed can be the same or different.

[0087] Further Figure 2 , 4 As shown in FIG. 6 , the gas phase feed subunit includes a gas phase feed pipe 31 and a gas phase distribution pipe 32. The gas phase feed pipe 31 extends in the radial direction of the reaction distillation tower 1. The gas phase distribution pipe 32 is annular (see FIG. Figure 4 , 5 ) or multiple layers of concentric rings (see Figure 6 The gas phase feed pipe 31 is orthogonal to the gas phase distribution pipe body 320 of the gas phase distribution pipe 32 (see Figure 4 ) or tangential intersection (see Figure 5 ), a plurality of gas phase channels 321 are provided on the wall of the gas phase distribution pipe 32, which are used to evenly distribute the gas phase feed to all directions at the bottom of the catalyst bed. Preferably, but not limiting, the gas phase distribution pipe 32 can be arranged below the catalyst bed, or can be arranged inside the catalyst bed. Figure 7 As shown, the gas-phase feed subunit further includes a gas-phase distribution plate 33, which is located at the bottom of the catalyst bed and is in the shape of a plate as a whole. The gas-phase distribution plate is uniformly and densely distributed with a plurality of holes 331. The gas-phase feed enters the reaction distillation tower 1 through the gas-phase feed pipes 31 of each layer, is distributed into the reaction distillation tower 1 through the annular gas-phase distribution pipe 32, and enters the catalyst bed upward through the gas-phase distribution plate 33 at the lower part of the catalyst support plate 19. The gas-phase feed pipe 31 enters the reaction distillation tower 1 in a radial direction, and intersects orthogonally or tangentially with the annular gas-phase distribution pipe 32. The annular gas-phase distribution pipe 32 is located below the catalyst bed. The annular diameter of the annular gas-phase distribution pipe 32 is smaller than the outer ring diameter of the catalyst bed, and the inner diameter is larger than the inner ring diameter of the catalyst bed. The several gas-phase channels 321 on the wall of the annular gas-phase distribution pipe 32 facilitate the uniform distribution of gas at various positions of the gas-phase distribution plate 33. The catalyst support plate 19 mainly supports the catalyst bed to ensure that the catalyst bed remains stable in the axial direction of the reaction distillation tower. The function of the gas phase distribution plate 33 is to ensure uniform distribution of the gas phase feed, while trying to avoid direct leakage of the liquid phase feed on the catalyst bed (using the gas phase distribution plate 33 of the present invention, the liquid leakage is less than 15%). When more than one concentric annular gas phase distribution pipe 32 with different diameters is arranged on the same plane, the distribution of the gas phase feed can be more uniform. Figure 2 In the embodiment of the present invention, the annular gas phase distribution pipe 32 is arranged below the catalyst bed layer. When the annular gas phase distribution pipe 32 is installed in the catalyst bed layer, the catalyst support plate 19 can Figure 8 The grid 191 is changed into a support plate, and the gas phase distribution plate 33 is cancelled.

[0088] In the hydrogenation reaction unit of the present invention, liquid phase feed and gas phase feed undergo hydrogenation reaction in the catalyst bed, gas phase products and unreacted gas phase feed rise through the gas phase channel 13 and leave the reaction system, and gas phase products generated after the reactants undergo chemical reaction in the catalyst bed will leave the reaction zone in time and will not enter the upper catalyst bed (isolated by the umbrella-shaped partition) again, thus avoiding secondary reaction of the target product and improving the selectivity of the reaction. At the same time, due to the departure of the product in the reaction zone, the driving force of the reaction is increased and the equilibrium conversion rate is improved.

[0089] The hydrogenation reaction distillation tower of the present invention can be a multi-layer plate tower structure. The number of catalyst beds in the hydrogenation reaction distillation tower is two or more layers. The hydrogenation reaction distillation tower 1 of the present invention is suitable for a reaction system in which at least one liquid feed and at least one gas feed undergo a chemical reaction on a solid catalyst, and at least one gas product is included in the reaction product. For example, hydrocracking of petroleum fractions and chemical synthetic oils, hydrodewaxing of diesel and lubricating oil fractions, and hydrotreatment processes of various petroleum fractions.

[0090] In addition to the downcomers, overflow weirs and liquid receiving tray 16, each tray of the hydrogenation reaction distillation tower 1 of the present invention is also provided with a liquid sealing baffle, which is connected to the gas phase channel. Adjacent trays are separated by umbrella-shaped partitions. Each tray is an annular structure, the inner edge of the annular ring is connected to the gas phase channel, and the outer edge is connected to the inner wall of the hydrogenation reaction distillation tower. The gas phase channel is a common channel for removing the gas phase products generated by the chemical reaction on each tray. In the embodiment of the present invention, the liquid feed positions are all on the upper part of a tray, and liquid feed can also be provided on some trays or each tray, and the gas phase feed has feed at the lower part of each tray. Above each tray is a catalyst loading area, and the liquid phase feed flows radially through the catalyst bed. The gas phase feed enters from the bottom of the tray and reacts under the action of the catalyst. The gas phase material generated after the reaction directly leaves the reaction system and enters the gas phase channel in the middle. After the liquid phase leaves the current bed, it enters the next bed through the downcomer. Since the reaction and separation are carried out simultaneously, the reaction equilibrium can be destroyed, effectively improving the conversion rate of the reactants and the selectivity of the target products.

[0091] Example 1

[0092] The above-mentioned device and process of the present invention are used to hydrogenate bio-oil and Fischer-Tropsch oil. The bio-oil raw material oil is soybean oil, and the Fischer-Tropsch oil is the diesel part in the Fischer-Tropsch full fraction.

[0093] A hydrogenation protective agent is placed in the upper part of the hydrodeoxygenation reactor, and a hydrodeoxygenation catalyst is loaded in the lower part. The volume ratio of the hydrogenation protective agent to the hydrodeoxygenation catalyst is 1:10. The biomass raw material undergoes olefin hydrogenation saturation and hydrodeoxygenation reactions in the hydrodeoxygenation reactor. The hydrogenation protective agent is FZC-105 / FZC-106, and the hydrodeoxygenation catalyst is FZC-41. The above catalysts are all produced by Sinopec Dalian Petrochemical Research Institute. The specific operating process conditions are shown in Table 2.

[0094] A hydrogenation protective agent is placed in the upper part of the hydrotreating reactor, and a hydrotreating catalyst is loaded in the lower part. The volume ratio of the hydrogenation protective agent to the hydrotreating catalyst is 1:10. The Fischer-Tropsch oil raw material undergoes olefin hydrogenation saturation and hydrodeoxygenation reactions in the hydrotreating reactor. The hydrogenation protective agent is FZC-105 / FZC-106, and the hydrotreating catalyst is FF-66. The above catalysts are all produced by Sinopec Dalian Petrochemical Research Institute. The specific operating process conditions are shown in Table 2.

[0095] The hydrogenation reaction distillation tower is provided with five trays, the upper part of the first tray is filled with filler, the lower part is filled with hydrocracking catalyst, and the remaining four trays are filled with hydrocracking catalyst. The filler is 5mm corundum Raschig ring, and the hydrocracking catalyst is FC-46 produced by Sinopec Dalian Petrochemical Research Institute. The specific operating process conditions are shown in Table 2.

[0096] The light fraction of the hydrogenation reaction distillation tower passes through a high-pressure separator to obtain high-pressure hydrogen-rich gas 4 for recycling, and the liquid enters a fractionation tower for fractionation to obtain gas, naphtha, and low aromatic solvent oil; the heavy fraction at the bottom of the tower is directly recycled back to the hydrogenation reaction distillation tower.

[0097] Product distribution and properties are shown in Table 3.

[0098] Example 2-3

[0099] 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.

[0100] Example 4

[0101] The biomass raw material oil was replaced with palm oil, as shown in Table 2. The product distribution and properties are shown in Table 3.

[0102] Table 1 Main properties of crude oil

[0103]

[0104] Table 2 Example process conditions

[0105]

[0106] Table 3 Product distribution and properties

[0107]

[0108] From the results in Table 3, it can be seen that by adopting the method of the present invention, animal and vegetable fats and Fischer-Tropsch oil can be hydrogenated to produce low aromatic solvent oil, and the aromatic content of the solvent oil is very low, which is a high-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 after hydrogenation and deoxygenation of bio-oil and a hydrorefining product of Fischer-Tropsch synthetic oil, 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 evenly from the bottom of each catalyst bed from bottom to top; The liquid feed and hydrogen undergo a hydrocracking reaction in a hydrocracking catalyst bed loaded in a hydrocracking 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 hydrocracking distillation tower through an independent gas phase channel, and the liquid product obtained by the hydrocracking is discharged from the bottom of the hydrocracking 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, characterized in that: The system also includes (5) a gas circulation unit, which is used to selectively concentrate the hydrogen-rich gas obtained from the separation unit and circulate it back to the hydrodeoxygenation reactor, the hydrofining reactor and the hydrogenation reaction distillation tower.

3. The hydrogenation device according to claim 1, characterized in that: The invention also comprises (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, characterized in that An annular downcomer is arranged at the end of the inclined partition plate, 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, characterized in that The liquid seal baffle comprises: a horizontal portion which is annular and flat and located above the overflow weir; a vertical portion which is cylindrical and is integrally formed with the horizontal portion, and a 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, characterized in that: The liquid phase feed subunit further includes: a liquid phase distribution pipe, which is annular and orthogonal or tangentially intersects with the liquid phase feed pipe, and a 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, 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 in an annular shape or a multi-layer concentric annular shape, and the gas-phase distribution pipe intersects the gas-phase feed pipe orthogonally or tangentially, and a plurality of holes are arranged on the wall surface 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 in the shape of a plate 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: The following steps are involved: (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 The method further comprises 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 The method further comprises 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 , the average reaction temperature is 180℃~450℃.

15. The method according to claim 10, characterized in that The final boiling 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, liquid hourly volume space velocity of 0.4h -1 ~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, liquid hourly volume space velocity of 0.1h -1 ~10.0h -1 .

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