A device and method for producing low aromatic solvent oil by hydrogenating bio-oil

Abstract

The present invention discloses a device and method for producing low aromatic solvent oil by hydrogenating bio-oil. Bio-oil is used as raw material, and a treatment mode combining hydrodeoxygenation and hydrocracking is adopted; wherein, the hydrodeoxygenation liquid product enters the hydrocracking section and contacts with hydrogen in countercurrent to carry out 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 technical field of oil refining, and relates to a hydrogenation distillation method, in particular to a method for producing low aromatic solvent oil from bio-oil by hydrogenation technology. Background Art

[0002] Bio-oils include animal oils and vegetable oils, which are renewable resources and one of the key development directions of renewable energy. They have received widespread attention from the world. Bio-oils can be used to produce lipid biodiesel through transesterification, or to remove oxygen through hydrogenation to obtain hydrocarbon products similar to petroleum products, such as the production of motor fuels such as jet fuel and diesel. Bio-oils are 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, the hydrodeoxygenation of vegetable oils and then hydrocracking to break large molecules into small molecules, and produce low-carbon alkanes that meet product requirements.

[0003] The existing bio-oil hydrogenation technology, US5705722 discloses a diesel blending component containing unsaturated fatty acids, fats, vegetable oils and animal oils 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.

[0004] The existing technology of hydrogenating bio-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, the bio-oil raw material is first subjected to a hydrodeoxygenation reaction under hydrogenation conditions, and then the resulting oil is dehydrated and hydrocracking is carried out under hydrocracking conditions to crack large molecules into small molecules to produce low aromatic solvent oil.

[0005] At present, the general process flow of bio-oil hydrogenation is usually that after the bio-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 cracking product enters the fractionation system to obtain low-carbon alkanes. The unconverted tail oil can be partially or fully recycled 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 all the raw materials are converted. Summary of the invention

[0006] In view of the shortcomings of the prior art, the present invention aims to provide a device and method for producing low aromatic solvent oil by hydrogenating bio-oil. The method improves the selectivity of the low aromatic solvent oil target product by using a hydrogenation distillation 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.

[0007] The first aspect of the present invention is to provide a device for producing low aromatic solvent oil by hydrogenating bio-oil.

[0008] A device for producing low aromatic solvent oil by hydrogenating bio-oil, comprising:

[0009] (1) a hydrodeoxygenation reactor, which is used to perform a hydrodeoxygenation reaction on the raw bio-oil;

[0010] (2) a hydrogenation reaction distillation tower, which receives the liquid phase product after the hydrodeoxygenation of the bio-oil as a liquid phase feed, and the liquid phase feed flows from top to bottom through each catalyst bed in the upper part of the reactor; the preheated hydrogen gas enters each catalyst bed from the bottom to the top uniformly from the bottom of each catalyst bed; wherein the liquid phase feed and hydrogen are subjected to a hydrocracking reaction in the hydrocracking catalyst bed loaded in the reactor; 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 the hydrocracking is discharged from the bottom of the hydrogenation reaction distillation tower;

[0011] (3) 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.

[0012] Furthermore, the above technical solution also includes (4) 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 and the hydrogenation reaction distillation tower.

[0013] Furthermore, the above technical solution also includes (5) a liquid circulation unit, which is used to circulate all or part of the liquid product obtained from the hydrogenation reaction distillation tower back to the hydrodeoxygenation reactor or the hydrogenation reaction distillation tower.

[0014] Furthermore, in the above technical solution, the hydrogenation reaction distillation tower comprises at least one reaction zone, and the 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 on the top of each catalyst bed, preferably an umbrella-shaped partition; each catalyst bed and the inclined partition are provided with a gas phase channel.

[0015] Furthermore, in the above technical solution, the hydrodeoxygenation reaction 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.

[0016] 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.

[0017] 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.

[0018] 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.

[0019] 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.

[0020] 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.

[0021] 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.

[0022] 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.

[0023] 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.

[0024] 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.

[0025] Furthermore, in the above technical solution, 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.

[0026] In the above technical solution, the uppermost end of the gas phase channel is connected to the outlet pipeline of the hydrogenation reaction distillation tower, and the lowermost end is close to the bottom of the hydrogenation reaction distillation tower. 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 reaction distillation tower.

[0027] 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.

[0028] 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.

[0029] Furthermore, in the above technical solution, the gas phase distribution pipe can be arranged below the catalyst bed or inside the catalyst bed.

[0030] Furthermore, in the above technical solution, each catalyst bed of the hydrogenation reaction distillation tower can selectively be placed with fillers and hydrocracking catalysts, and preferably solid fillers are placed on top of the catalyst.

[0031] Furthermore, in the above technical solution, a gas-liquid separation device needs to be provided between the hydrodeoxygenation reactor and the hydrogenation reaction distillation tower to remove water generated by the hydrodeoxygenation reaction.

[0032] The second aspect of the present invention is to provide a method for producing low aromatic solvent oil by hydrogenating bio-oil. The method comprises the following steps:

[0033] (1) The bio-oil raw material is mixed with hydrogen and enters the hydrodeoxygenation reactor, where it passes through the hydrodeoxygenation catalyst bed under hydrodeoxygenation conditions;

[0034] (2) The liquid product of the hydrodeoxygenation stream obtained in step (1) is dehydrated and enters the hydrogenation reaction distillation tower as a liquid feed, flows through each catalyst bed of the reactor from top to bottom, and enters from the bottom to top from the bottom of each catalyst bed after being preheated, and contacts and reacts on the catalyst of each bed;

[0035] (3) The gas phase product obtained by the reaction of each catalyst bed enters the gas phase channel, flows upward in the gas phase channel and flows out of the hydrogenation reaction distillation tower; the liquid phase product obtained by the reaction of each catalyst bed sequentially enters the next catalyst bed layer for reaction, and the final liquid product flows downward out of the hydrogenation reaction distillation tower;

[0036] (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.

[0037] Furthermore, the above technical solution also includes step (5): the hydrogen-rich gas obtained in step (4) is optionally concentrated and then returned to step (1) and step (2).

[0038] Furthermore, the above technical solution also includes step (6): the liquid product obtained in step (3) is recycled back to step (1) or step (2).

[0039] Furthermore, in the above technical solution, the bio-oil includes at least one of vegetable oil and animal oil. Vegetable oil raw materials include 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 tallow, lard, sheep oil and fish oil. The bio-oil raw material can also be acidified oil, waste cooking oil and the like.

[0040] Furthermore, in the above technical solution, the hydrodeoxygenation reactor described in step (1) is a conventional fixed bed reactor. The hydrodeoxygenation operating conditions are generally: reaction pressure of 3.0 MPa to 20.0 MPa, hydrogen to oil volume ratio of 200:1 to 3000:1, volume space velocity of 0.1 h -1 ~6.0h-1 The average reaction temperature is 180℃~450℃; the preferred operating conditions are: reaction pressure 4.0MPa~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℃.

[0041] Furthermore, in the above technical solution, in step (1), the reaction material first passes through a hydrogenation protective agent. The protective agent is a conventional protective agent. Only one protective agent can be selected, or two or more protective agents can be selected according to the order of contact with the crude oil, in the order of particle size from large to small, hydrogenation metal content from small to large, and 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 the fresh raw material of the protective agent is 3.0h -1 ~30.0h -1 , preferably 4.0h -1 ~20.0h -1 .

[0042] Furthermore, in the above technical scheme, the logistics through the hydrogenation protection agent continues to pass through the hydrodeoxygenation catalyst, and the hydrodeoxygenation catalyst can be used in one, or in two or more forms. The hydrogenation active components of the hydrodeoxygenation catalyst are generally Mo, W, Ni, Co, etc. of the VIB group and the VIII group, 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 hydrodeoxygenation 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 the two adjacent hydrodeoxygenation catalysts, the content of the hydrogenation active component in the upstream hydrodeoxygenation catalyst in terms of oxide weight is 2 to 10 percentage points lower than the content of the active component in the downstream catalyst in terms of oxide weight, preferably 3 to 8 percentage points lower. Generally, 2 to 5 hydrodeoxygenation catalyst beds can be set.

[0043] 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 of the two is 15:85~70:30.

[0044] 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.

[0045] Furthermore, in the above technical solution, the number of catalyst beds described in step (2) is n layers, and solid fillers can be preferably placed on the upper part of each bed, and hydrocracking catalysts are loaded on the lower part of each bed. 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 can 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 can 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 diameter or side length of the honeycomb catalyst hole is 1 to 50 mm, preferably 3 to 15 mm; the void ratio of the catalyst bed is recommended to be 15% to 85%, preferably 20% to 75%. It is further preferred that the cracking activity of the hydrocracking catalyst in different beds is different, and the cracking activity of the catalyst in the adjacent bed is lower than that of the catalyst in the upper bed, based on the cracking activity of the catalyst in the upper bed. That is, when the same single-pass conversion rate is achieved, usually with a single-pass conversion rate of 65 wt% at >165°C, the average reaction temperature of the hydrocracking catalyst in the downstream bed of the two adjacent catalyst beds is 3°C to 10°C lower than the average reaction temperature of the catalyst in the upstream bed. In this way, the temperature of the next catalyst bed is higher than that of the previous catalyst bed, which is conducive to the temperature gradient distribution of the reaction distillation tower, that is, the temperature gradually decreases from bottom to top.

[0046] Furthermore, in the above technical solution, the operating conditions of the catalytic distillation reactor in step (2) are: reaction temperature of 260°C to 450°C, reaction pressure of 3MPa to 20MPa, 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 4MPa to 15MPa, hydrogen to oil volume ratio of 100 to 1500, liquid hourly volume space velocity of 0.5h -1 ~10.0h -1 .

[0047] 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.

[0048] Furthermore, in the above technical solution, step (1) and step (2) may use catalysts of the same metal form or catalysts of different metal forms. Among them, the hydrodeoxygenation catalyst in step (1) may use a sulfided catalyst or a reduced catalyst. The hydrocracking catalyst in step (2) may use a sulfided catalyst or a reduced catalyst. If step (1) uses a sulfided hydrodeoxygenation catalyst and step (2) uses a sulfided hydrocracking catalyst, the two reactors may share a set of circulating hydrogen systems or use two sets of circulating hydrogen systems. If step (1) uses a sulfided hydrodeoxygenation catalyst and step (2) uses a reduced hydrocracking catalyst, the two reactors may use two sets of circulating hydrogen systems, and the two systems cannot be mixed or contaminated with each other to prevent hydrogen sulfide from poisoning the reduced catalyst. If step (1) uses a reducing hydrodeoxygenation catalyst and step (2) uses a sulfided hydrocracking catalyst, two reactors may use two circulating hydrogen systems, and the two systems may not be mixed or contaminated with each other to prevent hydrogen sulfide from poisoning the reducing catalyst. If step (1) uses a reducing hydrodeoxygenation catalyst and step (2) uses a reducing hydrocracking catalyst, the two reactors may share one circulating hydrogen system or use two circulating hydrogen systems.

[0049] Furthermore, the separation described in step (4) generally includes gas-liquid separation and fractionation processes. Gas-liquid separation generally includes high-pressure separation and low-pressure separation processes. High-pressure separation, low-pressure separation and fractionation processes are all operations well known to those skilled in the art.

[0050] 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.

[0051] 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.

[0052] Compared with the prior art, the present invention has the following beneficial effects:

[0053] 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 avoid 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.

[0054] 2. The components of bio-oil after hydrodeoxygenation contain more than 90% normal alkanes, and the rest are isoalkanes. The gas products produced after cracking of these alkanes are rich in a large number of low-carbon alkanes. These low-carbon alkanes are carried away from the distillation tower through the gas phase channel and will hardly be further cracked into smaller molecules on the subsequent cracking catalyst, thereby increasing the load of the subsequent cracking reaction zone.

[0055] 3. In the hydrogenation reaction distillation tower of the present invention, the umbrella-shaped partition provided can separate the gaseous feed and product gas between adjacent beds on the one hand, and play the role of guiding the liquid phase and the gas phase on the other hand. The liquid seal baffle provided can effectively isolate the gaseous feed and the gaseous product. The provision of multiple layers of concentric annular gas phase distribution pipes can keep the distribution of the gaseous feed uniform to the maximum extent.

[0056] 4. In the hydrogenation reaction distillation tower of the present invention, different catalyst beds are loaded with hydrocracking catalysts of different activities, and the reaction temperature gradually increases from top to bottom, which is beneficial to the operation of the distillation tower. The light components are successfully separated from each bed and leave the distillation tower from the top of each bed. BRIEF DESCRIPTION OF THE DRAWINGS

[0057] Figure 1 The figure is a schematic flow chart of a specific embodiment of the bio-oil hydrogenation distillation of the present invention.

[0058] Figure 2 It is a structural schematic diagram of a specific embodiment of the hydrogenation reaction distillation tower of the present invention.

[0059] Figure 3 It is a top view of the liquid phase distribution pipe in the hydrogenation reaction distillation tower of the present invention.

[0060] 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).

[0061] Figure 5It 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).

[0062] 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.

[0063] Figure 7 It is a top view of the gas phase distribution plate in the hydrogenation reaction distillation tower of the present invention.

[0064] Figure 8 It is a top view of the catalyst support plate in the hydrogenation reaction distillation tower of the present invention.

[0065] Description of main reference numerals:

[0066] 1-hydrogenation reaction distillation tower, 2-bio-oil, 3-new hydrogen, 4-high-pressure hydrogen-rich gas, 5-hydrogenation deoxygenation reactor, 6-hydrogenation deoxygenation logistics, 7-liquid product, 8-gas product, 9-high-pressure separator, 10-fractionation tower, 101-gas, 102-naphtha, 103-low aromatic solvent oil, 104-unconverted oil.

[0067] 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;

[0068] 21-liquid phase feed pipe, 22-liquid phase distribution pipe, 220-liquid phase distribution pipe body, 221-liquid phase channel;

[0069] 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

[0070] 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.

[0071] The present invention firstly provides a device for producing low aromatic solvent oil by hydrogenating biological oil.

[0072] Combination Figure 1-2 The present invention provides a device for producing low aromatic solvent oil by hydrogenating bio-oil, comprising:

[0073] A hydrodeoxygenation reactor 5, which is used to perform a hydrodeoxygenation reaction on the raw bio-oil;

[0074] A hydrogenation reaction distillation tower 1 receives a liquid phase product obtained by hydrogenating, deoxygenating and dehydrating bio-oil as a liquid phase feed, and the liquid phase feed flows from top to bottom through each catalyst bed in the reactor; the preheated hydrogen enters each catalyst bed from bottom to top uniformly from the bottom of each catalyst bed; the liquid phase feed and hydrogen undergo a hydrocracking reaction in the hydrocracking catalyst bed in the reactor; 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 13, and the liquid product 7 obtained by hydrocracking is discharged from the bottom of the hydrogenation reaction distillation tower;

[0075] The separation unit 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 4, dry gas 101, light naphtha 102, low aromatic solvent oil 103 and unconverted oil 104.

[0076] like Figure 2 As 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 catalytic reaction zone. Among them, the catalyst bed in the separation zone is used to place the filler 110, and the catalyst bed in the catalytic 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.

[0077] Each catalyst bed in the catalytic 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 passes through all catalyst beds from bottom to top.

[0078] like Figure 2, 3 As 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 catalytic 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 on 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.

[0079] 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.

[0080] 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 to a support plate, and the gas phase distribution plate 33 is cancelled. (Mainly the order is adjusted)

[0081] The second aspect of the present invention also provides a method for producing low aromatic solvent oil by hydrogenating bio-oil.

[0082] Figure 1The figure shows the process flow diagram of the biomass oil hydrogenation catalytic distillation process of the present invention. Figure 1 and Figure 2 As shown, the method comprises the following steps:

[0083] (1) The raw biomass oil 2 is mixed with hydrogen (including new hydrogen 3 and recycled hydrogen-rich gas 4) and then enters the hydrodeoxygenation reactor 5, and passes through the hydrodeoxygenation catalyst bed under hydrodeoxygenation conditions;

[0084] (2) The hydrodeoxygenation stream 6 obtained in step (1) is dehydrated and the liquid product is used as a liquid phase feed through the liquid phase feed pipe 21, and then evenly distributed through the liquid phase distribution pipe 22 to enter the hydrogenation reaction distillation tower 1, and flows through each catalyst bed 110 of the reactor from top to bottom, and contacts and reacts with the preheated hydrogen that enters the catalyst bed from the gas phase feed pipe 31 at the bottom of each catalyst bed from bottom to top on the catalyst in each bed;

[0085] (3) The gas phase product 8 obtained by the reaction of each catalyst bed enters the gas phase channel 13, flows upward in the gas phase channel and flows out of the hydrogenation reaction distillation tower 1; the liquid phase product obtained by the reaction of each catalyst bed sequentially enters the next catalyst bed 110 for reaction, and the final liquid product flows downward out of the hydrogenation reaction distillation tower 1;

[0086] (4) The gas phase product obtained in step (3) is separated in a separation unit to obtain hydrogen-rich gas 4, dry gas 101, light naphtha 102, low aromatic solvent oil 103, and unconverted oil 104.

[0087] Wherein, a hydrogenation protective agent and a hydrodeoxygenation catalyst are placed in the hydrodeoxygenation reactor in sequence from top to bottom.

[0088] The liquid phase stream 7 generated by the hydrocracking reaction flows out from the bottom of the hydrocracking reaction distillation tower, and the gas phase stream 8 from the gas phase channel 321 flows out from the top of the hydrocracking reaction 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, light naphtha 102, low aromatic solvent oil 103 and unconverted oil 104. The unconverted oil 104 is fully or partially recycled to the reaction distillation tower 1 to continue the hydrocracking reaction, and can also be drawn out of the device as a product.

[0089] In the hydrogenation reaction unit of the present invention, the liquid phase feed and the gas phase feed are subjected to catalytic reaction in the catalyst bed, and the gas phase product and the unreacted gas phase feed rise and leave the reaction system through the gas phase channel 13. The gas phase product 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, thereby avoiding the 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.

[0090] 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.

[0091] 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.

[0092] Example 1

[0093] The above-mentioned device and process of the present invention are used to carry out hydrogenation reaction and fractionation treatment on raw biological oil. The raw material oil is soybean oil, and its main properties are shown in Table 1.

[0094] 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 raw materials undergo 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-31 / FZC-41. The above catalysts are all produced by Sinopec Dalian Petrochemical Research Institute. The specific operating process conditions are shown in Table 2.

[0095] The reactive distillation tower is provided with four trays, and the upper part of all trays is filled with fillers, and the lower part is filled with hydrocracking catalyst. The fillers are 5 mm corundum Raschig rings, and the hydrocracking catalyst is FC-46 catalyst 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. After the liquid is reduced in pressure, it 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] The main properties of the product are shown in Table 3.

[0098] Example 2-3

[0099] The raw material oil is soybean oil, and some reaction conditions are changed, as shown in Table 2, and the other conditions are the same as those in Example 1. The main properties of the product are shown in Table 3.

[0100] Example 4

[0101] The raw material oil was replaced with palm oil, the main properties of which are shown in Table 1, and the specific reaction conditions are shown in Table 2. The main properties of the product are shown in Table 3.

[0102] Example 5

[0103] The raw material oil is soybean oil. Based on a single-pass conversion rate of 65 wt% at >165°C, the active temperatures of the two adjacent hydrocracking catalysts are 5°C lower from top to bottom. Among them, the FC-46 catalyst is selected for the second layer.

[0104] Comparative Example 1

[0105] A conventional two-stage hydrogenation process, i.e., hydrodeoxygenation + hydrocracking process, is adopted. Both the hydrodeoxygenation and cracking reactors adopt a reaction process in which the raw material and hydrogen flow in parallel from top to bottom. The soybean oil is first saturated with olefins in the hydrodeoxygenation reactor. After the hydrodeoxygenation and hydrodecontamination reactions, the hydrodeoxygenation logistics is dehydrated, wherein the refined high-pressure gas refluxes into the hydrorefining reactor, and the hydrodeoxygenation liquid enters the hydrocracking reactor for hydrocracking reaction. After the reaction is completed, it enters the second separation and fractionation system, and gas, naphtha, low aromatic solvent oil and unconverted oil are separated. The unconverted oil after separation is all refluxed into the hydrocracking reactor for hydrocracking reaction again. The catalyst and process conditions used are the same as those in Example 1.

[0106] Table 1 Main properties of crude oil

[0107]

[0108] Table 2 Example process conditions

[0109]

[0110] Table 3 Product distribution and properties

[0111]

[0112] From the results in Table 3, it can be seen that by adopting the method of the present invention, bio-oil can be hydrogenated to produce solvent oil, and the aromatic content is very low, which is a high-quality chemical product.

Claims

1. A device for producing low aromatic solvent oil by hydrogenation of bio-oil, characterized in that: include: (1) a hydrodeoxygenation reactor, which is used to perform a hydrodeoxygenation reaction on the raw bio-oil; (2) a hydrogenation reaction distillation tower, which receives the liquid phase product after the hydrogenation and deoxygenation of the bio-oil as a liquid phase feed, and the liquid phase feed flows from top to bottom through each catalyst bed 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; (3) 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 catalytic reaction zone, the catalytic reaction zone is a 2-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 on the top of each catalyst bed; Each catalyst bed and inclined partition is provided with a gas phase channel; The hydrogenation reaction distillation column comprises a liquid-phase feed subunit, which is arranged above the inclined partition of the topmost catalyst bed, through which the liquid-phase feed is guided to the catalyst bed; 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 layer; The hydrogenation reaction distillation tower 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; 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; Each of the catalyst beds is provided with an overflow weir arranged on one side close to the gas phase channel and a liquid sealing baffle arranged on the upper part of the overflow weir for isolating the gas phase feed from the gas phase product.

2. The device according to claim 1, characterized in that The system also includes (4) a gas circulation unit, which is used to selectively concentrate the hydrogen-rich gas obtained from the separation unit and then circulate it back to the hydrodeoxygenation reactor and the hydrogenation reaction distillation tower.

3. The device according to claim 1, characterized in that The invention also comprises (5) a liquid circulation unit, which is used to recycle all or part of the liquid product obtained from the hydrogenation reaction distillation tower back to the hydrodeoxygenation reactor or the hydrogenation reaction distillation tower.

4. 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.

5. The device according to claim 1, characterized in that The liquid phase feed subunit includes: a liquid phase feed pipe, which extends in 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 pipe wall of the liquid phase distribution pipe is provided with multiple channels for evenly distributing the liquid phase feed to all directions of the annular downcomer.

6. 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.

7. The device according to claim 6, 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.

8. The device according to claim 1, characterized in that The uppermost end of the gas phase channel is connected to the outlet pipeline of the hydrogenation reactor, and the lowermost end is close to the bottom of the hydrogenation reaction distillation tower. The port is immersed in the liquid phase product. A liquid level monitoring unit is arranged at the bottom of the hydrogenation reaction distillation tower.

9. A method for producing low aromatic solvent oil by hydrogenating bio-oil, characterized in that: Using the device of claim 1, the method comprises the following steps: (1) The bio-oil raw material is mixed with hydrogen and enters the hydrodeoxygenation reactor, where it passes through the hydrodeoxygenation catalyst bed under hydrodeoxygenation conditions; (2) After dehydration, the hydrodeoxygenation stream obtained in step (1) is used as a liquid phase feed to enter a hydrogenation reaction distillation tower, and flows from top to bottom through each catalyst bed of the hydrogenation reaction distillation tower, and contacts and reacts with the preheated hydrogen entering from the bottom of each catalyst bed from bottom to top on the catalyst of each bed; (3) The gas phase product obtained by the reaction of each catalyst bed enters the gas phase channel, flows upward in the gas phase channel and flows out of the hydrogenation reaction distillation tower; the liquid phase product obtained by the reaction of each catalyst bed sequentially enters the next catalyst bed layer for reaction, 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.

10. The method according to claim 9, characterized in that The method further comprises step (5): the hydrogen-rich gas obtained in step (4) is optionally concentrated and then returned to step (1) and step (2).

11. The method according to claim 9, characterized in that The method further comprises step (6): the liquid product obtained in step (3) is recycled back to step (1) or step (2).

12. The method according to claim 9, characterized in that The biological oil includes at least one of vegetable oil, animal oil, acidified oil, and waste cooking oil; the 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 seed oil and rice bran oil; the animal oil includes one or more of tallow, lard, mutton oil and fish oil.

13. The method according to claim 9, characterized in that The hydrodeoxygenation conditions in step (1) are as follows: reaction pressure of 3.0 MPa to 20.0 MPa, hydrogen to oil volume ratio of 200:1 to 3000:1, volume space velocity of 0.1 h -1 ~6.0h -1 The average reaction temperature is 180℃~450℃.

14. The method according to claim 9, characterized in that In step (1), a hydrogenation protection catalyst is loaded upstream of the hydrodeoxygenation catalyst bed, and the volume ratio of the hydrogenation protection agent to the hydrodeoxygenation agent is 15:85-70:

30.

15. The method according to claim 9, characterized in that Solid fillers are placed on the top of each catalyst bed in the hydrogenation reaction distillation tower, and hydrocracking catalyst is loaded on the bottom.

16. The method according to claim 9, characterized in that The operating conditions of the hydrogenation reaction distillation tower are: reaction temperature of 260°C to 450°C, reaction pressure of 3MPa to 20MPa, hydrogen-to-oil volume ratio of 100 to 2000, liquid hourly volume space velocity of 0.1h -1 ~10.0h -1 .

17. The method according to claim 9, characterized in that After the hydrodeoxygenation in step (1), the hydrodeoxygenation product is subjected to gas-liquid separation, and the separated liquid fraction enters the hydrogenation reaction distillation tower for further hydrogenation treatment.

18. The method according to claim 9, characterized in that The hydrodeoxygenation catalyst is selected from a sulfided catalyst or a reduced catalyst, and the hydrocracking catalyst is selected from a sulfided catalyst or a reduced catalyst.

19. The method according to claim 18, characterized in that Step (1) uses a sulfided hydrodeoxygenation catalyst, step (2) uses a sulfided hydrocracking catalyst, the hydrodeoxygenation reactor and the hydroreaction distillation tower share a set of circulating hydrogen systems, or use two sets of circulating hydrogen systems; or, step (1) uses a sulfided hydrodeoxygenation catalyst, step (2) uses a reduced hydrocracking catalyst, the hydrodeoxygenation reactor and the hydroreaction distillation tower use two sets of circulating hydrogen systems; or, step (1) uses a reduced hydrodeoxygenation catalyst, step (2) uses a sulfided hydrocracking catalyst, the hydrodeoxygenation reactor and the hydroreaction distillation tower use two sets of circulating hydrogen systems; or, step (1) uses a reduced hydrodeoxygenation catalyst, step (2) uses a sulfided hydrocracking catalyst, the hydrodeoxygenation reactor and the hydroreaction distillation tower use two sets of circulating hydrogen systems; or, step (1) uses a reduced hydrodeoxygenation catalyst, step (2) uses a reduced hydrocracking catalyst, the hydrodeoxygenation reactor and the hydroreaction distillation tower share a set of circulating hydrogen systems, or use two sets of circulating hydrogen systems.

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