CN116064128A

CN116064128A - Hydrogenation catalytic distillation device and method for Fischer-Tropsch synthetic oil

Info

Publication number: CN116064128A
Application number: CN202111268104.7A
Authority: CN
Inventors: 马蕊英; 赵玉琢; 郭兵兵; 王晶晶; 宣根海
Original assignee: China Petroleum and Chemical Corp; Sinopec Dalian Research Institute of Petroleum and Petrochemicals
Current assignee: Sinopec Dalian Petrochemical Research Institute Co ltd; China Petroleum and Chemical Corp
Priority date: 2021-10-29
Filing date: 2021-10-29
Publication date: 2023-05-05

Abstract

The invention discloses a Fischer-Tropsch synthesis oil hydrogenation catalytic distillation device and a Fischer-Tropsch synthesis oil hydrogenation catalytic distillation method. The device comprises: the hydrogenation refining reactor and the catalytic reaction distillation tower comprise at least one catalytic reaction zone, wherein the catalytic reaction zone has a 2-n layer tower plate structure, and a catalyst bed layer is filled on each layer of tower plate; an inclined baffle plate is arranged at the upper part of each catalyst bed layer; a gas phase passage is provided through each catalyst bed and the inclined partition. By adopting the device, the gas-phase product generated by the hydrocracking reaction can timely leave the cracking reaction zone and cannot enter the upper catalyst bed again, so that the gas-phase product is prevented from occupying the channels of the cracking catalyst, the secondary cracking and gasification of the target product are effectively prevented, and the reaction selectivity and the yield of the target product are improved.

Description

Hydrogenation catalytic distillation device and method for Fischer-Tropsch synthetic oil

Technical Field

The invention belongs to the field of hydrogenation reaction, and particularly relates to a hydrogenation catalytic distillation device and method for Fischer-Tropsch synthetic oil.

Background

With the rapid development of transportation industry, natural petroleum resources are increasingly being supplied, and various countries in the world are actively exploring and developing alternative products of petroleum resources. The energy structure of China is characterized in that 'rich coal, lack of oil and less gas', synthesis gas is produced by taking coal as raw material, and the synthesis gas is produced into liquid hydrocarbon by Fischer-Tropsch synthesis, which is one of important ways for solving the current demand of fuel supply.

The Fischer-Tropsch synthesis reaction is a technological process using hydrogen and carbon monoxide as raw materials and using the Fischer-Tropsch synthesis reaction to generate liquid fuel as main material under the action of a catalyst, and the raw materials are mainly coal and natural gas. The Fischer-Tropsch synthetic oil mainly comprises normal alkane, alkene and a certain amount of oxygen-containing compounds, has extremely low content of non-ideal components such as sulfur, nitrogen, aromatic hydrocarbon and the like, and has great difference from conventional mineral petroleum in terms of hydrocarbon composition and main properties. The naphtha fraction in the Fischer-Tropsch synthetic oil is basically free of sulfur and nitrogen, the octane number is very low (the olefin and alkane are mostly linear), and the olefin content and the oxygen content are the highest; the diesel fraction has extremely low sulfur, nitrogen and aromatic hydrocarbon contents, high cetane number, intermediate olefin content and oxygen content and the lowest olefin content and oxygen content of heavy oil. Accordingly, fischer-Tropsch synthesis oil must be subjected to a corresponding hydrotreatment to obtain a petrochemical feedstock or transportation fuel that meets the specifications for use.

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

The conventional fixed bed hydrogenation mode is still adopted in the heavy oil hydrocracking reaction in the Fischer-Tropsch synthesis oil at present, raw oil and hydrogen flow into a reactor together in parallel, and materials flow downwards while reacting in the bed propelling process. In the whole reaction process, gas-phase reaction products cannot leave the reactor in time, so that the gas-phase reaction products occupy the pore channels of the catalyst, and the cracking reaction occurs again, so that the selectivity of the reaction and the yield of target products are reduced. In addition, as the normal paraffin content in the heavy oil fraction of Fischer-Tropsch synthetic oil is up to more than 80%, olefin saturation occurs quickly in the olefin product after normal paraffin cracking, the olefin saturation heat release amount is large, the whole reaction is much larger than that of the traditional petroleum-based crude oil, the hydrogen consumption is increased, the bed temperature is difficult to control, and the risk of 'flying temperature' is easily generated.

Disclosure of Invention

Aiming at the defects of the prior art, the invention aims to provide a Fischer-Tropsch synthesis oil hydrogenation device and a Fischer-Tropsch synthesis oil hydrogenation method. The method overcomes the defect that the gas-phase reaction product cannot be effectively separated from the reaction zone in time in the prior art, improves the yield and selectivity of the target product, prevents a great deal of olefin saturation, and eliminates the risk of bed temperature flying.

In a first aspect, the present invention provides a fischer-tropsch synthesis oil hydrogenation catalytic distillation apparatus comprising:

(1) The hydrofining reactor is used for carrying out hydrofining reaction on the raw material Fischer-Tropsch synthetic oil;

(2) The catalytic reaction distillation tower comprises at least one catalytic reaction zone, wherein the catalytic reaction zone is of a 2-n-layer tower plate structure, and each layer of tower plate is filled with a catalyst bed layer, wherein n is an integer greater than 2; an inclined baffle, preferably an umbrella-shaped baffle, is arranged at the upper part of each catalyst bed layer; a gas phase passage is provided through each catalyst bed and the inclined partition.

Further, in the above technical scheme, the hydrofining reaction includes olefin saturation, hydrodeoxygenation, hydrodemetallization, and the like.

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

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

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

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

Further, in the above technical scheme, the catalyst bed is provided with: an overflow weir provided on one side close to the gas phase passage; and the liquid sealing baffle is arranged at the upper part of the overflow weir and is used for isolating the gas-phase feeding material from the gas-phase product.

Further, in the above technical scheme, the liquid seal baffle includes: a horizontal part which is in an annular flat plate shape and is positioned above the overflow weir; and the vertical part is cylindrical, the vertical part and the horizontal part are integrally formed, and the lower end of the vertical part is spaced from the bottom of the catalyst bed by a certain distance.

Further, in the above technical solution, the liquid phase feeding subunit further includes: a liquid phase feed tube extending in a radial direction of the catalyst bed; and the liquid phase distributing pipe is annular and is orthogonal or tangentially intersected with the liquid phase feeding pipe, and the pipe wall of the liquid phase distributing pipe is provided with a plurality of pore channels for uniformly distributing liquid phase feeding materials to all directions of the annular downcomer.

Further, in the above technical solution, the gas phase feeding subunit further includes: a gas phase feed tube extending in a radial direction of the catalyst bed; the gas distribution pipe is annular or multi-layer concentric annular, the gas distribution pipe is orthogonal or tangential to the gas inlet pipe, and a plurality of pore channels are arranged on the wall surface of the gas distribution pipe and are used for uniformly distributing gas phase feeding materials to all directions at the bottom of the catalyst bed.

Further, in the above technical solution, the gas phase feeding subunit may further include: the gas phase distribution plate is positioned at the bottom of the catalyst bed and is in a disc shape as a whole, and a plurality of holes are uniformly distributed on the gas phase distribution plate.

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

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

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

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

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

Further, in the above technical solution, a separation zone is provided at the upper part of the catalytic reaction distillation column; the number ratio of the separation zone to the catalyst bed layer of the catalytic reaction zone is 1:1-1:n-1. A packing or hydrocracking catalyst, preferably a solid packing, may be selectively placed on the catalyst bed in the separation zone. The catalyst bed layer in the catalytic reaction zone is filled with a hydrocracking catalyst.

Further, in the above technical scheme, a fractionating tower may be optionally provided between the hydrofining reactor and the catalytic reaction distillation tower.

The second aspect of the invention provides a Fischer-Tropsch synthesis oil hydrogenation method, which comprises the following steps:

(1) The Fischer-Tropsch synthetic oil and hydrogen are mixed and then enter a hydrofining reactor, and hydrofining reaction is carried out under the action of a hydrofining catalyst;

(2) The material flow obtained by hydrofining in the step (1) enters a catalytic reaction distillation tower, and goes down in the tower to enter a plurality of catalyst beds in sequence to contact with hydrogen entering the catalyst beds upwards for reaction; the gas phase product obtained by each catalyst bed reaction enters a gas phase channel, and flows out of the catalytic reaction distillation tower upwards in the gas phase channel; and the liquid phase product obtained by each catalyst bed layer reaction sequentially enters the next catalyst bed layer for reaction, and finally the obtained liquid product flows downwards out of the catalytic reaction distillation tower.

In the technical scheme, after the gas phase product flowing out of the top of the catalytic reaction distillation tower is mixed with the liquid phase product flowing out of the bottom of the catalytic reaction distillation tower, high-pressure hydrogen-rich gas, naphtha, diesel oil and heavy oil are obtained through separation and fractionation, and all or part of the heavy oil is returned to the catalytic reaction distillation tower for hydrocracking reaction.

Furthermore, in the technical scheme, the obtained high-pressure hydrogen-rich gas is recycled.

In the technical scheme, the hydrofining material flows into the catalytic reaction distillation tower, goes down in the tower and goes into a separation area to be separated into light components and heavy components, the heavy components enter a catalyst bed layer downwards to react, and the light components are directly mixed with gas-phase products and flow out of the catalytic reaction distillation tower from the top.

Further, in the above technical scheme, the light component in step (2) generally refers to C1-C4 gas, naphtha and kerosene fractions, and the heavy component generally refers to diesel oil and wax oil fractions.

Further, in the above technical scheme, the fischer-tropsch synthesis oil includes high temperature fischer-tropsch synthesis full distillate oil and low temperature fischer-tropsch synthesis full distillate oil, preferably low temperature fischer-tropsch synthesis full distillate oil. The Fischer-Tropsch synthetic oil has the following properties: density of 0.6g/cm ³ ～1.0g/cm ³ Preferably 0.7g/cm3 to 0.95g/cm ³ The method comprises the steps of carrying out a first treatment on the surface of the The final distillation point is 650 ℃ to 750 ℃, preferably 680 ℃ to 720 ℃.

Further, in the above technical scheme, the hydrogenation reactor in the step (1) is a conventional fixed bed hydrogenation reactor, which contains a hydrofining catalyst bed layer. The hydrogenation protecting agent is preferably placed at the upper part of the hydrogenation refining catalyst bed layer, the hydrogenation refining catalyst is placed at the lower part of the hydrogenation refining catalyst bed layer, and the volume ratio of the hydrogenation protecting agent to the hydrogenation refining catalyst is 0-3:10, preferably 1:20-1:5. The hydrogenation protecting agent and the hydrofining catalyst are conventional protecting agent and refining catalyst. Generally, alumina or silicon-containing alumina is used as a carrier, and metals of a VIB group and a VIII group are used as active components, for example, one, two or more of W, mo, co, ni are selected, and the hydrogenation protecting agent has larger pore diameter than a hydrofining catalyst.

Further, in the above technical scheme, the operation conditions of the hydrofining process of the Fischer-Tropsch synthetic oil in the step (1) are as follows: the reaction temperature is 200-400 ℃, the reaction pressure is 3-18 MPa, the hydrogen-oil volume ratio is 100-1000, and the liquid hourly space velocity is 0.5h ^-1 ～10.0h ^-1 . Preferred operating conditions are: the reaction temperature is 250-390 ℃, the reaction pressure is 4-15 MPa, the hydrogen-oil volume ratio is 200-1000, and the liquid hourly space velocity is 0.5h ^-1 ～8.0h ^-1 。

In the above technical scheme, the number of the catalyst beds in the step (2) is n, the upper layer can be selectively placed with a filler or a hydrocracking catalyst, preferably with a solid filler, the lower layer is filled with the hydrocracking catalyst, and the layer ratio of the solid filler to the hydrocracking catalyst is preferably 1:1-1:n-1. The filler is in a conventional form in the field, for example, one or more random fillers such as pall rings, raschig rings, saddle-shaped, open pore ring types, semi-rings, ladder rings, double arcs, halfpace rings, conjugated rings, flat rings, flower rings and the like can be selected, and the filler can also be selected from metal or ceramic corrugated fillers. The hydrocracking catalyst generally comprises an active component and a carrier, wherein the carrier component comprises one or more of alumina, silica-containing alumina and molecular sieves, preferably molecular sieves, and the molecular sieves can be Y-type molecular sieves; the active component is one or more of VIB group and VIII group metals, wherein the VIB group metals are Mo and/or W, and the VIII group metals are Co and/or Ni. The hydrocracking catalyst shape may be any conventional existing hydrocracking catalyst shape, preferably porous, shaped and/or honeycomb catalysts. The pore diameter of the porous catalyst is 1-50 mm, preferably 4-20 mm; the average particle diameter of the shaped catalyst is 2-50 mm, preferably 4-30 mm; the diameter or the side length of the honeycomb catalyst holes is 1-50 mm, preferably 3-15 mm; the void fraction of the catalyst bed is recommended to be 15-85%, preferably 20-75%.

Further, in the above technical scheme, the operation conditions of the catalytic distillation reactor in the step (2) are as follows: the reaction temperature is 260-450 ℃, the reaction pressure is 3-20 MPa, the hydrogen-oil volume ratio is 100-2000, and the liquid hourly space velocity is 0.1h ^-1 ～10.0h ^-1 . Preferred operating conditions are: the reaction temperature is 300-450 ℃, the reaction pressure is 4-15 MPa, the hydrogen-oil volume ratio is 100-1500, and the liquid hourly space velocity is 0.5h ^-1 ～10.0h ^-1 。

In the technical scheme, the hydrofined product can be selectively fractionated after the hydrofined product in the step (1), and naphtha, diesel oil and heavy oil fractions are obtained after the hydrofined product is fractionated, wherein the heavy oil fractions enter a catalytic reaction distillation tower for further hydrotreating.

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

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

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

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

The heavy component of Fischer-Tropsch synthetic oil contains more than 80% of normal paraffins, and the gas phase product generated after the normal paraffins are cracked is rich in a large amount of olefins, so that the olefins can timely leave the distillation tower, the saturation of the olefins is effectively prevented, hydrogen is consumed, a large amount of heat is generated, the hydrogen consumption and the reaction temperature rise are reduced, and the risk of 'flying temperature' is prevented. In addition, the gas phase product rich in olefin belongs to high-quality chemical raw materials.

Further, the light components of the hydrofinishing stream of the present invention can be separated directly from the catalytic reactive distillation column via a gas phase passage as they pass through a separation zone within the catalytic reactive distillation column. Even if a small amount of light components are entrained into the hydrocracking reaction zone, the heavy components can be preferentially adsorbed on the surface of the cracking catalyst to generate cracking reaction due to the large polarity of the large molecules in the heavy components, the light components are influenced by competitive adsorption, and the light components can be further carried out of the distillation tower through a gas phase channel by means of the stripping action of hydrogen, so that the light components are hardly cracked into smaller molecules on the cracking catalyst, and the load of the cracking reaction zone is increased. Therefore, the invention can realize the separation of light and heavy components of hydrofining material flow by the design of the catalytic reaction distillation tower without specially designing a set of fractionating system after the refining reactor, thereby simplifying the reaction flow.

Furthermore, the umbrella-shaped partition plate can separate gas phase feeding and product gas between adjacent bed layers, and plays a role in guiding liquid phase and gas phase. The liquid sealing baffle can effectively isolate the gas phase feeding material from the gas phase product. The arrangement of the multi-layer concentric annular gas phase distributing pipe can keep the distribution of the gas phase feeding material uniform to the maximum extent.

Drawings

FIG. 1 is a schematic flow diagram of one embodiment of the Fischer-Tropsch synthesis oil hydrogenation catalytic distillation of the present invention.

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

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

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

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

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

FIG. 7 is a top view of a vapor distribution tray in a catalytic reactive distillation column of the present invention.

FIG. 8 is a top view of a catalyst support tray in a catalytic reaction distillation column according to the present invention.

The main reference numerals illustrate:

1-catalytic reaction distillation tower, 2-Fischer-Tropsch synthetic oil, 3-new hydrogen, 4-high-pressure hydrogen-rich gas, 5-hydrofining reactor, 6-hydrofining stream, 7-liquid phase stream, 8-gas phase stream, 9-high-pressure separator, 10-fractionating tower, 101-gas, 102-naphtha, 103-diesel oil and 104-heavy oil.

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

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

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

Detailed Description

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

FIG. 1 is a schematic diagram of the hydrogenation catalytic distillation process flow of Fischer-Tropsch synthesis oil. As shown in fig. 1 and 2, the fischer-tropsch synthesis oil 2, the fresh hydrogen 3 and the circulating pressurized hydrogen-rich gas 4 are mixed and then enter a hydrofining reactor 5, and a hydrogenation protecting agent and a hydrofining catalyst are sequentially placed in the hydrofining reactor from top to bottom. The hydrofining stream 6 enters a liquid phase feed pipe 21 of the catalytic reaction distillation column 1, is uniformly distributed through a liquid phase distribution pipe 22 and then flows downwards along a packing or catalyst bed. When the hydrofining stream 6 passes through the packing of the separation zone, light components flow upwards through the gas phase channel 321, heavy components continue to flow downwards, and when entering the catalyst bed of the catalytic reaction zone, after reacting with hydrogen which enters the catalyst bed upwards from the gas phase feeding pipe 31, gas phase products directly enter the gas phase channel 321, and heavy components continue to flow downwards along the catalyst bed and react.

The liquid phase material flow 7 generated by the hydrocracking reaction flows out from the bottom of the catalytic reaction distillation column, the gas phase material flow 8 from the gas phase channel 321 flows out from the top of the catalytic reaction distillation column, the gas phase material flow 8 and the gas phase material flow are mixed and enter the high-pressure separator 9 for gas-liquid separation, the separated high-pressure hydrogen-rich gas 4 and the new hydrogen 3 are mixed and then used as circulating hydrogen, the separated liquid enters the fractionating column 10 for fractionation to obtain gas 101, naphtha 102, diesel 103 and heavy oil 104, and the heavy oil 104 is completely or partially circulated back to the catalytic reaction distillation column 1 for continuous hydrocracking reaction, and can also be led out of the device for producing lubricating oil base oil raw materials.

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

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

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

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

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

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

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

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

Example 1

The device and the flow are adopted to carry out catalytic distillation treatment on the raw material Fischer-Tropsch synthetic oil. The properties of the raw Fischer-Tropsch oil are shown in Table 1.

The upper part in the hydrofining reactor is provided with a hydrogenation protecting agent, the lower part is filled with a hydrofining catalyst, and the volume ratio of the hydrogenation protecting agent to the hydrofining catalyst is 1:10. The raw materials are subjected to olefin saturation, deoxidation and impurity removal reactions in a hydrofining reactor. The hydrogenation protecting agent is FZC-105, the hydrofining catalyst is FHUDS-5, and both are produced by China petrochemical industry institute of great company petrochemical industry. The specific operating process conditions are shown in Table 2.

Five layers of tower plates are arranged in the catalytic reaction distillation tower, the upper part of the first layer of tower plates is filled with filler, the lower part is filled with hydrocracking catalyst, and the rest four layers of tower plates are filled with hydrocracking catalyst. The filler is 5mm corundum Raschig ring produced by Baisheng chemical filler Co., ltd, the hydrocracking catalyst is FC-14 produced by China petrochemical industry institute of great company petrochemical industry, and the specific operation process conditions are shown in Table 2.

The whole fraction of the catalytic reaction distillation tower is subjected to high-pressure reflux of hydrogen-rich gas 4 through a high-pressure separator, and the rest products enter a fractionating tower for fractionation, so that gas, naphtha, diesel oil fraction and heavy oil fraction are obtained, wherein the heavy oil fraction is partially refluxed and partially thrown outwards. The high pressure separator was operated at 85 ℃, the fractionation column was operated at 0.1MPa, the column top temperature was 70 ℃, and the column bottom temperature was 390 ℃.

The product distribution and properties are shown in Table 3.

Example 2

The difference with example 1 is only that a fractionating tower is added between the hydrofining reactor and the catalytic reaction distillation tower. Fractionating the hydrofined material flow in a fractionating tower to obtain naphtha, diesel oil and heavy oil fraction, wherein the heavy oil fraction enters a catalytic reaction distillation tower to continue hydrocracking reaction. The operation condition of the added fractionating tower is 0.1MPa, the temperature of the top of the tower is 60 ℃, and the temperature of the bottom of the tower is 400 ℃.

Examples 3 to 4

The reaction conditions were changed as shown in Table 2, and the other conditions were the same as in example 1.

Comparative example 1

Adopts the conventional two-stage hydrogenation method, namely refining and cracking process. Both the refining reactor and the cracking reactor adopt a reaction process that raw materials and hydrogen flow in parallel flow from top to bottom. The Fischer-Tropsch synthesis oil is subjected to olefin saturation in a hydrofining reactor, after deoxidation and impurity removal reaction, hydrofining material flows enter a first separation and fractionation system to obtain refined high-pressure gas, naphtha, diesel oil and heavy oil, wherein the refined high-pressure gas flows back into the hydrofining reactor, the heavy oil enters a hydrocracking reactor to carry out hydrocracking reaction, after the reaction is finished, the heavy oil enters a second separation and fractionation system to obtain gas, naphtha and diesel oil fractions through separation, and all the separated uncracked heavy components flow back into the hydrocracking reactor to carry out hydrocracking reaction again. The catalyst and process conditions used were the same as in example 1.

TABLE 1 Main Properties of raw oil

Project	Data
		Density, g/cm ³	0.807
Distillation range, DEG C
		Initial point/10%	68/203
30％/50％	314/372
		70％/90％	461/558
95％	649
		Sulfur content, vol%	3.9
Nitrogen content, μg/g	3.2
		Oxygen content, wt%	0.79

Table 2 example process conditions

TABLE 3 product distribution and Properties

	Example 1	Example 2	Example 3	Example 4	Comparative example 1
						Naphtha fraction
Distillation range, DEG C	65～160	65～160	65～160	65～160	65～160
						Yield, wt%	43.1	36.7	40.2	38.5	16.8
Density, g/cm ³	0.703	0.705	0.712	0.710	0.712
						Composition, wt%
N-alkanes	88.21	87.18	85.47	88.15	84.50
						Isoparaffin(s)	10.19	11.60	12.36	10.22	12.21
Diesel oil fraction
						Distillation range, DEG C	160～370	160～370	160～370	160～370	160～370
Yield, wt%	55.2	62.8	58.1	60.7	78.1
						Density, g/cm ³	0.781	0.773	0.765	0.775	0.793
Condensation point, DEG C	-3	-4	-5	-5	-2
						Cetane number	83	83	84	84	78

As shown in Table 3, the yield of naphtha was significantly higher than that of comparative example 1, and the naphtha fraction consisted mainly of alkane, which was a good raw material for producing ethylene and propylene by steam cracking.

Example 5

The laboratory performed simulated calculations of the bed reaction temperature profiles of the two-stage hydrogenation reactors (hydrocracking reactors) of examples and comparative examples using ansys version 19.0 software. The simulation conditions were as actual data inputs for examples and comparative examples. Simulation results show that the center temperature of the traditional fixed bed is highest, and the temperature change is normally distributed from the inlet end to the outlet end, so that the temperature of the reactor bed is relatively uniform. The simulated temperature rise change of the bed is shown in Table 4.

Table 4 bed simulated temperature rise variation

Bed temperature point	Example 1	Example 2	Example 3	Example 4	Comparative example 1
						Maximum radial temperature difference, DEG C	1.6	1.3	1.5	2.0	26.1
Average temperature, DEG C	386.0	370.9	376.1	401.2	402.7

As can be seen from the results of Table 4, the temperature difference of the catalyst beds in examples 1 to 3 according to the present invention was significantly lower than that of comparative example 1, and the temperature difference was reduced from 26.1℃to less than 2.0℃in the conventional fixed bed, and the difference between the average temperature and the control temperature of the examples was minimized, indicating that the reactor according to the present invention has eliminated the overheating phenomenon of the hydrocracking reaction.

Claims

1. A fischer-tropsch synthesis oil hydrocatalytic distillation apparatus, said apparatus comprising:

2. The apparatus of claim 1, wherein the catalytic reactive distillation column comprises a liquid phase feed subunit disposed above an inclined partition of the topmost catalyst bed through which liquid phase feed is directed to the catalyst bed.

3. The apparatus of claim 1 or 2, wherein the inclined partition ends with an annular downcomer having a bottom spaced from the bottom of the catalyst bed such that the liquid phase feed enters the catalyst bed in a radial direction.

4. The apparatus of claim 1, wherein the catalytic reactive distillation column comprises a gas phase feed subunit disposed between a catalyst bed of an upper layer and an inclined partition of a lower layer.

5. The apparatus of claim 1 wherein the lower ends of the gas phase channels through the previous catalyst bed are connected to inclined baffles on the next catalyst bed.

6. The apparatus according to claim 1, wherein the catalyst bed is provided with: an overflow weir provided on one side close to the gas phase passage; and the liquid sealing baffle is arranged at the upper part of the overflow weir and is used for isolating the gas-phase feeding material from the gas-phase product.

7. The apparatus of claim 6, wherein the liquid seal barrier comprises: a horizontal part which is in an annular flat plate shape and is positioned above the overflow weir; and the vertical part is cylindrical, the vertical part and the horizontal part are integrally formed, and the lower end of the vertical part is spaced from the bottom of the catalyst bed by a certain distance.

8. The apparatus of claim 2, wherein the liquid phase feed subunit comprises: a liquid phase feed tube extending in a radial direction of the catalyst bed; and the liquid phase distributing pipe is annular and is orthogonal or tangentially intersected with the liquid phase feeding pipe, and the pipe wall of the liquid phase distributing pipe is provided with a plurality of pore channels for uniformly distributing liquid phase feeding materials to all directions of the annular downcomer.

9. The apparatus of claim 4, wherein the gas phase feed subunit comprises: a gas phase feed tube extending in a radial direction of the catalyst bed; the gas distribution pipe is annular or multi-layer concentric annular, the gas distribution pipe is orthogonal or tangential to the gas inlet pipe, and a plurality of pore channels are arranged on the wall surface of the gas distribution pipe and are used for uniformly distributing gas phase feeding materials to all directions at the bottom of the catalyst bed.

10. The apparatus according to claim 4 or 9, wherein the gas phase feed subunit comprises: the gas phase distribution plate is positioned at the bottom of the catalyst bed and is in a disc shape as a whole, and a plurality of holes are uniformly distributed on the gas phase distribution plate.

11. The apparatus according to claim 1, wherein the gas phase channels are located in the middle of the catalytic reaction unit and run through all catalyst beds from bottom to top.

12. The apparatus of claim 6 wherein the upper edge of the weir is 10 to 100mm above the upper surface of the catalyst bed.

13. The apparatus according to claim 1, wherein the catalytic reaction distillation column is provided with a separation zone in the upper part; the number ratio of the separation zone to the catalyst bed layer of the catalytic reaction zone is 1:1-1:n-1; a packing or hydrocracking catalyst, preferably a solid packing, may be selectively placed on the catalyst bed of the separation zone; the catalyst bed layer in the catalytic reaction zone is filled with a hydrocracking catalyst.

14. A process for the hydrogenation of fischer-tropsch synthesis oil using the apparatus of any one of claims 1 to 13, comprising the steps of:

15. The process of claim 14 wherein the hydrofinishing stream enters a catalytic reactive distillation column, proceeds down the column first to a separation zone, separates into light components and heavy components, the heavy components enter a catalyst bed downwardly for reaction, the light components are mixed directly with the vapor phase product and exit the catalytic reactive distillation column from the top.

16. The process of claim 14, wherein the operating conditions of the fischer-tropsch synthesis oil hydrofinishing process in step (1) are as follows: the reaction temperature is 200-400 ℃, the reaction pressure is 3-18 MPa, the hydrogen-oil volume ratio is 100-1000, and the liquid hourly space velocity is 0.5h ^-1 ～10.0h ^-1 The method comprises the steps of carrying out a first treatment on the surface of the Preferred operating conditions are: the reaction temperature is 250-390 ℃, the reaction pressure is 4-15 MPa, the hydrogen-oil volume ratio is 200-1000, and the liquid hourly space velocity is 0.5h ^-1 ～8.0h ^-1 。

17. The process according to claim 14, wherein the catalytic distillation reactor in step (2) is operated under the following conditions: the reaction temperature is 260-450 ℃, the reaction pressure is 3-20 MPa, the hydrogen-oil volume ratio is 100-2000, and the liquid hourly space velocity is 0.1h ^-1 ～10.0h ^-1 The method comprises the steps of carrying out a first treatment on the surface of the Preferred operating conditions are: the reaction temperature is 300-450 ℃, the reaction pressure is 4-15 MPa, the hydrogen-oil volume ratio is 100-1500, and the liquid hourly space velocity is 0.5h ^-1 ～10.0h ^-1 。