CN116064131A - A Fischer-Tropsch synthetic oil hydrotreating device and method - Google Patents

Abstract

The invention discloses a Fischer-Tropsch synthesis oil hydrotreatment device and a Fischer-Tropsch synthesis oil hydrotreatment method. The device comprises a separator, a reaction cavity and a heavy oil bin which are sequentially communicated from top to bottom; the reaction chamber is internally provided with: a hydrofining catalyst bed, a separation cavity, a hydrocracking catalyst bed and a hydrogen distribution cavity. According to the invention, the characteristics of large difference of hydrocarbon composition, olefin and oxygen content of each fraction of Fischer-Tropsch synthesis oil are fully considered, the Fischer-Tropsch synthesis oil with wide distillation range distribution is divided into a lighter part and a heavier part by arranging the separation cavity in the hydrotreatment device, the lighter part with high olefin and oxygen content is upwards fed into the hydrofining reaction zone to carry out olefin saturation, deoxidization and deacidification reactions, and the heavier part is directly downwards fed into the hydrocracking reaction zone to carry out hydrocracking reactions, so that the occurrence of excessive cracking of the lighter part is avoided, and the yield of high-quality fuel oil is ensured to the greatest extent.

Description

A Fischer-Tropsch synthetic oil hydrotreating device and method

technical field

The invention belongs to the field of hydrogenation reaction, and in particular relates to a hydroprocessing device and method for Fischer-Tropsch synthetic oil.

Background technique

The Fischer-Tropsch synthesis reaction is a reaction in which hydrogen and carbon monoxide are used as raw materials to generate hydrocarbons under the action of a synthesis catalyst. Fischer-Tropsch synthesis raw material synthesis gas has a wide range of sources and can be converted from coal, natural gas, coal bed methane, biomass, etc. Fischer-Tropsch synthetic oil is mainly composed of normal alkanes, olefins and a certain amount of oxygen-containing compounds, and its non-ideal components such as sulfur, nitrogen, and aromatics are extremely low, and it is quite different from conventional mineral oil in terms of hydrocarbon composition and main properties. difference. The heavy naphtha fraction of Fischer-Tropsch synthetic oil basically does not contain sulfur and nitrogen, the octane number is very low (alkenes and alkanes are mostly straight chain), and the content of olefins and oxygen is relatively high; the content of sulfur, nitrogen and aromatics in the diesel fraction is extremely low , The cetane number is very high, the olefin content and oxygen content are in the middle, and the heavy oil has low olefin content and oxygen content. Therefore, Fischer-Tropsch synthetic oil must undergo corresponding hydrogenation and upgrading in order to obtain petrochemical raw materials or transportation fuels that meet the usage specifications.

The Fischer-Tropsch synthetic oil hydrotreating and upgrading technology reported in the industry mainly focuses on obtaining middle distillate diesel and aviation kerosene. Such as CN104711019A, CN1854265A, CN1854266A, CN101177626A, US5689031, etc., these processes generally include two parts of hydrofining and hydrocracking, using a fixed bed reactor, Fischer-Tropsch synthetic oil is generally first hydrofined saturated olefins and oxygenated Compounds, and then separate naphtha, diesel and heavy oil according to the distillation range, and the heavy oil is hydrocracked to obtain naphtha and high-quality diesel blending components. Fischer-Tropsch synthetic oil after hydrocracking, the naphtha obtained by fractionation has a high content of normal paraffins and high saturation, which is an ideal raw material for steam cracking to produce olefins; the obtained diesel cetane number reaches 70-80, not Containing sulfur, nitrogen and aromatics, it is an ideal vehicle diesel component. These processes are complicated, and the two parts of hydrotreating and hydrocracking are completely separated, and two fractionation systems are set up, which requires high equipment investment and high energy consumption.

US5378348, US6309432, CN101230291A and other patents first divide the Fischer-Tropsch synthetic oil into two or three different fractions, and the heavy fraction enters the hydroisomerization cracking reactor for isomerization cracking reaction, and then mixes with the light fraction into the hydrofining reactor The reaction is carried out, and the hydrogenated product enters the fractionation tower for fractionation. Although this flow process reduces the device load and simplifies the flow process, these inventions all use the traditional hydrogenation reaction fixed-bed reaction zone, and its aspect ratio (the ratio of the total height of the reactor bed to the diameter) is generally selected from 2 to 10. The temperature rise is high, especially the olefin content in Fischer-Tropsch synthetic oil is much higher than that of conventional mineral oil, which aggravates the risk of "flying temperature" in the catalyst bed. In order to lower the temperature, a large amount of cold hydrogen is injected between the various beds of the reactor to lower the reaction temperature, the reactor consumes a lot of energy, and in the initial stage and when the operation is abnormal, it is difficult for the cold hydrogen to take too much heat out of the reactor in time, so The service life of the catalyst is reduced, and in severe cases, the catalyst may fail and be scrapped.

The traditional hydrogenation reaction fixed bed has a high height-to-diameter ratio to ensure full contact between the gas-liquid material and the solid catalyst to achieve the required reaction depth and efficiency. Dong Fangliang and others mentioned in "First Heavy Technology" 1998.1 (Total 75), "Determination of Main Structural Parameters of Hydrogenation Reactor", in order to avoid small height-to-diameter ratio "to make fluid distribution uneven and lead to poor catalyst contact efficiency" , so "traditional fixed-bed reactors have an aspect ratio of 4-9 more". Patent CN109679689 A also mentions that the aspect ratio of existing liquid-phase hydrogenation reactors is generally 2.5-12. The design of the height-to-diameter ratio of the above-mentioned hydrogenation reactor bed has become the solidified cognition of those skilled in the art. The application of a large number of industrial practices has also proved that the design is reasonable and generally adaptable. The extensive industrial success may also lead to the unsatisfactory understanding of the technical personnel. For a more comprehensive and in-depth study of whether there are other better options for different types of reactions, there have been no relevant research reports for a long time, or only research reports that demonstrate that the existing height-to-diameter ratio is an appropriate design.

In addition, Fischer-Tropsch synthetic oil is mainly composed of straight-chain high-carbon alkanes and olefins, especially light fractions with a higher content of heavy olefins, which can reach more than 50%, while traditional petroleum-based catalytic cracking feedstocks are mainly composed of polycyclic hydrocarbons containing side chains. The composition of hydrocarbons, the molecular structure of the two is very different. Compared with polycyclic hydrocarbon molecules with side chains, the cracking activation energy of long-chain hydrocarbon molecules is lower, and the cracking reaction is easier to occur. Therefore, the cracked products may undergo cracking reactions again during the cracking process, making the reaction depth uncontrollable and producing a large amount of Gases and olefins, and can cause coking and plugging of the reactor. However, none of the existing methods can eliminate the coking and clogging of traditional reactors. This is because the product cannot leave the system during the flow of raw material oil and hydrogen in the entire reactor bed, and the material flows downward while reacting during the advancing process of the bed. During this period, both the fresh material reacts under the action of the catalyst, and the cracked product undergoes cracking reaction again.

Contents of the invention

In view of the deficiencies in the prior art, the object of the present invention is to provide a Fischer-Tropsch synthetic oil hydrotreating device and method. In order to effectively control the reaction depth, greatly increase the yield of heavy naphtha and diesel oil, eliminate the risk of "flying temperature" in the bed, and prolong the service life of the catalyst.

The first aspect of the present invention is to provide a Fischer-Tropsch synthetic oil hydrotreating device, said device comprising:

(1) Separator, reaction chamber and heavy oil tank connected sequentially from top to bottom;

(2) The reaction chamber is arranged sequentially from top to bottom: a hydrorefining catalyst bed, a separation chamber, a hydrocracking catalyst bed and a hydrogen distribution chamber.

In the above technical scheme, the ratio of the equivalent diameter of the cross-sectional area of the reaction chamber (the equivalent diameter formula is de=4A/L, A is the cross-sectional area of the reaction chamber, and the perimeter of the L bed) to the height is 2:1 to 10:1, Preferably it is 3:1-6:1.

In the above technical solution, the height of the reaction chamber is generally 100 mm to 5000 mm, preferably 200 mm to 1000 mm. With a higher diameter-to-height ratio, the flux of materials passing through the catalyst bed in the reaction chamber can be greatly increased, and at the same time, the residence time of materials and heat in the catalyst bed can be reduced.

In the above technical solution, the cavity of the reaction chamber may be in the form of a horizontal tank or a cylinder, preferably in the form of a horizontal tank. There are heads on both sides of the cavity, which is convenient for loading and unloading the catalyst. Catalysts are filled in the reaction chamber to form a reaction zone, which is divided into multiple reaction zone units by mesh partitions.

In the above technical solution, the height of the separation chamber provided in the reaction chamber is generally 10% to 30% of the height of the reaction chamber. The ratio of the bed height of the hydrotreating catalyst to the bed height of the hydrocracking catalyst is 5:1-1:8, preferably 2:1-1:6.

In the above technical solution, a liquid distribution component is arranged in the separation chamber, and the liquid distribution component includes a liquid distributor, a liquid distribution plate and a distribution cone arranged above the liquid distributor. It is used to disperse the liquid generated by the hydrogenation reactor into small droplets. Under the action of hydrogen stripping, the lighter part is taken out upwards, and the heavier part enters the hydrocracking reaction zone. The lighter fractions are usually naphtha and diesel fractions and the heavier fractions are usually heavy diesel and wax oil fractions.

In the above technical solution, the liquid distributor is a conventional distributor in the field, such as a shower head distributor, a coil distributor, a porous straight pipe distributor, a straight pipe baffle distributor, and a baffle distributor. Distributor, tangential circulation distributor, rotating vane distributor, double row vane distributor, etc. In the present invention, the liquid distributor is preferably a porous pipe distributor or a straight pipe baffle distributor, and the diameter of the pipe distributor is 0.5-20 mm, preferably 2-10 mm. The farther away from the feed oil inlet end, the larger the pore size. The height of the liquid distributor from the top of the reactor bed is 1-1000 mm, preferably 50-500 mm. The height is related to the nature, temperature and pressure of the raw oil. Generally speaking, when the temperature is higher, the distance between the liquid distributor and the bed is higher, so that the distributor can make the raw materials fall more evenly on the bed surface in a higher space. Similarly, the higher the pressure, the larger the injection angle of the liquid distributor, and the lower the height from the top of the reactor bed can be, saving more space.

In the above technical solution, the shape of the liquid distribution plate is the same as the cross-section of the catalyst bed in the hydrogenation main reactor, and the area of the liquid distribution plate is 10% to 100% of the cross-sectional area of the catalyst bed, preferably 60% to 100%.

In the above technical solution, a plurality of first through holes are evenly opened on the liquid distribution plate, a first overflow ring is provided around the first through holes, and an overflow portion is provided on the outer edge of the distribution plate; the opening ratio of the distribution plate is 5% ~90%, the diameter of the first through hole is 5-100mm, and the height of the first overflow ring is 1-30mm.

In the above technical solution, the inner side of the first overflow ring is provided with a sawtooth portion, the sawtooth portion is bent downward, and a diversion groove is provided on the sawtooth portion.

In the above technical solution, the distribution cone is arranged at the upper center of the liquid distribution plate, the distribution cone is provided with a plurality of second through holes, and a second overflow ring is provided around the second through holes. The apex angle of the distribution cone is greater than 90°, the opening ratio of the distribution cone is 5% to 80%, the height of the second overflow ring is 1 to 30mm; the bottom area of the distribution cone is 2% to 15% of the area of the liquid distribution plate %.

In the above technical solution, the ratio of the diameter or equivalent diameter of the separator to the diameter or equivalent diameter of the lower reaction chamber is 1:1.2-1:50, preferably 1:2-1:10. The diameter of the upper separator becomes smaller, so that the light fraction load under high pressure completely matches the tray, and the separation efficiency of the tray is high, which can completely replace the fractionation tower.

In the above technical solution, the separator is composed of a mixing section, a separation section and a stabilization section from bottom to top, the height of the mixing section is 20% to 35% of the total height of the separator, and the height of the separation section is 55% to 50% of the total height of the separator. 70%; the height of the stable section is 5% to 10% of the total height of the separator.

In the above technical scheme, packing or trays are placed in the separation section. Packing or trays are conventional forms in the field, such as packing can choose Pall ring, Raschig ring, saddle ring, saddle, open ring type, half ring, stepped ring, double arc, Haier ring, conjugated ring One or several kinds of random packing such as flat ring, rosette, etc. The packing can also choose metal or ceramic corrugated packing. The trays are one or more of the trays with downcomers such as bubble cap trays, sieve trays, valve trays, mesh plates, tongue-shaped plates, oriented sieve trays, and multi-downcomer trays. It is a tray without a downcomer such as a through-flow sieve tray and a through-flow corrugated plate. High-efficiency trays such as guided valves and sieve trays are preferred. The mixing section and the stabilizing section do not limit whether fillers are placed, and the reaction area can be increased according to the process requirements.

In the above technical solution, the hydrogen distribution chamber and the catalyst bed in the reaction chamber are provided with a plurality of partitions parallel to the vertical direction, and the plurality of partitions divide the hydrogen distribution chamber into a plurality of intake units, and each intake unit The bottom of the unit is provided with at least one hydrogen inlet. A plurality of holes are distributed on each separator; the separator extends upward to the catalyst layer, the aperture ratio of the separator below the catalyst bed is less than 70%, and the aperture ratio of the separator in the catalyst layer is greater than 50%.

In the above technical solution, a gas distributor is provided at the hydrogen inlet. In the present invention, the gas distributor is preferably a tangential circulation distributor or a rotating vane distributor. The gas distributor can make the flow rate of the gas entering the entire catalyst bed interface more uniform, and avoid bias flow and channel flow.

The hydrogen distribution chamber of the hydroprocessing device according to the present invention is provided with a hydrogen feed pipe, and the hydrogen feed pipe is provided with multiple inlets, and each hydrogen feed pipe corresponds to the catalyst bed area between two partitions, so that the hydrogen After exiting each sparger, it can pass upwards through the reaction zone at the top. The connection between the perforated baffle and the bottom of the main reactor has at least one channel.

In the above technical scheme, there are 1 to 3 gas distributors corresponding to the bottom of the catalyst bed between two adjacent partitions, and the distribution area of the hydrogen from all the gas distributors in the partition area to the bottom of the bed in the corresponding area should be Cover the entire area bottom of the bed. Further, the separator is in the shape of a ring or a segment.

In the above technical solution, the heavy oil tank is arranged at the center of the bottom of the reaction chamber, and the heavy oil tank communicates with the plurality of air intake units.

In the above technical solution, the hydrotreating device further includes: a reboiler, one end of which is connected to the outlet of the heavy oil tank, and the other end is connected to the hydrogen distribution chamber. Through the reboiler, the temperature of the heavy oil tank is kept at the temperature required by the reaction bed.

In the above technical solution, the hydrogenation treatment device also includes: multi-stage auxiliary reaction chambers, each auxiliary reaction chamber is fed with hydrogen independently, the bottom center is provided with a separate heavy oil tank, and the liquid raw material inlet of each auxiliary reaction chamber is connected to the The heavy oil tanks of the upper stage are connected, and the tops of the multi-stage auxiliary reaction chambers are all connected to the separator.

A second aspect of the present invention is to provide a Fischer-Tropsch synthetic oil hydrotreating method, comprising the steps of:

(1) The Fischer-Tropsch synthetic oil enters the separation chamber in the hydrotreating unit, and the lighter part is carried away by the hydrogen flowing upward from the bottom to enter the hydrorefining reaction zone; the heavier part enters the hydrocracking downward In the reaction zone, there is a cracking reaction with the countercurrent upward hydrogen, and the light components produced by the reaction leave the hydrocracking reaction zone upwards and then enter the hydrofining reaction zone, and the heavy components flow downward for reflux or part of them are discharged as tail oil;

(2) The stream after hydrofinishing enters the separator arranged on the upper part of the hydrofinishing reaction zone upwards, and is separated to obtain naphtha and diesel oil fractions, and the uncracked heavy components after separation enter the hydrocracking reaction zone downward again for hydrogenation Hydrogen cracking reaction.

In the above technical solution, after the Fischer-Tropsch synthetic oil enters the separation chamber in step (1), it is preferably dispersed into lighter parts and heavier parts through flash evaporation and liquid distribution components.

In the above-mentioned technical scheme, the respective equivalent diameters of the catalyst bed in the hydrofinishing reaction zone and the hydrocracking reaction zone in step (1) (the equivalent diameter formula is de=4A/L, A is the bed cross-sectional area, L bed circumference length) to the total height of the respective catalyst beds in a ratio of 2:1 to 10:1, preferably 3:1 to 6:1.

When there are no special instructions herein, the equivalent diameter-to-height ratio refers to the ratio of the equivalent diameter to the total height of the catalyst bed in the reactor. When there are multiple catalyst beds, the total height of the catalyst bed refers to the height of multiple catalyst beds. Sum. The bed cross-sectional area mentioned in the present invention refers to the cross-sectional area of the reactor bed. The reactor bed is preferably isodiametric, ie the cross-sectional area is the same at different locations throughout the catalyst bed. The cross-sectional area of the catalyst bed is usually the same as the cross-sectional area of the reaction chamber in the reactor, and the cross-section refers to the cross-section in a top view, that is, the cross-section perpendicular to the vertical line in the reaction chamber. If there is a difference in the cross-sectional area of the reactor within the height range of the catalyst bed, the cross-sectional area here refers to the average value of the cross-sectional area of the catalyst bed or the cross-sectional area of the reaction chamber within the range of the catalyst bed.

The Fischer-Tropsch synthetic oil in the present invention includes high-temperature Fischer-Tropsch synthetic whole distillate oil and low-temperature Fischer-Tropsch synthetic whole distillate oil, preferably low-temperature Fischer-Tropsch synthetic whole distillate oil. The properties of the Fischer-Tropsch synthetic oil are as follows: a density of 0.6g/cm ³ to 1.0g/cm ³ , preferably 0.7g/cm ³ to 0.95g/cm ³ ; an end boiling point of 650°C to 750°C, preferably 680°C to 720°C .

In the above technical solution, the hydrofinishing reaction zone described in step (1) is filled with a hydrofinishing catalyst. Hydrofining catalysts are generally based on alumina or silicon-containing alumina as the carrier, and the metals of Group VIB and Group VIII as active components, such as one, two or more of W, Mo, Co, and Ni.

In the above technical solution, the hydrocracking reaction zone described in step (1) is filled with a hydrocracking catalyst. The catalyst generally includes an active component and a carrier, and 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 components It 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 hydrorefining catalyst and the hydrocracking catalyst can be any conventional existing hydrocracking catalyst shape, preferably a porous catalyst, a heterogeneous catalyst and/or a honeycomb catalyst. The pore diameter of the porous catalyst is 1-50mm, preferably 4-20mm; the average particle diameter of the special-shaped catalyst is 2-50mm, preferably 4-30mm; the pore diameter or hole side length of the honeycomb catalyst is 1-50mm, preferably 3-15mm; the catalyst bed The porosity of the layer is recommended to be 15% to 85%, preferably 20% to 75%.

In the above technical scheme, the operating conditions of the Fischer-Tropsch synthetic oil hydrogenation reaction process (including hydrofinishing and hydrocracking) described in step (1) are as follows: the reaction temperature is 380°C to 450°C, and the reaction pressure is 3MPa to 15MPa , the top reflux ratio is 1.2-4.5, the hydrogen-oil volume ratio is 300-2000, and the liquid hourly volume space velocity is 0.1h ^-1 -10.0h ^-1 . The preferred operating conditions are: reaction temperature 400°C-450°C, reaction pressure 4MPa-12MPa, top reflux ratio 1.5-3.0, hydrogen-oil volume ratio 400-1500, liquid hourly volume space velocity 0.5h ^-1 ～ 10.0h ^-1 .

In the above technical solution, the separator draws out the product at the side line or at the top. Gases are drawn overhead, and naphtha and diesel fractions are drawn in the sideline.

In the above technical solution, it is preferable to open 2 to 3 side lines in the separation section of the separator for extracting the required products. At the top of the separator, the extraction temperature is 60°C to 80°C, and the extracted components are gas and light naphtha fractions. The extraction temperature of the first side line is 150°C to 190°C, preferably 160°C to 180°C, and the component extracted from this side line is a heavy naphtha fraction. The second side line extraction temperature is 300°C to 380°C, preferably 310°C to 370°C, and the component extracted from the side line is diesel fraction.

Further, in the above technical solution, the side line extraction line can be provided with backflow.

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, due to the rapid increase of gas phase amount, it occupies a large number of bed voids, which greatly increases the liquid phase flow rate. According to the traditional design, although sufficient contact between the gas-liquid-solid three-phase can be ensured, 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) increases the contact probability with the catalyst. , for a system that requires more transformation in the liquid phase and secondary reactions controlled by the gas phase, the overall reaction effect is limited to a certain extent, and it is generally difficult to further improve the reaction conversion rate and selectivity.

It has been found through research that when the overall space velocity is similar, for the gas-liquid-solid three-phase hydrogenation reaction in which the amount of liquid phase decreases rapidly and the amount of gas phase increases rapidly during the reaction process, when the hydrogen and the raw material oil, gas, and liquid are contacted in a countercurrent manner, the catalyst bed in the reactor When the layer diameter-to-height ratio is significantly higher than that of the existing conventional technology, the generated gas phase quickly leaves the catalyst bed, and the cumulative effect of the adverse effects of the gas phase is small, and the liquid phase can have a more sufficient chance of reacting on the catalyst, thereby overcoming The traditional understanding that a small aspect ratio will bring adverse effects such as poor contact effect has achieved a significant increase in the yield of the target product (heavy naphtha, the target product in the hydrogenation technology of inferior diesel oil), and at the same time solved the problem of countercurrent reactors. Easy liquid flooding, limited hydrogen-to-oil ratio, etc.

Compared with the prior art, the present invention has the advantages of:

(1) The present invention fully considers the characteristics of the large difference in hydrocarbon composition, olefin and oxygen content of each fraction of Fischer-Tropsch synthetic oil, and separates the Fischer-Tropsch synthetic oil with a wide distillation range distribution by setting a separation chamber in the hydrotreating device. The lighter part and the heavier part, the lighter part with high olefin and oxygen content enters the hydrorefining reaction zone upwards for olefin saturation, deoxygenation and deacidification reaction, and the heavier part directly enters the hydrocracking reaction zone downward for hydrocracking reaction , reducing the amount of hydrorefining and hydrocracking catalysts, avoiding the occurrence of excessive cracking of lighter parts on hydrocracking catalysts, and ensuring the yield of high-quality fuel oil to the greatest extent. At the same time, the light products produced after the hydrocracking of the heavier part contain a small amount of olefins, and the olefins in it can be hydrogenated and saturated by the upward hydrorefining catalyst, which ensures the stability of product quality.

(2) The design of the liquid distribution component can make the heavy component liquid flowing downward in the separation chamber dispersed into suitable small droplets, and by means of the stripping effect of hydrogen, the lighter part can be taken out directly, without Then enter the hydrocracking reaction zone to participate in the hydrocracking reaction. At the same time, through the design of the liquid distribution components, the heavy components entering the hydrocracking reaction zone can be distributed more evenly, which solves the problem of poor contact of reactants on the catalyst bed with a large diameter to height ratio in the traditional reactor. By adopting the reactor of the present invention, under the same process conditions and product index requirements, the void ratio of the reactor of the present invention can be smaller. The temperature rise of the reaction bed is controlled smaller and the allowable feed load is larger. The improvement of the pore structure can improve the flooding characteristics and at the same time ensure good mass transfer performance. Finally, the product properties have good controllability.

(3) The present invention can realize timely extraction of naphtha and diesel oil products through the setting of the separator through flashing and steam stripping, which has not only taken away a large amount of heat released by the reaction, but also controls the degree of reaction to prevent naphtha The excessive cracking and gasification of oil and diesel products ensures the maximum yield of naphtha and diesel products. At the same time, because the partial pressure of the product has been kept low, the reaction speed is accelerated, the hidden danger of the catalyst being blocked by by-products is eliminated, the yield of the target product is improved, and the service life of the catalyst is prolonged.

Description of drawings

Fig. 1 is the structure schematic diagram of Fischer-Tropsch synthetic oil hydrotreating device of the present invention;

Fig. 2 is a schematic diagram of the Fischer-Tropsch synthetic oil hydrotreating process flow diagram and a structural schematic diagram of a hydrotreating device of the present invention;

Among them: 1-Fischer-Tropsch synthetic oil; 2-hydrogen; 3-hot feed oil; 4-hot hydrogen; 5-inlet pipeline of hydrogenation reactor; 6-reaction chamber of hydrogenation reactor; 7-heavy oil of hydrogenation reactor silo; 8—hydrogenation reactor separation chamber; 9—isomerization depreciation catalyst bed in mixing section; 10—stabilization section; 11—hydrofining catalyst bed; 12—hydrocracking catalyst bed; 13—separation section; 14 －Liquid distribution component; 15－Mixing section; 16－Separator; 17－Separator overhead condenser; 18－Separator tank; 19－Hydrogen-rich and non-condensable gas; ; 22-light naphtha fraction; 23-hydrogenated heavy fraction; 24-reflux heavy fraction; 25-circulation oil pump; 26-tail oil; 27-mesh baffle;

Fig. 3 is a side structural schematic diagram of a circular partition according to an embodiment of the present invention.

Fig. 4 is a schematic side view of a circular partition according to another embodiment of the present invention, the partition is only located in the hydrogen distribution chamber, and there is no circular partition inside the catalyst bed.

Fig. 5 is a schematic top view of an annular partition according to an embodiment of the present invention.

Fig. 6 is a schematic side view of the annular partition according to two embodiments of the present invention. The partition in Fig. 6-2 is only located in the hydrogen distribution chamber, and there is no annular partition inside the catalyst bed.

Fig. 7 is a schematic top view of a liquid distribution assembly according to an embodiment of the present invention.

Fig. 8 is a top view and a side view schematic diagram of a liquid distribution assembly according to another embodiment of the present invention.

Fig. 9 is a schematic top view of the first overflow ring according to an embodiment of the present invention.

FIG. 10 is a schematic perspective view of the first overflow ring in FIG. 8 .

Detailed ways

The present invention will be further described below in conjunction with the accompanying drawings.

Figure 1 and Figure 2 show the Fischer-Tropsch synthetic oil hydrotreating device and process flow of the present invention. The Fischer-Tropsch synthetic oil 1 is sent to the heating furnace 28, heated to 380°C-450°C to obtain the hot raw material oil 3, which enters the hydrogenation reactor separation chamber 8 in the hydrogenation reactor reaction chamber 6 through the hydrogenation reactor inlet pipeline 5 Among them, after being dispersed by the flashing and liquid distribution assembly 14, the lighter part is carried away by the hydrogen flowing upward from the bottom and enters the hydrotreating catalyst bed 11 (including 11-1～11-6); the heavier part A part is uniformly sprayed downward on the hydrocracking catalyst bed 12 (including 12-1 to 12-6). At the same time, the hot hydrogen 4 after the hydrogen gas 2 is heated by the heating furnace 28 is sent into the reaction chamber 6 of the hydrogenation reactor. The bottom of the zone moves upwards in countercurrent contact with the heavier fraction sprayed from the liquid distribution assembly within the hydrocracking catalyst bed 12 and cocurrently within the hydrofinishing catalyst bed 11 with the lighter fraction flowing upward from the separation chamber up. Under the operating pressure of 4MPa-15MPa, the heavier component of Fischer-Tropsch synthetic oil undergoes hydrocracking reaction on the bed of hydrocracking catalyst, some long-chain molecules are broken into short-chain molecules, and polycyclic aromatic hydrocarbons are also partially broken; The light components of the Tropical synthetic oil, the light components of the cracking reaction and the hydrogen are subjected to reactions such as olefin hydrogenation saturation, hydrodeoxygenation, and impurity removal on the hydrotreating catalyst bed 11 .

Smaller molecular hydrocarbons or other gases produced by the hydrogenation reaction are quickly brought to the separator 16 by the hydrogen, and after the separation of the mixing section 15 and the separation section 13, a part of the heavier fraction passes downward through the hydrotreating catalyst bed 11. The separation chamber 8 of the hydrogenation reactor and the liquid distribution component 14 fall evenly onto the surface of the hydrocracking catalyst bed 12 . Another part of the light component continues to move to the stable section 10 at the top of the separator, flows out from the top of the separator 16, continues to be cooled by the condenser 17 at the top of the separator, and then passes through the liquid separation tank 18 for gas-liquid separation. Hydrogen and non-condensable gas 19 are discharged, and liquid hydrocarbons are refluxed or extracted as light naphtha fraction 22 . The heavy naphtha product 20 is extracted from the side line of the separator at 150°C to 190°C, and the diesel product 21 is extracted from the side line of the separator at 300°C to 380°C.

The hydrogenated heavy fraction 23 in the hydrogenation reactor flows out from the heavy oil tank 7 of the hydrogenation reactor, and is used as the reflux heavy fraction 24 through the circulating oil pump 25 and mixed with the raw material oil as the feed to the hydrogenation reactor, and can also be partially used as tailings. The oil is discharged from the device through the circulating oil pump 25.

Further, in one or more exemplary embodiments of the present invention, the reaction chamber 6 of the hydrogenation reactor can be a horizontal storage tank, as shown in FIG. With head. Further, in one or more exemplary embodiments of the present invention, the reaction chamber 6 may also be a flat cylindrical tank, the axial direction of which is arranged along the longitudinal direction.

Further, in one or more exemplary embodiments of the present invention, the shape of the partition matches the bottom of the reaction chamber 6, and when the reaction chamber 6 is a horizontal storage tank, the partition is a round partition, such as As shown in Figure 3 and Figure 4; when the reaction chamber 6 is a flat cylindrical tank, the plurality of partitions are coaxial annular partitions, as shown in Figures 5-7. Further, in one or more exemplary embodiments of the present invention, a plurality of round holes are distributed on each separator. Further, in one or more exemplary embodiments of the present invention, as shown in FIG. 1 , a plurality of mesh baffles 27 may extend upwards to the hydrotreating catalyst bed 11 , and are not connected to the catalyst bed 11 at the bottom. The opening ratio of the contacting separator is less than 70%, and the lower opening ratio is beneficial to increase the resistance, so that the hydrogen gas enters the hydrotreating catalyst bed 11 as far as possible upwards, and further functions as a gas distributor. The opening ratio of the separator in the hydrorefining catalyst bed layer 11 is greater than 50%, which is beneficial to utilize the catalyst more fully.

Further, in one or more exemplary embodiments of the present invention, the separator 16 includes a mixing section 15 , a separating section 13 and a stabilizing section 10 from bottom to top.

Further, in one or more exemplary embodiments of the present invention, the liquid distribution assembly 14 provided at the bottom of the separation chamber 8 of the hydrogenation reactor includes a liquid distribution pan and a distribution cone. The liquid distribution plate is arranged above the liquid distributor, the shape of the liquid distribution plate is the same as that of the top surface of the porous catalyst layer, a plurality of first through holes are evenly opened on the liquid distribution plate, and a first overflow ring is arranged around the first through holes, An overflow portion (not shown in the figure) is provided on the outer edge of the distribution plate. The distribution cone is arranged at the upper center of the liquid distribution plate, and the distribution cone is provided with a plurality of second through holes, and a second overflow ring (not shown in the figure) is arranged around the second through holes.

Preferably but not limitatively, in one or more exemplary embodiments of the present invention, as shown in FIG. 9 and FIG. 10 , the inner side of the first overflow ring is provided with a sawtooth portion 14-1, and the sawtooth portion is bent downward , The sawtooth part is provided with diversion groove 14-2. Exemplarily, the guide groove is opened along the center of the serrated part.

In one or more embodiments of the present invention, the hydrogenation reactor further includes an auxiliary reaction chamber. It should be understood that the auxiliary reaction chamber may be multi-stage. Each level of auxiliary reaction chamber is fed with hydrogen separately, and the center of the bottom is equipped with a separate heavy oil tank. The liquid raw material inlet of each level of auxiliary reaction chamber is connected to the heavy oil tank of the previous level, and the top of the multi-level auxiliary reaction chamber is connected to to separator 16.

The present invention will be further described below in conjunction with specific examples, but it should be understood that the protection scope of the present invention is not limited by the specific embodiments.

Example 1

Using the device and process flow shown in Fig. 1 and Fig. 2 of the present invention, the raw oil is Fischer-Tropsch synthetic oil, and its properties are shown in Table 1. After the raw oil and hydrogen are heated to 430°C by the heating furnace, they enter the hydrogenation reactor. In the hydrogenation reactor, the hydrorefining catalyst FHUDS-5 bed produced by Sinopec Dalian Petrochemical Research Institute and the The bed of hydrocracking catalyst FC-14, the ratio of equivalent diameter to height of each catalyst bed is 4:1, the height of hydrofining catalyst bed and hydrocracking catalyst bed are 800mm respectively. The height of the separation chamber between the hydrorefining catalyst bed and the hydrocracking catalyst bed is 15% of the height of the reaction chamber. Ring-shaped partitions are arranged in the catalyst bed, and the number of partitions is four. A plurality of holes are distributed on the separator; the separator extends upward to the catalyst layer, the aperture ratio of the separator below the catalyst layer is 40%, and the aperture ratio of the separator in the catalyst layer is 70%.

The raw material liquid is firstly separated from the light and heavy components through the liquid distribution component. The light components enter the separator after being hydrotreated upwards, and the heavy components enter the four different sub-catalyst bed areas in the hydrocracking catalyst bed. The hydrogen entering from the bottom will undergo reactions such as hydrocracking and olefin saturation under the action of the hydrocracking catalyst. The light fraction produced by the reaction quickly leaves the reaction system upwards and enters the separator after hydrofining. The separator consists of a mixing section, a separating section and a stabilizing section from bottom to top, and the height of the mixing section is 30% of the total height of the separator , the height of the separation section is 65% of the total height of the separator; the height of the stable section is 5% of the total height of the separator. Raschig rings are placed in the mixing section and the stabilizing section. After hydrofinishing, the stream is mixed with the light components in the lower part of the mixing section of the separator, and then continues upwards. Under the action of the separation tray in the mixing section and the separation section, the heavy fractions re-enter the hydrocracking catalyst bed for cracking reaction downwards. The light fraction is taken out as a heavy naphtha fraction and diesel fraction in the side line and sent out as a product, and the non-condensable gas continues to be cooled by the condenser, and after gas-liquid separation in the liquid separation tank, the liquid hydrocarbon is extracted as a light naphtha fraction, and the gas enters the desulfurization The ammonia removal equipment is purified and recycled. The heavy distillates that are not sufficiently cracked flow out from the bottom of the heavy oil tank and pass through the circulation pump, and then all enter the inlet of the hydrotreating unit as circulating oil. The specific operating process conditions are shown in Table 2, and the product distribution and properties are shown in Table 4.

The liquid distribution assembly includes a liquid distributor, a liquid distribution plate and a distribution cone arranged above the liquid distributor. The shape of the liquid distribution plate is the same as that of the top surface of the catalyst bed, and the area of the liquid distribution plate is 70% of the cross section of the catalyst bed. A plurality of first through holes are evenly opened on the liquid distribution plate, a first overflow ring is arranged around the first through holes, an overflow portion is provided on the outer edge of the distribution plate, the opening rate of the liquid distribution plate is 50%, and the first overflow ring is arranged around the first through hole. The diameter of a through hole is 10mm, and the height of the first overflow ring is 10mm. The distribution cone is arranged on the upper center of the liquid distribution plate, and the distribution cone is provided with a plurality of second through holes, and a second overflow ring is arranged around the second through holes; the apex angle of the distribution cone is 120°, and the opening of the distribution cone The ratio is 50%, the height of the second overflow ring is 10mm; the bottom area of the distribution cone is 10% of the area of the liquid distribution plate.

Example 2

The difference between this example and Example 1 is that the equivalent diameter-to-height ratio of the hydrofining and hydrocracking catalyst beds is 5:1, and the number of partitions in the catalyst bed is 6. A plurality of holes are distributed on the separator; the separator extends upward to the catalyst layer, the aperture ratio of the separator below the catalyst layer is 30%, and the aperture ratio of the separator in the catalyst layer is 80%. The area of the liquid distribution tray is 90% of the cross-section of the catalyst bed. The opening ratio of the liquid distribution plate is 80%, the diameter of the first through hole is 20 mm, and the height of the first overflow ring is 20 mm. The distribution cone is arranged on the upper center of the liquid distribution plate, and the distribution cone is provided with a plurality of second through holes, and a second overflow ring is arranged around the second through holes; the apex angle of the distribution cone is 150°, and the opening of the distribution cone The ratio is 70%, the height of the second overflow ring is 20mm; the bottom area of the distribution cone is 15% of the area of the liquid distribution plate. All the other conditions are the same as in Example 1.

Example 3

The difference between this example and Example 1 is that the equivalent diameter-to-height ratio of the hydrofining and hydrocracking catalyst bed is 6:1, and the number of partitions in the catalyst bed is 6. The area of the liquid distribution tray is 80% of the cross-section of the catalyst bed. The opening ratio of the liquid distribution plate is 70%, the diameter of the first through hole is 30 mm, and the height of the first overflow ring is 20 mm. The distribution cone is arranged on the upper center of the liquid distribution plate, and the distribution cone is provided with a plurality of second through holes, and a second overflow ring is arranged around the second through holes; the apex angle of the distribution cone is 140°, and the opening of the distribution cone The ratio is 60%, the height of the second overflow ring is 20mm; the bottom area of the distribution cone is 15% of the area of the liquid distribution plate. All the other conditions are the same as in Example 1.

Example 4

The difference between this embodiment and Example 1 is that the height of the mixing section of the separator is 25% of the total height of the separator, the height of the separation section is 65% of the total height of the separator; the height of the stabilizing section is 10% of the total height of the separator. All the other conditions are the same as in Example 1.

Comparative example 1

The conventional two-stage hydrogenation method is adopted, that is, refining + cracking process. Both refining and cracking reactors adopt a reaction process in which raw materials and hydrogen flow in parallel from top to bottom. Fischer-Tropsch synthetic oil is first saturated with olefins in the hydrotreating reactor, and after deoxygenation and impurity removal reactions, the hydrofinishing stream enters the first separation and fractionation system to obtain refined high-pressure gas, naphtha, diesel oil and heavy oil, among which refined high-pressure The gas flows back into the hydrotreating reactor, and the heavy oil enters the hydrocracking reactor for hydrocracking reaction. After the reaction is completed, it enters the second separation and fractionation system. After separation, gas, naphtha and diesel fractions are obtained. After separation, the uncracked All the heavy components are refluxed into the hydrocracking reactor to carry out the hydrocracking reaction again. The catalyst embodiment 1 adopted is the same, and the processing conditions are shown in Table 3.

Table 1 Main Properties of Raw Oil

project data <![CDATA[Density, g/cm³]]> 0.815 Distillation range, ℃ Initial boiling point/10% 65/203 30%/50% 312/371 70%/90% 460/557 95% 617 Sulfur content, vol% 4.1 Nitrogen content, μg/g 4.9 Oxygen content, wt% 0.82

Table 2 embodiment technological conditions

Example 1 Reaction temperature, °C 410 Reaction pressure, MPa 8.0 Volume Hydrogen Oil Ratio 1000:1 <![CDATA[volumetric space velocity, h^-1]]> 1.5

Table 3 Comparative Example 1 process conditions

Example 1 One stage hydrofining reactor Reaction temperature, °C 300 Reaction pressure, MPa 8.0 Volume Hydrogen Oil Ratio 300:1 <![CDATA[volumetric space velocity, h^-1]]> 3.0 Second stage hydrocracking Reaction temperature, °C 410 Reaction pressure, MPa 8.0 Volume Hydrogen Oil Ratio 1000:1 <![CDATA[volumetric space velocity, h^-1]]> 1.5

Table 4 Product distribution and properties

Example 1 Example 2 Example 3 Example 4 Comparative example 1 naphtha fraction Distillation range, ℃ 65～160 65～160 65～160 65～160 65～160 Yield, wt% 32.2 31.3 33.8 32.3 18.4 <![CDATA[Density, g/cm³]]> 0.708 0.705 0.710 0.706 0.714 Composition, wt% n-alkanes 91.20 92.30 91.57 91.74 87.57 Isoparaffin 6.54 6.52 6.81 7.02 10.24 diesel fraction Distillation range, ℃ 160～370 160～370 160～370 160～370 160～370 Yield, wt% 66.1 67.2 65.0 66.8 78.1 <![CDATA[Density, g/cm³]]> 0.781 0.773 0.765 0.780 0.793 freezing point, ℃ -3 -4 -5 -3 -2 cetane number 83 83 84 83 78

From the result of table 4, adopt the method of the present invention, the yield of diesel oil is obviously higher than that of comparative example 1, and cetane number is greater than 83, is high-quality diesel blending component; Naphtha fraction is mainly made up of alkane, is high-quality Raw materials for steam cracking of ethylene and propylene.

Example 5

The laboratory simulated and calculated the bed reaction temperature distribution of the examples and comparative examples by using the ANSYS version 19.0 software. The simulation conditions are input according to the actual data of the examples and comparative examples. The simulation results show that the center temperature of the traditional fixed bed is the highest, and the temperature change from the inlet end to the outlet end shows a normal distribution, while the reactor bed temperature of the present invention is relatively uniform. Table 5 shows the simulated temperature rise of the bed.

Table 5 Bed simulated temperature rise changes

bed temperature point Example 1 Example 2 Comparative example 1 Maximum radial temperature difference, ℃ 3.1 3.5 28.1 average temperature, ℃ 411.2 411.5 431.6

As can be seen from the results in Table 5, the temperature difference of the catalyst bed in Embodiments 1 to 2 of the present invention is significantly lower than that of Comparative Example 1, from the temperature difference of 28.1° C. in the traditional fixed bed down to 3.5° C. The difference between the temperature and the control temperature is small, indicating that the reactor of the present invention has eliminated the overheating phenomenon of the hydrocracking reaction.

Claims (16)

1. A fischer-tropsch synthesis oil hydrotreater, the apparatus comprising:

(1) The separator, the reaction cavity and the heavy oil bin are sequentially communicated from top to bottom;

(2) The reaction chamber is internally provided with: a hydrofining catalyst bed, a separation cavity, a hydrocracking catalyst bed and a hydrogen distribution cavity.

2. The apparatus according to claim 1, wherein the ratio of reaction chamber cross-sectional area equivalent diameter to height is 2:1 to 10:1, preferably 3:1 to 6:1.

3. The device according to claim 1, characterized in that the reaction chamber has a height of 100mm to 5000mm, preferably 200mm to 1000mm.

4. The device according to claim 1, wherein the height of the separation chamber arranged in the reaction chamber is 10% -30% of the height of the reaction chamber; the ratio of the height of the hydrofining catalyst bed to the height of the hydrocracking catalyst bed is 5:1-1:8, preferably 2:1-1:6.

5. The apparatus of claim 1, wherein a liquid distribution assembly is disposed within the separation chamber, the liquid distribution assembly comprising a liquid distributor and a liquid distribution tray and cone disposed above the liquid distributor.

6. The apparatus of claim 5, wherein the liquid distribution tray has the same shape as the cross section of the catalyst bed of the hydrogenation main reactor, and the area of the liquid distribution tray is 10% to 100%, preferably 60% to 100%, of the cross section area of the catalyst bed.

7. The device according to claim 5, wherein the liquid distribution plate is uniformly provided with a plurality of first through holes, a first overflow ring is arranged around the first through holes, and an overflow part is arranged at the outer edge of the distribution plate; the aperture ratio of the distribution plate is 5-90%, the diameter of the first through hole is 5-100 mm, and the height of the first overflow ring is 1-30 mm.

8. The apparatus of claim 5, wherein the dispensing cone is disposed in the upper center of the liquid dispensing tray, the dispensing cone having a plurality of second through holes and a second overflow ring disposed around the second through holes. The vertex angle of the distribution cone is larger than 90 degrees, the aperture ratio of the distribution cone is 5-80%, and the height of the second overflow ring is 1-30 mm; the bottom area of the distribution cone is 2% -15% of the area of the liquid distribution disc.

9. The apparatus according to claim 1, wherein the ratio of separator diameter or equivalent diameter to the diameter or equivalent diameter of the lower reaction chamber is 1:1.2 to 1:50, preferably 1:2 to 1:10.

10. The device according to claim 1, wherein the separator comprises a mixing section, a separating section and a stabilizing section from bottom to top, the height of the mixing section is 20% -35% of the total height of the separator, and the height of the separating section is 55% -70% of the total height of the separator; the height of the stabilizing section is 5-10% of the total height of the separator.

11. A process for the hydroprocessing of fischer-tropsch synthesis oil using the apparatus of any one of claims 1 to 10, comprising the steps of:

(1) The Fischer-Tropsch synthesis oil enters a separation cavity in the hydrotreater, wherein a lighter part is carried away upwards by hydrogen flowing upwards from the bottom and enters a hydrofining reaction zone; the heavier part enters a hydrocracking reaction zone downwards to carry out a cracking reaction with countercurrent upward hydrogen, and the light component generated by the reaction is separated upwards from the hydrocracking reaction zone and then enters a hydrofining reaction zone, and the heavy component flows downwards to be refluxed or partially thrown outwards to be discharged as tail oil;

(2) The material flow after hydrofining upwards enters a separator arranged at the upper part of the hydrofining reaction zone, naphtha and diesel oil fractions are obtained through separation, and the separated uncracked heavy components downwards enter the hydrocracking reaction zone again for hydrocracking reaction.

12. The process according to claim 11, wherein the ratio of the equivalent catalyst bed diameter of each of the hydrofinishing reaction zone and the hydrocracking reaction zone to the total height of the respective catalyst bed in step (1) is from 2:1 to 10:1, preferably from 3:1 to 6:1.

13. The method according to claim 11, characterized in that the fischer-tropsch synthesis oil properties are as follows: density of 0.6g/cm ³ ～1.0g/cm ³ Preferably 0.7g/cm ³ ～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 ℃.

14. The process of claim 11, wherein the operating conditions of the fischer-tropsch synthesis oil hydrogenation process of step (1) are as follows: the reaction temperature is 380-450 ℃, the reaction pressure is 3-15 MPa, the reflux ratio of the tower top is 1.2-4.5, the hydrogen oil volume ratio is 300-2000, 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 400-450 ℃, the reaction pressure is 4-12 MPa, the reflux ratio of the top of the tower is 1.5-3.0, the volume ratio of hydrogen to oil is 400-1500, and the liquid hourly space velocity is 0.5h ^-1 ～10.0h ^-1 。

15. The process of claim 11, wherein the separator withdraws product at the side or top, gas at the top, naphtha and diesel fractions at the side; preferably 2 to 3 side lines are cut off in the separation section of the separator for withdrawing the desired product.

16. The process according to claim 15, wherein at the top of the separator, the temperature of extraction is 60 ℃ to 80 ℃, and the components extracted are gas and light naphtha fraction; the temperature of the 1 st side draw is 150 ℃ to 190 ℃, preferably 160 ℃ to 180 ℃, and the components drawn from the side draw are heavy naphtha fraction; the temperature of the 2 nd side draw is 300 ℃ to 380 ℃, preferably 310 ℃ to 370 ℃, and the component drawn from the side draw is diesel oil distillate.

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