A high-silicon naphtha processing system and processing technology

Abstract

The present invention discloses a high-silicon naphtha processing system and a processing technology. The high-silicon naphtha processing system of the present invention comprises a fixed bed reactor, a desiliconization reactor, a heat exchanger and a fractionation unit; the fixed bed reactor is composed of a gas phase reaction zone, a gas-liquid separation zone, a gas-liquid countercurrent zone, a gas-liquid contact zone and a liquid phase reaction zone from top to bottom. The high-silicon naphtha processing technology comprises: hydrogen first enters the gas-liquid contact zone at the bottom of the fixed bed reactor, the high-silicon naphtha and the olefin-rich raw material enter the gas-liquid separation zone from the top of the fixed bed reactor, the materials react in each reaction zone, the desiliconized product is mixed with the bottom effluent of the liquid phase reaction zone after heat exchange, and stripping and fractionation are performed to obtain qualified products. The present invention not only realizes the efficient removal of silicon in high-silicon raw materials, but also can simultaneously process conventional coking raw materials, obtain qualified chemical raw materials and automotive diesel products, and at the same time extend the operation cycle of the coking oil processing device.

Description

Technical Field

[0001] The invention belongs to the field of clean oil refining, and in particular relates to a high-silicon naphtha processing technology. Background Art

[0002] At present, the hydrogenation process for coking oil products has the problem of short unit operation cycle. On the one hand, the raw materials contain cyclosiloxane silicon species. Although they can be removed under certain conditions, silicon is easy to deposit and cause catalyst poisoning. On the other hand, the olefin content in coking oil products is too high, and the intense heat release during the reaction causes coking and carbon deposition on the catalyst surface, and causes bed pressure drop. The above two reasons cause the operation cycle of the hydrogenation unit of coking oil products to be significantly lower than that of other oil hydrogenation units.

[0003] With the overcapacity of my country's fuel oil market, oil refining companies are more inclined to convert naphtha into ethylene cracking raw materials and reforming feeds through hydrogenation technology to create more value. Naphtha fractions are lighter and have less impurities, so they are generally not difficult to process. Conventional catalysts can meet the requirements of reforming feeds and ethylene cracking raw materials under certain process conditions. However, due to the complex sources of naphtha, the types of impurities in it are also complex, especially the various types of silicon-containing compounds, which require strict desiliconization conditions, which will inevitably have an adverse effect on other components in the oil. It is difficult to convert such difficult-to-remove, high-content silicon-containing oils into qualified reforming or ethylene feeds with current technology.

[0004] CN201911212127.9 discloses a grading method and a hydrofining method for a coking gasoline hydrofining catalyst. The method comprises: a replaceable zone and a main reaction zone are sequentially arranged in series along the direction of logistics, wherein the replaceable zone is sequentially loaded with a diolefin saturation catalyst in the first reactor and a silicon trap in the second reactor along the direction of logistics, wherein the second reactor is a parallel dual reactor and can be switched online; the main reaction zone is sequentially loaded with an arsenic removal agent and a gasoline hydrogenation catalyst along the direction of logistics, and the invention solves the problem of catalyst coking.

[0005] CN200810166851.8 discloses a method for extending the service life of a catalyst for hydroprocessing of coked gasoline and diesel. After the coked gasoline and diesel are mixed with hydrogen at the reactor inlet, they enter the hydrogenation reactor from the upper part of the reactor and enter two reaction zones in sequence. The first reaction zone is loaded with a desiliconization catalyst with high desiliconization activity and high silicon capacity, and the second reaction zone is loaded with a conventional hydroprocessing catalyst. The silicon in the coked gasoline and diesel is removed and deposited on the desiliconization catalyst before entering the conventional hydroprocessing catalyst, thereby protecting the hydroprocessing catalyst from being contaminated by silicon. The invention uses graded loading of silicon traps to extend the service life of the main hydrogenation catalyst.

[0006] The above patented technologies are applied to the processing technology of naphtha containing high content of silicon-containing compounds that are difficult to remove, and the obtained products cannot meet the problem of high-quality reforming feed and ethylene cracking raw material.

[0007] In summary, the difficulty in processing conventional coke oil and high-silicon naphtha still needs to be solved in order to provide refineries with more suitable reforming and ethylene raw materials. Summary of the invention

[0008] In view of the shortcomings of the prior art, the present invention proposes a high-silicon naphtha processing system and processing technology. The present invention matches the raw material characteristics with the process flow, and through zoned reaction control and full utilization of reaction heat, it not only achieves efficient removal of silicon in high-silicon raw materials, but also can process conventional coking raw materials at the same time, thereby obtaining qualified chemical raw materials and automotive diesel products, and at the same time extending the operation cycle of the coking oil processing device.

[0009] The existing naphtha desiliconization technology is generally used for desiliconization of oil products with a silicon content of less than 100ppm and in the form of epoxysilane. Such silicon species are easier to remove under appropriate temperature and pressure. However, in recent years, with the change in crude oil quality, the properties of naphtha raw materials have also changed. The silicon content in some naphtha raw materials is as high as 300-1000ppm, which greatly exceeds the conventional silicon content range. The inventors have found through research that the types of silicon compounds have also changed, with more alkylsilanes, silanols or silane ethers containing silicon compounds. Such silicon species are enriched in the light components of naphtha. After the raw oil products are treated according to the desiliconization methods and processes of the prior art, it is difficult to make the silicon content meet the standard.

[0010] The high-silicon naphtha processing system of the present invention comprises a fixed bed reactor, a desiliconization reactor, a heat exchanger and a fractionation unit;

[0011] The fixed bed reactor is composed of a gas phase reaction zone, a gas-liquid separation zone, a gas-liquid countercurrent zone, a gas-liquid contact zone and a liquid phase reaction zone from top to bottom;

[0012] The gas-liquid separation zone is used to convert the materials entering the zone into gas-phase raw materials and liquid-phase raw materials; the gas-phase raw materials enter the gas-phase reaction zone, and the liquid-phase raw materials enter the gas-liquid countercurrent zone; the gas-phase components are generally fractions below 80°C, and the liquid-phase components are generally fractions above 80°C; the materials include a certain proportion of high-silicon naphtha raw materials and olefin-rich raw materials;

[0013] The gas phase reaction zone is used for low temperature hydrogenation and de-diolefination reaction of the gas phase components from the gas-liquid separation zone;

[0014] The gas-liquid countercurrent zone is used for further hydrogenation refining reaction of liquid raw materials, generally including the removal of silicic acid cyclosiloxanes, hydrogenation saturation of monoolefins and the removal of sulfur and nitrogen in naphtha fractions;

[0015] The gas-liquid contact zone is used for heat exchange between the hydrogen entering the zone and the liquid effluent from the gas-liquid countercurrent zone, and to mix and disperse the liquid effluent and hydrogen to form a hydrogen-dissolving flow; the gas-liquid contact zone is connected to the hydrogen source, and a mixed hydrogen-dissolving component or equipment is generally arranged in the gas-liquid contact zone, such as a membrane-tube hydrogen-dissolving component, a high-efficiency hydrogen mixer, a micro-bubble generator, and a bubble fractalizer, etc.;

[0016] The liquid phase reaction zone is used for deep hydrogenation refining reaction of the hydrogen-dissolved stream flowing out of the gas-liquid contact zone, and further deep desulfurization and dearomatization reactions;

[0017] The heat exchanger is used for heat exchange between the bottom effluent of the liquid phase reaction zone and the gas phase component from the gas phase reaction zone where dienes are removed;

[0018] The desiliconization reactor is used for the desiliconization reaction of the gas phase component for removing diolefins after heat exchange;

[0019] The fractionation unit is used for stripping fractionation of the gas phase effluent of the desiliconization reactor and the bottom effluent of the liquid phase reaction zone after heat exchange; it is generally in the form of a fractionation tower.

[0020] In the system of the present invention, a gas phase outlet is arranged at the top of the gas phase reaction zone, and a liquid phase outlet is arranged at the bottom of the liquid phase reaction zone. The effluent at the bottom of the liquid phase reaction zone passes through a heat exchanger and is connected to the outlet pipeline of the desiliconization reactor, and then enters the fractionation unit together.

[0021] The high-silicon naphtha processing technology of the present invention comprises the following contents:

[0022] (1) Hydrogen first enters the gas-liquid contact zone at the bottom of the fixed bed reactor, contacts and mixes with the liquid phase product flowing out from the bottom of the gas-liquid countercurrent zone, and flows upward through the gas-liquid countercurrent zone in sequence, and enters the gas phase reaction zone together with the gas phase components separated from the gas-liquid separation zone, and then flows out from the top of the fixed bed reactor with the gas phase components of the diene removal;

[0023] (2) High-silicon naphtha and olefin-rich feedstock enter the gas-liquid separation zone from the upper part of the fixed bed reactor and are converted into gas-phase feedstock and liquid-phase feedstock under certain conditions; the gas-phase feedstock enters the gas-phase reaction zone upward for hydrogenation and de-diolefination reaction; the liquid-phase feedstock enters the gas-liquid countercurrent zone downward for countercurrent contact with the upward hydrogen and undergoes a hydrogenation refining reaction; the liquid-phase effluent obtained by hydrogenation enters the gas-liquid contact zone for contact and heat exchange with hydrogen; the obtained hydrogen-dissolved stream enters the liquid-phase reaction zone for deep hydrogenation refining reaction; the bottom effluent of the liquid-phase reaction zone is heat exchanged with the gas-phase component flowing out of the top of the fixed bed reactor in a heat exchanger; the gas-phase component enters the desiliconization reactor for desiliconization; the desiliconized product is mixed with the bottom effluent of the liquid-phase reaction zone after heat exchange and subjected to stripping fractionation to obtain a qualified product.

[0024] In the process of the present invention, hydrogen and oil raw materials are fed at medium temperature (100°C~150°C) of the heating furnace when the operation is started. After the bed layer shows a significant temperature rise, the heating furnace is stopped and hydrogen and oil raw materials are fed at room temperature (5°C~30°C).

[0025] In the process of the present invention, the high-silicon naphtha has a silicon content of 1-3000 μg / g, preferably 5-800 μg / g, a sulfur content of ≯1000 μg / g, a nitrogen content of ≯200 μg / g, an olefin content of 1-30wt%, preferably 2-15wt%, and is generally at least one of straight-run naphtha, coker naphtha or hydrogenated reformed naphtha; at least one silicon-containing compound other than cyclic siloxane compounds and silane compounds exists in the high-silicon naphtha; wherein the cyclic siloxane compound is a cyclic silicon-oxygen bond (-Si-O-) n , n≥3 as the main chain of organic silicon compounds, such as those having the general formula (H 2 SiO) n , n≥3, or a compound in which H in the general formula is replaced by an alkyl group, a halogen, etc.; the silane compound is a compound having Si n H 2n+2 , an organosilicon compound of the general formula: n≥1.

[0026] Furthermore, the silicon-containing compound in the high-silicon naphtha includes but is not limited to at least one of alkylsilane, silanol or silane ether silicon-containing compounds.

[0027] Furthermore, as a specific embodiment, the silicon-containing compounds in the crude oil include but are not limited to at least one of tetramethylsilane, triethylsilane, tetraethylsilane, tetrapropylsilane, trimethylsilanol, dimethoxydimethylsilane, diethoxydimethylsilane, hexamethyldisiloxane, dimethyloctylchlorosilane, octamethyltrisiloxane and n-octyltriethoxysilane.

[0028] In the process of the present invention, the properties of the olefin-rich raw material are as follows: olefin content 5-60wt%, preferably 10-40wt%, sulfur content ≯15000μg / g, nitrogen content ≯400μg / g, polycyclic aromatic hydrocarbon content ≯50wt%. Generally, the olefin-rich raw material includes at least one gasoline and diesel fraction obtained by catalytic cracking, catalytic cracking, steam cracking, delayed coking, ebullated bed residue oil hydrogenation and other processes, and one or more of the gasoline or diesel fractions obtained by catalytic cracking, catalytic cracking, steam cracking, delayed coking, ebullated bed residue oil hydrogenation and other processes, preferably the content of coker gasoline is not less than 60wt%.

[0029] In the process of the present invention, the proportion of the olefin-rich feedstock to the total feedstock (high-silicon naphtha and olefin-rich feedstock) is ≮70wt%, preferably 40wt%~60wt%.

[0030] In the process of the present invention, the operating conditions of the gas-liquid separation zone are: pressure 1.0~10.0MPa, preferably 2.0~6.0MPa, wherein the hydrogen partial pressure accounts for 45%~80% of the total pressure; feed temperature 50~240℃, preferably 100~150℃.

[0031] In the process of the present invention, the operating conditions of the gas phase reaction zone are: pressure 0.1~6.0MPa, preferably 0.5~3.0MPa, wherein the hydrogen partial pressure accounts for 40%~70% of the total pressure; volume space velocity 0.1~6.0h -1 , preferably 0.5~3.0h -1 The reaction temperature is 80-200°C, preferably 120-160°C. The reaction zone is filled with a light distillate oil hydrogenation catalyst with low-temperature hydrogenation activity, the active metal is a metal oxide of Group VIB and / or a metal oxide of Group VIII, the carrier is alumina or modified alumina, based on the weight of the catalyst, the content of the metal oxide of Group VIB is 5%-30%, preferably 5%-15%, and the content of the metal oxide of Group VIII is 1%-15%, preferably 2%-6%; the specific surface area is 100-500m 2 / g, preferably 300-500m 2 / g, pore volume is 0.3-1.2mL / g, preferably 0.4-0.8mL / g. Based on the total catalyst loading in the fixed bed reactor, the catalyst loading volume ratio in the gas phase reaction zone is 1%~50%, preferably 5%~35%;

[0032] The gas phase reaction zone is used for low-temperature removal of dienes from the gas phase components from the gas-liquid separation zone. At this time, non-cyclic siloxanes in high-silica naphtha will also enter the gas phase reaction zone due to their lighter fractions. However, the inventors have found that their removal requires a higher reaction temperature, so this type of silicide does not react in the gas phase reaction zone.

[0033] In the process of the present invention, the operating conditions of the gas-liquid countercurrent zone are generally: pressure 1.0~8.0MPa, preferably 2.0~6.0MPa, wherein the hydrogen partial pressure accounts for 50%~90% of the total pressure; volume space velocity 0.1~10.0h -1 , preferably 0.5~5.0h -1 ; The reaction temperature is 200~400℃, preferably 220~300℃.

[0034] The gas-liquid countercurrent zone is filled with a graded silicon scavenger and a light distillate oil hydrogenation catalyst. The filling volume ratio of the two catalysts is 1:5~3:1. The active metal of the silicon scavenger is a metal oxide of Group VIB or / and a metal oxide of Group VIII, and the carrier is alumina or alumina modified by an additive. Based on the weight of the silicon scavenger, the content of the metal oxide of Group VIB is 5%-30%, preferably 5%-15%, and the content of the metal oxide of Group VIII is 1%-15%, preferably 2%-6%; the specific surface area is 100-500m 2 / g, preferably 300-500m 2 / g, pore volume is 0.3-1.2mL / g, preferably 0.4-0.8mL / g. Light distillate oil hydrogenation catalyst is well known to those skilled in the art, and is the same type as the catalyst loaded in the gas phase reaction zone, and the active metal content can be adjusted accordingly according to actual needs. The catalyst loading volume ratio in the gas-liquid countercurrent zone is 1%~60%, preferably 5%~40%;

[0035] In the process of the present invention, the operating conditions of the gas-liquid contact zone are 3.0-10.0 MPa, preferably 5.0-8.0 MPa, wherein the hydrogen partial pressure accounts for 100% of the total pressure; the feed temperature is 100-300°C, preferably 160-260°C at the initial start-up. After the device runs smoothly, the feed can be switched to normal temperature.

[0036] In the process of the present invention, the operating conditions of the liquid phase reaction zone are generally: pressure 1.0~8.0MPa, preferably 3.0~7.0MPa, this is a pure liquid phase reaction zone, the volume ratio of standard hydrogen to oil is 2~50, preferably 10~30; volume space velocity 0.1~6.0h -1 , preferably 0.5~3.0h -1; Reaction temperature 260~400℃, preferably 300~360℃. The catalyst loading volume ratio in the liquid phase reaction zone is 1%~80%, preferably 30%~60%. The liquid phase reaction zone is loaded with diesel hydrotreating catalyst, which is well known to those skilled in the art, wherein the active metal is a metal oxide of Group VIB or / and a metal oxide of Group VIII, based on the weight of the hydrogenation catalyst, the content of Group VIB metal oxide is 5%-30%, preferably 15%-25%, and the content of Group VIII metal oxide is 1%-15%, preferably 3%-8%; the carrier is alumina or additive-modified alumina, and the additive is one or more of B, P, Mg, Zr or Si, based on the weight of the carrier, the additive is 3%-15% in terms of oxide, preferably 3%-10%.

[0037] In the process of the present invention, the temperature of the gas phase component after heat exchange with the effluent of the liquid phase reaction zone is 200-340°C, and then enters the desiliconization reactor, in which a silicon scavenger is loaded. The type of silicon scavenger is the same as that of the gas-liquid countercurrent zone, and will not be described in detail. The operating conditions of the desiliconization reactor are generally: pressure 0.2-4.0MPa, preferably 1.0-3.0MPa, volume space velocity 0.1-4.0h -1 , preferably 0.2~2.0h -1 ; Reaction temperature 150~400℃, preferably 220~300℃.

[0038] Furthermore, according to the silicon content in the high-silicon naphtha oil, if the Si content is above 500 ppm, it is preferred to set up two desiliconization reactors for online switching operation.

[0039] The system and process of the present invention aims at the problem of difficulty in desiliconization of high-silicon naphtha and limited long-term operation of coking devices, and designs a low-energy consumption and low-cost process line, which can meet the processing of inferior oil products and provide qualified chemical raw materials and automotive diesel products. The advantages are as follows:

[0040] (1) The present invention uses the feed method, the gas-liquid separation zone, and the gas-liquid countercurrent zone to enrich the diolefins and high-silicon substances in the feed in the gas phase reaction zone for separate reactions, and separates the diolefins and monoolefins by fraction cutting, thereby avoiding the chain reaction caused by the large amount of saturation of monoolefins after the exothermic hydrogenation of the diolefins, which is conducive to controlling the catalyst bed temperature. In addition, compared with the conventional coking oil hydrogenation process, the setting of the pre-protection reactor to remove the diolefins can be omitted, simplifying the process flow.

[0041] (2) A process method in which high-silicon naphtha is mixed with coke oil (olefin-rich raw material). Since the silicides in high-silicon naphtha require a higher reaction temperature to be removed, the conditions for hydrogenation alone are harsh and the energy consumption is high. By processing it together with coke oil products, the heat of olefin saturation reaction in coke oil products can be fully utilized, saving energy and reducing consumption. However, if conventional hydrogenation technology is used, the larger bed temperature rise of coke oil hydrogenation mainly occurs in the middle and lower part of the reactor. If the high-silicon species can only be removed in the middle and lower part of the reactor, it will cover the surface of the main catalyst, causing it to be quickly poisoned and deactivated. The present invention utilizes the characteristic that the high-silicon component is enriched in the light component, effectively separates it through the setting of the partition reaction, and then increases the reaction temperature through heat exchange with the heavy component, and enters the gas phase reactor for separate hydrogenation. It not only utilizes the heat of the coke oil hydrogenation reaction, but also avoids the poisoning effect on the main catalyst caused by the co-reaction with the coke oil, thereby extending the operation cycle of the device. At the same time, the present invention only needs to activate the heating furnace to heat the feed in the initial stage of operation. After the catalyst bed temperature rises significantly, the heating furnace can be stopped to further save energy and reduce consumption.

[0042] (3) The present invention can better distribute and balance the heat of the three reaction zones and achieve coupling by setting up a gas-liquid countercurrent zone. In the present invention, the large amount of saturation of monoolefins mainly occurs in the gas-liquid countercurrent zone, which is also the reactor with the most obvious concentrated heat release. Through the countercurrent flow of the gas phase and the liquid phase in this area, a large amount of reaction heat can be quickly carried to the gas phase reaction zone and the liquid phase reaction zone, which not only meets the reaction temperature requirements of the two reaction zones, but also can better control the catalyst bed temperature in the gas-liquid countercurrent zone and delay the coking and deactivation of the catalyst surface. At the same time, the coupling effect of the three reaction zones set up in the present invention can make the reaction system more stable. While the gas-liquid countercurrent zone strengthens the mass transfer process, it is accompanied by backmixing, which requires stable pressure control. The uppermost gas phase reaction zone has a large compressible gas phase space, which has a good buffering effect on stabilizing the bed pressure and stabilizing the fluid flow state. By flexibly controlling the outlet gas volume, the gas phase flow rate and liquid layer thickness in the gas-liquid countercurrent zone can be adjusted, thereby increasing the gas velocity range for liquid flooding. The effluent from the gas-liquid countercurrent zone contacts pure hydrogen in the gas-liquid contact zone, undergoes heat exchange and cooling, and quickly enters the liquid phase reaction zone below, thus avoiding the escape of hydrogen after entering the reactor during the hydrogen mixing process of the conventional liquid phase hydrogenation process. At the same time, the liquid phase reaction zone below is a liquid phase space, which can well control the logistics state of the reactor outlet. If there is no such area, it will cause the problem of hydrogen being carried out of the reactor without reacting through the catalyst bed, which can eliminate the high-pressure separator in the conventional process flow and simplify the process. BRIEF DESCRIPTION OF THE DRAWINGS

[0043] Figure 1 The figure is a schematic diagram of the process flow of high-silicon naphtha processing of the present invention.

[0044] In the figure: 1-raw material, 2-hydrogen, 3-gas phase reaction zone, 4-gas-liquid separation zone, 5-gas-liquid countercurrent zone, 6-gas-liquid mixing zone, 7-liquid phase reaction zone, 8-gas phase reaction zone effluent, 9-liquid phase reaction zone effluent, 10-heat exchanger, 11-desiliconization reactor, 12-fractionation unit. DETAILED DESCRIPTION

[0045] The following non-limiting embodiments may enable a person skilled in the art to more fully understand the present invention, but are not intended to limit the present invention in any way.

[0046] The composition of the catalyst provided by the present invention can be characterized by combining inductively coupled plasma ICP and XRF energy spectrum to characterize the total content of the VIB group metal and the total content of the VIII group metal in the catalyst. The specific surface area and pore volume of the catalyst provided by the present invention are analyzed by nitrogen physical adsorption. The silicon content and type in the raw oil and the product of the present invention are analyzed by ICP and nuclear magnetic resonance, wherein the silicon content in the raw oil and the product is analyzed by ICP, and then the nuclear magnetic resonance 29Si MASNMR is used to analyze the silicon type in the raw oil and the product. According to nuclear magnetic resonance analysis, the silicon types in the high-silicon naphtha are tetramethylsilane, dimethoxydimethylsilane, and triethylsilane; the silicon types in the coker oil (olefin-rich raw material) are dimethylsiloxane, trimethylsiloxane, and tetramethylsiloxane.

[0047] by Figure 1 The implementation process of the processing technology of high-silicon naphtha of the present invention is described as follows: the reaction raw material 1 enters the hydrogenation reactor from the gas-liquid separation zone 4. It is separated into a gas phase and a liquid phase in the gas-liquid separation zone 4. The gas phase flows upward into the gas phase reaction zone 3, and the liquid phase flows downward into the gas-liquid countercurrent zone 5. Hydrogen 2 enters the hydrogenation reactor in the gas-liquid mixing zone 6, and after mixing and contacting with the liquid phase material flowing downward from the gas-liquid countercurrent zone 5, it flows upward into the gas-liquid countercurrent zone 5, and the liquid phase material carrying hydrogen flows downward into the liquid phase reaction zone 7.

[0048] A gas phase reaction occurs in the gas phase reaction zone 3, mainly saturating the dienes in the fraction below 80°C, generating the gas phase reaction zone effluent 8. A gas-liquid two-phase reaction occurs in the gas-liquid countercurrent zone 5, the liquid phase is the naphtha heavy component and the diesel fraction flowing downward, the gas phase is the hydrogen flowing upward, and the gas-liquid reverse contact occurs the hydrodesulfurization and desiliconization reaction of the naphtha fraction. The liquid phase logistics after the reaction in the gas-liquid countercurrent zone 5 flows downward into the gas-liquid contact zone 6, and after contacting and exchanging heat with hydrogen, enters the liquid phase reaction zone 7, where the liquid phase reaction occurs, mainly deep desulfurization and dearomatization of the diesel fraction. The liquid phase reaction zone 7 is a liquid phase reaction, and the deep desulfurization and dearomatization of the diesel fraction occurs. The liquid phase reaction zone effluent 9 flows out of the device and passes through the heat exchanger 10 to exchange heat with the gas phase reaction zone effluent 8, and then enters the desiliconization reactor 11. After desiliconization, it enters the subsequent fractionation unit 12 together with the liquid phase reaction zone effluent 9 to obtain qualified chemical raw materials naphtha and automotive diesel.

[0049] Embodiment 1~3

[0050] exist Figure 1 In the process flow, a 100mL fixed bed hydrogenation reactor, a heat exchanger, a desiliconization reactor and a fractionation unit are respectively set. The 100mL fixed bed hydrogenation reactor is set up with a gas phase reaction zone, a gas-liquid countercurrent zone, and a catalyst bed in the liquid phase reaction zone from top to bottom. 20mL of light distillate oil hydrogenation catalyst A is loaded in the gas phase reaction zone, and silicon scavenger B and light distillate oil hydrogenation catalyst A are loaded in the gas-liquid countercurrent zone, with the loading volumes of 10mL and 20mL respectively, and 50mL of diesel hydrofining catalyst C is loaded in the liquid phase reaction zone. 50mL of silicon scavenger B is loaded in the desiliconization reactor. High-silicon naphtha and coking gasoline and diesel mixed oil are used as raw materials, and the mass ratio of high-silicon naphtha, coking gasoline and coking diesel is 4:4:2. The properties of the catalyst are shown in Table 1, the properties of the raw oil are shown in Table 2, and the reaction process conditions and results are shown in Table 3.

[0051] Comparative Example 1

[0052] A conventional coking gasoline and diesel hydrogenation process was adopted, and a pre-protection reactor and a main reactor were respectively set in the process flow. The pre-protection reactor was filled with 20 mL of light distillate oil hydrogenation catalyst A, and the main reactor was filled with 10 mL of silicon trap B, 20 mL of light distillate oil hydrogenation catalyst A and 50 mL of diesel hydrogenation catalyst C. The properties of the raw materials were the same as those in the embodiment, the reaction conditions of the pre-protection reactor were the same as those of the gas phase hydrogenation reaction zone in embodiment 3, and the reaction conditions of the main reactor were the same as those of the liquid phase reaction zone in embodiment 3.

[0053] Comparative Example 2

[0054] The high-silicon naphtha, coking gasoline and diesel are hydrogenated separately.

[0055] High-silicon naphtha is hydrogenated and desiliconized in a gas phase reactor, and the reactor is filled with 50 mL of silicon trap B and 20 mL of light distillate oil hydrogenation catalyst A. The high-silicon naphtha raw material needs to be heated to the required reaction temperature in a heating furnace.

[0056] The coking gasoline and diesel hydrogenation adopts conventional process, and the pre-protection reactor and the main reactor are respectively set in the process flow. The pre-protection reactor is filled with 20mL of light distillate oil hydrogenation catalyst A, and the main reactor is filled with 10mL of silicon trap B and 50mL of diesel hydrogenation catalyst C. The raw material properties are the same as those in the embodiment, and the reaction conditions are the same as those in the liquid phase reaction zone of Example 3.

[0057] Table 1 Raw oil properties

[0058]

[0059] Table 2 Catalyst properties

[0060]

[0061] Table 3 Process conditions and results

[0062]

[0063] The evaluation results in Table 3 can illustrate that the high-silicon naphtha processing technology of the present invention can simultaneously process high-content silicides that are difficult to remove, control the silicon content in the refined naphtha to less than 1.0 μg / g, and better control and utilize the reaction heat, with a simple process and energy saving and consumption reduction. At the same time, there is no obvious silicon deposition phenomenon in the unloading of the device of the present invention, and there is no coking and carbon deposition phenomenon, which proves that it can be beneficial to the long-term operation of the device. When the conventional coking gasoline and diesel hydrogenation process is used to process high-silicon naphtha, the silicide can only be effectively removed in the lower part of the reactor, resulting in rapid deactivation of the main catalyst and difficulty in reaching the standard for the refined product. When the high-silicon naphtha and coking gasoline and diesel hydrogenation processes are used separately, the high-silicon naphtha desiliconization reactor needs to be equipped with a separate heating furnace, which has high energy consumption and cannot fully utilize the reaction heat of the coking oil product; the coking gasoline and diesel hydrogenation process needs to be equipped with a pre-protection reactor, and the problem of concentrated heat release due to olefin saturation is difficult to solve, and there is still a problem of short operation cycle.

Claims

1. A high-silicon naphtha processing technology, characterized in that The method includes the following contents: (1) hydrogen first enters the gas-liquid contact zone at the bottom of the fixed bed reactor, contacts and mixes with the liquid product flowing out from the bottom of the gas-liquid countercurrent zone, and flows upward through the gas-liquid countercurrent zone in sequence, and enters the gas-phase reaction zone together with the gas-phase raw material separated from the gas-liquid separation zone, and then flows out from the top of the fixed bed reactor with the gas-phase components of the de-diene; (2) high-silicon naphtha and olefin-rich raw materials enter the gas-liquid separation zone from the top of the fixed bed reactor, and are converted into gas-phase raw materials and liquid-phase raw materials under certain conditions; the gas-phase raw materials enter the gas-phase reaction zone upward to undergo hydrogenation de-diene reaction; The liquid raw material enters the gas-liquid countercurrent zone downward, contacts with the upward hydrogen countercurrent, and undergoes a hydrofining reaction. The liquid effluent obtained by hydrogenation enters the gas-liquid contact zone, contacts with the hydrogen for heat exchange, and the obtained hydrogen-dissolved stream enters the liquid reaction zone for deep hydrofining reaction. The gaseous component flowing out of the top of the fixed bed reactor exchanges heat with the bottom effluent of the liquid reaction zone in a heat exchanger, and then enters the desiliconization reactor for desiliconization. The desiliconized product is mixed with the bottom effluent of the liquid reaction zone after heat exchange, and subjected to stripping fractionation to obtain qualified products. The fixed bed reactor comprises a gas phase reaction zone, a gas-liquid separation zone, a gas-liquid countercurrent zone, a gas-liquid contact zone and a liquid phase reaction zone from top to bottom; the operating conditions of the gas-liquid separation zone are as follows: a pressure of 1.0-10.0 MPa, wherein the hydrogen partial pressure accounts for 45%-80% of the total pressure; a feed temperature of 50-240°C; the silicon content of the high-silicon naphtha is 1-3000 μg / g, the sulfur content is ≯1000 μg / g, the nitrogen content is ≯200 μg / g, and the olefin content is 1-30wt%; the properties of the olefin-rich raw material are as follows: the olefin content is 5-60wt%, the sulfur content is ≯15000 μg / g, the nitrogen content is ≯400 μg / g, and the polycyclic aromatic hydrocarbon content is ≯50wt%; a silicon scavenger and a light distillate oil hydrogenation catalyst are graded and loaded in the gas-liquid countercurrent zone; and the loading volume ratio of the two catalysts is 1:5-3:

1.

2. The processing technology according to claim 1, characterized in that: Hydrogen, high-silicon naphtha and olefin-rich raw materials are fed at a heating furnace at a temperature of 100~150℃ when the production starts. After the bed layer shows a significant temperature rise, the heating furnace is stopped and hydrogen, high-silicon naphtha and olefin-rich raw materials are fed at room temperature of 5℃~30℃.

3. The processing technology according to claim 1, characterized in that: The high-silicon naphtha is at least one of straight-run naphtha, coker naphtha or hydrogenated naphtha.

4. The processing technology according to claim 1, characterized in that: The high-silicon naphtha contains at least one silicon-containing compound other than cyclic siloxane compounds and silane compounds; wherein the cyclic siloxane compound is a compound with a cyclic silicon-oxygen bond (-Si-O-) n , n≥3 is an organic silicon compound with a main chain; the silane compound is an organic silicon compound with Si n H 2n+2 , an organosilicon compound of the general formula: n≥1.

5. The processing technology according to claim 1, characterized in that: The olefin-rich feedstock includes one or more of gasoline or diesel fractions obtained by catalytic cracking, catalytic cracking, steam cracking, delayed coking, and ebullated bed residue hydrogenation processes.

6. The processing technology according to claim 1, characterized in that: The content of coker gasoline in the olefin-rich feedstock is not less than 60wt%.

7. The processing technology according to claim 1, characterized in that: The proportion of olefin-rich raw materials to the total raw materials high-silicon naphtha and olefin-rich raw materials is ≮70wt%.

8. The processing technology according to claim 1, characterized in that: The operating conditions of the gas phase reaction zone are: pressure 0.1~6.0MPa, hydrogen partial pressure accounts for 40%~70% of the total pressure, volume space velocity 0.1~6.0h -1 , reaction temperature 80~200℃; the gas phase reaction zone is filled with a light distillate oil hydrogenation catalyst with low temperature hydrogenation activity, the active metal is a VIB group metal oxide and / or a VIII group metal oxide, the carrier is alumina or modified alumina, based on the weight of the catalyst, the VIB group metal oxide content is 5%-30%, the VIII group metal oxide content is 1%-15%; the specific surface area is 100-500m 2 / g, pore volume is 0.3-1.2mL / g; Based on the total catalyst loading in the fixed bed reactor, the catalyst loading volume ratio in the gas phase reaction zone is 1%~50%.

9. The processing technology according to claim 1, characterized in that: The operating conditions of the gas-liquid countercurrent zone are: pressure 1.0~8.0MPa, hydrogen partial pressure accounts for 50%~90% of the total pressure, volume space velocity 0.1~10.0h -1 , reaction temperature is 200~400℃; the catalyst loading volume ratio in the gas-liquid countercurrent zone is 1%~60%; the loading volume ratio of silicon scavenger and light distillate oil hydrogenation catalyst is 1:5~3:

1.

10. The processing technology according to claim 1, characterized in that: The active metal of the silicon scavenger is a metal oxide of Group VIB and / or a metal oxide of Group VIII, the carrier is alumina or alumina modified by an additive, the content of the metal oxide of Group VIB is 5%-30%, the content of the metal oxide of Group VIII is 1%-15% based on the weight of the silicon scavenger, and the specific surface area is 100-500m 2 / g, and the pore volume is 0.3-1.2mL / g.

11. The processing technology according to claim 1, characterized in that: The operating conditions of the gas-liquid contact zone are a pressure of 3.0-10.0 MPa, wherein the hydrogen partial pressure accounts for 100% of the total pressure; the feed temperature is 50-300°C at the initial start-up, and after the device runs smoothly, switch to normal temperature feed.

12. The processing technology according to claim 1, characterized in that: The operating conditions of the liquid phase reaction zone are: pressure 1.0~8.0MPa, this is a pure liquid phase reaction zone, the volume ratio of standard hydrogen to oil is 2~50, and the volume space velocity is 0.1~6.0h -1 , reaction temperature 260~400℃; the catalyst filling volume ratio in the liquid phase reaction zone is 1%~80%, the liquid phase reaction zone is filled with diesel hydrotreating catalyst, the active metal is VIB group metal oxide and / or VIII group metal oxide, based on the weight of the diesel hydrotreating catalyst, the VIB group metal oxide content is 5%-30%, and the VIII group metal oxide content is 1%-15%; the carrier is alumina or additive-modified alumina, the additive is one or more of B, P, Mg, Zr or Si, based on the weight of the carrier, the additive is 3%-15% in terms of oxide.

13. The processing technology according to claim 1, characterized in that: The temperature of the gas phase component after heat exchange with the effluent of the liquid phase reaction zone is 200~340℃, and then enters the desiliconization reactor. The desiliconization reactor is filled with a silicon scavenger. The type of silicon scavenger is the same as that of the gas-liquid countercurrent zone. The operating conditions of the desiliconization reactor are: pressure 0.2~4.0MPa, volume space velocity 0.1~4.0h -1 , reaction temperature 150~400℃.

14. The processing technology according to claim 1, characterized in that: When the Si content in high-silicon naphtha oil is above 500 ppm, two desiliconization reactors are set up for online switching operation.

15. A system for implementing the high-silicon naphtha processing process according to any one of claims 1 to 14, characterized in that It includes a fixed bed reactor, a desiliconization reactor, a heat exchanger and a fractionation unit; wherein the fixed bed reactor is composed of a gas phase reaction zone, a gas-liquid separation zone, a gas-liquid countercurrent zone, a gas-liquid contact zone and a liquid phase reaction zone from top to bottom; The gas-liquid separation zone is used to convert the materials entering the zone into gas-phase raw materials and liquid-phase raw materials; The gas phase reaction zone is used for low temperature hydrogenation and de-diolefination reaction of the gas phase raw material from the gas-liquid separation zone; The gas-liquid countercurrent zone is used for further hydrogenation refining reaction of liquid raw materials; The gas-liquid contact zone is used for the contact and heat exchange between the hydrogen entering the zone and the liquid effluent from the gas-liquid countercurrent zone, and for mixing and dispersing the liquid effluent and the hydrogen to form a hydrogen-dissolved flow; The liquid phase reaction zone is used for deep hydrogenation refining reaction of the hydrogen-dissolved stream flowing out of the gas-liquid contact zone; The heat exchanger is used for heat exchange between the bottom effluent of the liquid phase reaction zone and the gas phase component from the gas phase reaction zone where dienes are removed; The desiliconization reactor is used for the desiliconization reaction of the gas phase component of the diene removal after heat exchange; The fractionation unit is used for stripping and fractionating a mixture of the gas phase effluent of the desiliconization reactor and the bottom effluent of the liquid phase reaction zone after heat exchange.

16. The system according to claim 15, characterized in that: The gas phase raw material enters the gas phase reaction zone, and the liquid phase raw material enters the gas-liquid countercurrent zone; the gas phase raw material is a fraction below 80°C, and the liquid phase raw material is a fraction above 80°C.

17. The system according to claim 15, characterized in that: The gas-liquid contact area is connected with a hydrogen source, and a mixed hydrogen dissolving device is arranged in the gas-liquid contact area.

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