CN113563923A - Gasoline hydrogenation method and gasoline hydrogenation device - Google Patents

Gasoline hydrogenation method and gasoline hydrogenation device Download PDF

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Publication number
CN113563923A
CN113563923A CN202110954904.8A CN202110954904A CN113563923A CN 113563923 A CN113563923 A CN 113563923A CN 202110954904 A CN202110954904 A CN 202110954904A CN 113563923 A CN113563923 A CN 113563923A
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China
Prior art keywords
gas
phase hydrogenation
liquid
hydrogenation
hydrogen
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CN202110954904.8A
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Chinese (zh)
Inventor
李鹏程
李治
宋军超
陈新宇
高银福
刘文新
董晓伟
田进锋
马强
黄乐毅
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China Petroleum and Chemical Corp
Sinopec Engineering Group Co Ltd
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China Petroleum and Chemical Corp
Sinopec Engineering Group Co Ltd
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Priority to CN202110954904.8A priority Critical patent/CN113563923A/en
Publication of CN113563923A publication Critical patent/CN113563923A/en
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    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G7/00—Distillation of hydrocarbon oils
    • C10G7/02—Stabilising gasoline by removing gases by fractioning
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G45/00—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds
    • C10G45/02—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds to eliminate hetero atoms without changing the skeleton of the hydrocarbon involved and without cracking into lower boiling hydrocarbons; Hydrofinishing
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G49/00—Treatment of hydrocarbon oils, in the presence of hydrogen or hydrogen-generating compounds, not provided for in a single one of groups C10G45/02, C10G45/32, C10G45/44, C10G45/58 or C10G47/00
    • C10G49/002—Apparatus for fixed bed hydrotreatment processes
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G49/00—Treatment of hydrocarbon oils, in the presence of hydrogen or hydrogen-generating compounds, not provided for in a single one of groups C10G45/02, C10G45/32, C10G45/44, C10G45/58 or C10G47/00
    • C10G49/22—Separation of effluents
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G7/00—Distillation of hydrocarbon oils
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/10—Feedstock materials
    • C10G2300/1037—Hydrocarbon fractions
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/20—Characteristics of the feedstock or the products
    • C10G2300/201—Impurities
    • C10G2300/202—Heteroatoms content, i.e. S, N, O, P
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
    • C10G2400/02—Gasoline

Abstract

The invention discloses a gasoline hydrogenation method and a gasoline hydrogenation device, and relates to the technical field of chemical industry. A gasoline hydrogenation process by performing rectification, gas phase hydrogenation and liquid phase hydrogenation in a combined bed tower apparatus, comprising: mixing raw oil and hydrogen, and then introducing the mixture into combined bed tower equipment for rectification to obtain a gas-phase component and a liquid-phase component; the gas-phase components go upward in the combined bed tower type equipment to carry out gas-phase hydrogenation reaction, and the liquid-phase components go downward in the combined bed tower type equipment to carry out liquid-phase hydrogenation reaction. The gas-phase hydrogenation reaction has the advantages of low hydrogen consumption, low reaction heat and the like; the invention can complete the reaction without arranging a hydrogen circulation system, and the liquid phase hydrogenation reaction has the advantages of small bed temperature rise, high catalyst utilization rate, low investment cost and operation cost, and the like.

Description

Gasoline hydrogenation method and gasoline hydrogenation device
Technical Field
The invention relates to the technical field of chemical industry, in particular to a gasoline hydrogenation method and a gasoline hydrogenation device.
Background
At present, the rigidity requirement of gasoline in the market keeps relatively rapid growth speed, but with the increasing strictness of environmental protection requirements, the sulfur content in the gasoline needs to be strictly controlled so as to meet the requirements of producing clean fuel and controlling emission pollution, and the progress and development of hydrogenation technology is a key technology for solving the problem.
With the expansion of the scale of delayed coking units and ethylene processing, the scale and the scale of hydrotreaters such as coker gasoline and pyrolysis gasoline are increasing. The coking gasoline has high contents of olefin, dialkene, sulfur, heavy metal, coke powder and the like, and the problems of coking at the top of the reactor, pressure drop rise of a bed layer, a heat exchanger structure of a reaction part and the like easily occur during single hydrogenation, so that the operation period can be shortened. Pyrolysis gasoline hydrogenation is usually processed by a two-stage hydrogenation method, and the equipment investment is also large.
The existing hydrogenation process mainly has the following problems: (1) the investment and operation cost of the device are high, and more fractionating equipment and hydrogenation equipment exist; (2) the reaction conditions are harsh, and the reaction pressure required by the liquid phase hydrogenation method is high. Based on the above problems, it is the key point of the current oil hydrogenation technology upgrading to improve the oil hydrogenation process, reduce the investment and prolong the device operation period.
In addition, at present, domestic hydrofining is mainly carried out on a trickle bed, but the traditional trickle bed hydrogenation has a series of defects: (1) the mass transfer process of hydrogen to a liquid phase has the influence of a wetting factor; (2) the temperature gradient of the reactor is large, reaction hot spots and coking exist, and the service cycle of the catalyst is influenced; (3) the pressure drop in the reactor and the temperature difference between the bed layers are large, and side reactions such as hydrocracking and the like exist, so that the liquid yield is influenced; (4) the process flow is complex, and the investment of the device and the energy consumption of the operation are high.
In view of this, the invention is particularly proposed.
Disclosure of Invention
The invention aims to provide a gasoline hydrogenation method, aiming at realizing deep hydrofining, effectively reducing the sulfur content of a refined product and reducing the hydrogenation production cost.
Another object of the present invention is to provide a gasoline hydrogenation apparatus, which aims to reduce the investment and operation cost of the apparatus.
The invention is realized by the following steps:
in a first aspect, the present invention provides a gasoline hydrogenation process by performing rectification, gas phase hydrogenation and liquid phase hydrogenation in a combined bed tower apparatus, comprising: mixing raw oil and hydrogen, and then introducing the mixture into combined bed tower equipment for rectification to obtain a gas-phase component and a liquid-phase component; the gas-phase components go upward in the combined bed tower type equipment to carry out gas-phase hydrogenation reaction, and the liquid-phase components go downward in the combined bed tower type equipment to carry out liquid-phase hydrogenation reaction.
In an optional embodiment, the feeding temperature of the mixed raw oil and hydrogen entering the combined bed tower equipment is 220-350 ℃, and preferably 220-340 ℃.
In an optional embodiment, at least 1-3 layers of gas-phase hydrogenation catalyst beds are filled in the gas-phase hydrogenation reaction area, and 2-4 layers of liquid-phase hydrogenation catalyst beds are filled in the liquid-phase hydrogenation reaction area; preferably, hydrogen is replenished between two adjacent catalyst beds in the liquid phase hydrogenation reaction zone.
In an optional embodiment, the operation temperature of the combined bed tower type equipment is controlled to be 240-350 ℃, and the operation pressure is controlled to be 3-10 MPa; the hydrogen-oil ratio is 200-1000, and the volume space velocity is 0.5-5.0 h-1。
In an optional embodiment, a gas-phase hydrogenation product is output from the top of the combined bed tower type equipment, and the gas-phase hydrogenation product is subjected to gas-liquid separation to obtain an exhaust gas and a liquid-phase product; preferably, a side draw is performed in the region of the liquid phase hydrogenation reaction to collect the exit gas and output it mixed with the exit gas.
In an optional embodiment, a liquid-phase hydrogenation product is output from the bottom of the combined bed tower equipment, after the liquid-phase hydrogenation product exchanges heat with the raw oil, a part of the liquid-phase hydrogenation product is mixed with the raw oil and hydrogen to enter the combined bed tower equipment for rectification, and the rest of the liquid-phase hydrogenation product is mixed with the liquid-phase product and output.
In a second aspect, the present invention provides a gasoline hydrogenation apparatus for implementing the gasoline hydrogenation method according to any one of the foregoing embodiments, including a combined bed-tower apparatus, which includes, from top to bottom, a gas-phase hydrogenation reaction section, a rectification section and a liquid-phase hydrogenation reaction section in sequence, wherein the rectification section is provided with feed inlets for feeding raw oil and hydrogen into the tower, the gas-phase hydrogenation reaction section is filled with a gas-phase hydrogenation catalyst bed, and the liquid-phase hydrogenation reaction section is filled with a liquid-phase hydrogenation catalyst bed.
In an optional embodiment, the system further comprises a heat exchanger for heating the raw oil, a static mixer for mixing the raw oil and the hydrogen, and a hydrogen conveying pipeline for conveying the hydrogen, wherein a raw oil outlet of the heat exchanger is communicated with a feed inlet of the static mixer, and an outlet of the hydrogen conveying pipeline is respectively communicated with the feed inlet of the static mixer and a hydrogen inlet on the liquid-phase hydrogenation reaction section; the discharge hole of the static mixer is communicated with the feed inlet on the rectifying section.
In an optional embodiment, a heat source inlet and a heat source outlet for flowing in and out of a heat source are arranged on the heat exchanger, a liquid-phase hydrogenation product outlet is arranged at the bottom of the liquid-phase hydrogenation reaction section and is communicated with the heat source inlet of the heat exchanger, and a heat source outlet on the heat exchanger is respectively communicated with a first material conveying pipeline and a second material conveying pipeline so as to divide the liquid-phase hydrogenation product into two paths, wherein the outlet of the first material conveying pipeline is communicated with the steam stripping fractionation unit, and the outlet of the second material conveying pipeline is communicated with a feed inlet of the static mixer so as to utilize part of the liquid-phase hydrogenation product to flow back to the combined bed tower type equipment.
In an optional embodiment, the reactor further comprises a separator for gas-liquid separation, a gas-phase hydrogenation product outlet is arranged at the top of the gas-phase hydrogenation reaction section, the gas-phase hydrogenation product outlet is communicated with a feed inlet of the separator, and a liquid-phase outlet at the bottom of the separator is communicated with the first material conveying pipeline.
The invention has the following beneficial effects: the method has the advantages that the rectification, the gas-phase hydrogenation reaction and the liquid-phase hydrogenation reaction are carried out in a combined bed tower type device, the gas-phase hydrogenation reaction has the advantages of low hydrogen consumption, low reaction heat and the like, the sufficient dissolved hydrogen amount is achieved when raw oil is fed, and the reaction can be completed without adopting a hydrogen circulation system. Therefore, the invention can complete the reaction without arranging a hydrogen circulation system, and the liquid phase hydrogenation reaction has the advantages of small bed temperature rise, high catalyst utilization rate, low investment cost and operation cost, and the like.
Drawings
In order to more clearly illustrate the technical solutions of the embodiments of the present invention, the drawings needed to be used in the embodiments will be briefly described below, it should be understood that the following drawings only illustrate some embodiments of the present invention and therefore should not be considered as limiting the scope, and for those skilled in the art, other related drawings can be obtained according to the drawings without inventive efforts.
Fig. 1 is a schematic structural diagram of a gasoline hydrogenation apparatus provided in an embodiment of the present invention.
Description of the main element symbols: 100-a gasoline hydrogenation unit; 110-a combined bed-tower plant; 111-gas phase hydrogenation reaction section; 112-a rectification section; 113-a liquid phase hydrogenation reaction section; 120-feedstock pump; 130-a heat exchanger; 140-a static mixer; 150-a separator; 001-feedstock oil transfer line; 002-hydrogen transfer line; 003-a first material conveying pipeline; 004-second material conveying pipeline.
Detailed Description
In order to make the objects, technical solutions and advantages of the embodiments of the present invention clearer, the technical solutions in the embodiments of the present invention will be clearly and completely described below. The examples, in which specific conditions are not specified, were conducted under conventional conditions or conditions recommended by the manufacturer. The reagents or instruments used are not indicated by the manufacturer, and are all conventional products available commercially.
Referring to fig. 1, an embodiment of the present invention provides a gasoline hydrogenation method, which performs rectification, gas-phase hydrogenation and liquid-phase hydrogenation in a combined-bed tower apparatus, including: mixing raw oil and hydrogen, and then introducing the mixture into combined bed tower equipment for rectification to obtain a gas-phase component and a liquid-phase component; the gas-phase components go upward in the combined bed tower type equipment to carry out gas-phase hydrogenation reaction, and the liquid-phase components go downward in the combined bed tower type equipment to carry out liquid-phase hydrogenation reaction.
Specifically, the mixture of raw oil and hydrogen can be divided into gas-liquid two phases by rectification in a combined bed-tower type device, and hydrogenation reactions are respectively carried out above and below. The inventor creatively carries out the two-phase hydrogenation reaction in a tower type device, and the reaction can be completed on the premise of not arranging a hydrogen circulation system by depending on the sufficient hydrogen dissolving amount when raw oil is fed, and the liquid phase hydrogenation reaction has the advantages of small bed temperature rise, high catalyst utilization rate, low investment cost, low operation cost and the like.
Specifically, the raw oil is not limited, and may be distillate oil, secondary processed oil such as catalytic cracking and delayed coking, or oil produced by an ethylene cracking device.
In some embodiments, the feed temperature of the mixed raw oil and hydrogen entering the combined bed tower equipment is 220-350 ℃, preferably 220-340 ℃. Under the temperature condition, the preset gas-liquid ratio of rectification can be met.
In a preferred embodiment, the mixing of the feedstock oil and hydrogen gas is performed in a static mixer, which allows the hydrogen gas in the feedstock oil to reach the maximum hydrogen solubility of the current feedstock oil at that temperature and pressure.
Further, at least 1-3 layers of gas phase hydrogenation catalyst bed layers are filled in the gas phase hydrogenation reaction area, and 2-4 layers of liquid phase hydrogenation catalyst bed layers are filled in the liquid phase hydrogenation reaction area; generally, it is suitable to control the number of beds for the gas phase hydrogenation reaction to be smaller than the number of beds for the liquid phase hydrogenation reaction, for example, the number of beds for the gas phase hydrogenation reaction is 1 catalyst bed, and the number of beds for the liquid phase hydrogenation reaction is 3 catalyst beds.
Specifically, the types of catalysts filled in the gas-phase hydrogenation catalyst bed layer and the liquid-phase hydrogenation catalyst bed layer are slightly different, for example, the gas-phase hydrogenation catalyst bed layer is filled with conventional gas-phase hydrogenation catalysts such as FF-66 and RS-20, and the liquid-phase hydrogenation catalyst bed layer is filled with conventional liquid-phase hydrogenation catalysts such as FDS-1 and FUDS-8.
In some embodiments, hydrogen is supplemented between two adjacent catalyst beds in the liquid phase hydrogenation reaction zone, for example, in fig. 1, the liquid phase hydrogenation reaction is 3 catalyst beds, and hydrogen supplementation is performed between two adjacent beds at the middle and lower part of the tower, so as to meet the hydrogen requirement of the liquid phase hydrogenation reaction.
Further, the operation temperature of the combined bed tower type equipment is controlled to be 240-350 ℃, and the operation pressure is controlled to be 3-10 MPa. Specifically, the operation temperature is the approximate reaction temperature of the combined bed-tower apparatus, and may be 240 ℃, 250 ℃, 260 ℃, 270 ℃, 280 ℃, 290 ℃, 300 ℃, 310 ℃, 320 ℃, 330 ℃, 340 ℃, 350 ℃ or the like, or may be any value between the above adjacent temperature values; the operating pressure may be 3MPa, 4MPa, 5MPa, 6MPa, 7MPa, 8MPa, 9MPa, 10MPa, etc., or may be any value between the above adjacent pressure values.
Further, the hydrogen-oil ratio of the combined bed tower type equipment is controlled to be 200-1000, and the volume airspeed is 0.5-5.0 h-1. Specifically, the hydrogen-oil ratio is the ratio of the volume flow rate of the total amount of hydrogen to the flow rate of the raw material oil, and may be 200300, 400, 500, 600, 700, 800, 900, 1000, etc.; the volume space velocity can be 0.5h-1、1.0h-1、2.0h-1、3.0h-1、4.0h-1、5.0h-1And the like.
Further, a gas-phase hydrogenation product is output from the tower top of the combined bed tower type equipment, and the gas-phase hydrogenation product is subjected to gas-liquid separation to obtain a discharge gas and a liquid-phase product. In some embodiments, a side line collection is performed in the liquid phase hydrogenation reaction area to collect the discharged gas and output the gas by mixing with the discharged gas, wherein the gas collected by the side line is mainly hydrogen sulfide, hydrogen and other gases.
Further, a liquid phase hydrogenation product is output from the tower bottom of the combined bed tower type equipment, after the liquid phase hydrogenation product exchanges heat with the raw oil, part of the liquid phase hydrogenation product is mixed with the raw oil and hydrogen to enter the combined bed tower type equipment for rectification, and the rest of the liquid phase hydrogenation product is mixed with the liquid phase product and output. The energy utilization rate is improved by utilizing the heat exchange between the liquid-phase hydrogenation product and the raw oil, partial liquid-phase hydrogenation product flows back to the combined bed tower type equipment, the liquid-phase circulating hydrogenation technology is adopted, a hydrogen circulating system is not arranged in the technology, and the hydrogen required by the hydrogenation reaction of the fresh raw material is provided by depending on the dissolved hydrogen carried into the reaction system when the liquid-phase product is circulated in a large quantity.
The embodiment of the invention provides a gasoline hydrogenation device 100 which comprises a combined bed tower type device 110, wherein the combined bed tower type device 110 sequentially comprises a gas phase hydrogenation reaction section 111, a rectification section 112 and a liquid phase hydrogenation reaction section 113 from top to bottom, the rectification section 112 is provided with a feed inlet for inputting raw oil and hydrogen into a tower, the gas phase hydrogenation reaction section 111 is filled with a gas phase hydrogenation catalyst bed layer, and the liquid phase hydrogenation reaction section 113 is filled with a liquid phase hydrogenation catalyst bed layer.
In some embodiments, the rectifying section 112 is located at the middle upper part of the combined bed tower apparatus 110, and a set of reactive rectifying combined bed system is formed by the gas phase hydrogenation reaction section 111, the rectifying section 112 and the liquid phase hydrogenation reaction section 113, so that the gas phase hydrogenation reaction is performed in the top region, and the liquid phase hydrogenation reaction is performed in the bottom region.
In some embodiments, the gasoline hydrogenation apparatus 100 further comprises a heat exchanger 130 for heating the feedstock oil, a static mixer 140 for mixing the feedstock oil and the hydrogen, and a hydrogen conveying line 002 for conveying the hydrogen, wherein a feedstock oil outlet of the heat exchanger 130 is communicated with a feed inlet of the static mixer 140, and an outlet of the hydrogen conveying line 002 is respectively communicated with a feed inlet of the static mixer 140 and a hydrogen inlet on the liquid-phase hydrogenation reaction section 113; the discharge port of the static mixer 140 is communicated with the feed port on the rectifying section 112. The raw material oil conveying pipeline 001 conveys the raw material to the raw material pump, the raw material oil is conveyed to the heat exchanger 130 through the raw material pump 120 to be heated, and then the raw material oil enters the static mixer 140 to be uniformly mixed with hydrogen and then enters the rectifying section 112 to be rectified.
Further, a heat source inlet and a heat source outlet for flowing in and out of a heat source are arranged on the heat exchanger 130, a liquid-phase hydrogenation product outlet is arranged at the bottom of the liquid-phase hydrogenation reaction section 113, the liquid-phase hydrogenation product outlet is communicated with the heat source inlet of the heat exchanger 130, the heat source outlet of the heat exchanger 130 is respectively communicated with the first material conveying pipeline 003 and the second material conveying pipeline 004 to divide the liquid-phase hydrogenation product into two paths, wherein the outlet of the first material conveying pipeline 003 is communicated with the steam stripping fractionation unit, and the outlet of the second material conveying pipeline 004 is communicated with the feed inlet of the static mixer 140 to reflux part of the liquid-phase hydrogenation product to the combined bed tower device 110. Utilize liquid phase hydrogenation product as the heat source of heat exchanger 130 heating for heating raw oil, liquid phase hydrogenation product after the cooling divide into two the tunnel, carry to the steam stripping fractionating unit through first defeated material pipeline 003, realize liquid phase circulation hydrogenation (SRH) through second defeated material pipeline 004 backward flow to static mixer 140, can be under the prerequisite that does not set up hydrogen circulation system, carry into reaction system's dissolved hydrogen when relying on liquid phase product to circulate in a large number and provide the hydrogen that fresh raw materials carried out the hydrogenation and needed. Thus, the liquid phase hydrogenation part has the advantages of small bed temperature rise, high catalyst utilization rate, low investment cost and operation cost, and the like.
In some embodiments, the gasoline hydrogenation apparatus 100 further comprises a separator 150 for gas-liquid separation, the top of the gas-phase hydrogenation reaction section 111 is provided with a gas-phase hydrogenation product outlet, the gas-phase hydrogenation product outlet is communicated with the feed inlet of the separator 150, and the bottom liquid-phase outlet of the separator 150 is communicated with the first material conveying pipeline 003. The separator 150 performs gas-liquid separation on the gas-phase hydrogenation product, and is divided into a top outlet and a bottom outlet, wherein the top outlet is a gas-phase outlet and is used as exhaust gas for output, and the bottom outlet is a liquid-phase outlet and enters the stripping fractionation unit after being communicated with the first material conveying pipeline 003.
Compared with the prior art, the gasoline hydrogenation process provided by the embodiment of the invention has the following advantages:
(1) the embodiment of the invention combines the rectification part and the light and heavy component hydrodesulfurization and olefin removal reaction part in a tower to form a set of reaction rectification combined bed system, thereby reducing the occupied area and equipment investment and lowering the equipment maintenance cost.
(2) The liquid phase hydrogenation part enables the amount of hydrogen dissolved in the raw material to be far higher than the amount of hydrogen required by hydrogenation reaction through the circulation of liquid products, so that a hydrogen circulating compressor with high price is not required to be used in the device, and the investment of the device is greatly reduced. In addition, high-pressure equipment such as a hot high-pressure separator, a cold high-pressure separator and the like is also eliminated, the running safety risk of the device is reduced, and the safety performance is obviously improved.
(3) The liquid phase hydrogenation part has good reaction effect and can still realize the purpose of deep hydrofining at higher airspeed.
(4) Compared with the traditional hydrogenation technology, the liquid phase hydrogenation technology has lower operation energy consumption and reduces the hydrogenation production cost.
(5) The temperature gradient of the reactor is greatly reduced, the operation is close to isothermal operation, the phenomena of local overheating and bed layer temperature runaway are avoided, reaction hot spots are eliminated, carbon deposition is reduced, and the service life of the catalyst is prolonged.
The features and properties of the present invention are described in further detail below with reference to examples.
The following example uses the gasoline hydrogenation apparatus 100 in fig. 1 to perform the reaction, and specifically includes the following steps:
the raw oil is conveyed to the heat exchanger 130 by the raw material pump 120 to be heated, then the raw oil is mixed with the hydrogen conveyed by the hydrogen conveying pipeline 002 in the static mixer 140, and then the mixture enters the rectifying section 112 to be rectified into gas-phase components and liquid-phase components, wherein the gas-phase components are subjected to hydrogenation reaction in the gas-phase hydrogenation reaction section 111, and the liquid-phase components are subjected to hydrogenation reaction in the liquid-phase hydrogenation reaction section 113.
The hydrogen is split into three streams, one stream of hydrogen enters the static mixer 140, and the other two streams are used for supplying hydrogen to the liquid phase hydrogenation reaction section 113.
The liquid phase hydrogenation product output from the tower bottom of the combined bed tower type equipment 110 is used as a heat source heated by the heat exchanger 130 for heating raw oil, the liquid phase hydrogenation product after being cooled is divided into two paths, and the two paths are conveyed to a steam stripping fractionation unit through a first material conveying pipeline 003 and flow back to the static mixer 140 through a second material conveying pipeline 004 to realize liquid phase circulating hydrogenation (SRH).
The gas phase hydrogenation product output from the top of the combined bed tower type equipment 110 enters a separator 150 for gas-liquid separation, a gas phase outlet is output as exhaust gas, and a liquid phase outlet at the bottom is communicated with a first material conveying pipeline 003 and then enters a stripping fractionation unit. Side line collection is carried out in the liquid phase hydrogenation reaction section to collect the discharged gas and mix the discharged gas with the discharged gas for output.
Example 1
The embodiment provides a gasoline hydrogenation method, and the specific parameters are as follows: cracking gasoline with 350 mug/g of sulfur and 29 mug/g of nitrogen is used as a raw material, the reaction temperature is controlled to be 250 ℃, the reaction pressure is controlled to be 3MPa, the hydrogen-oil ratio is controlled to be 600, and the volume space velocity is controlled to be 1.0h-1。
The sulfur content of the refined pyrolysis gasoline obtained in the embodiment is less than 10 mug/g through detection.
Example 2
The embodiment provides a gasoline hydrogenation method, and the specific parameters are as follows: the coking gasoline with the sulfur content of 524 mu g/g and the nitrogen content of 117 mu g/g is used as a raw material, the reaction temperature is controlled to be 260 ℃, the reaction pressure is controlled to be 6MPa, the hydrogen-oil ratio is controlled to be 800, and the volume space velocity is controlled to be 1.0h-1。
Through detection, the sulfur content of the refined coker gasoline obtained in the embodiment is less than 10 mug/g.
Comparative example 1
This comparative example was carried out by conventional trickle bed hydrogenation, and the feed and reaction conditions were as in example 1.
The sulfur content of the refined pyrolysis gasoline obtained by the comparative example is 21 mug/g and is difficult to be reduced to below 10 mug/g.
Comparative example 2
This comparative example was carried out by conventional trickle bed hydrogenation, and the feed and reaction conditions were as in example 2.
The detection proves that the sulfur content of the refined coker gasoline obtained by the comparative example is 17 mug/g, and is difficult to reduce to below 10 mug/g.
Comparative example 3
The only difference from example 1 is that the reaction temperature is reduced to 200 ℃ and the remaining parameters are the same. The sulfur content in the refined pyrolysis gasoline obtained after hydrogenation is 309 mug/g, which is far larger than the standard 10 mug/g of gasoline products.
Comparative example 4
The only difference from example 1 is that the reaction temperature is increased to 370 ℃ and the remaining parameters are the same. After a period of testing, the pressure drop across the bed increases and the catalyst activity decreases until it becomes inactive.
Comparative example 5
The difference from example 2 is only that the hydrogen-oil ratio is reduced to 180, and the remaining parameters are the same. The sulfur content of the refined coking gasoline obtained after hydrogenation is 112 mug/g, which is far larger than the standard 10 mug/g of gasoline products.
Comparative example 6
The difference from example 2 is only that the hydrogen to oil ratio is increased to 1200 and the remaining parameters are the same. The sulfur content in the refined coking gasoline obtained after hydrogenation is 14 mu g/g which is more than 10 mu g/g of the standard of gasoline products.
Comparative example 7
The difference from example 2 is only that the pressure is reduced to 2MPa, the remaining parameters being the same. The sulfur content of the refined coking gasoline obtained after hydrogenation is far greater than the standard 10 mug/g of gasoline products.
Comparative example 8
The only difference from example 2 is that the pressure is increased to 12MPa, the remaining parameters being the same. The hydrodesulfurization effect of the coker gasoline is similar to that of example 2, but the increase in pressure leads to poor economic benefit.
Test example 1
When the combined bed reactor is stably operated for 1000 hours under the condition of the example 1, the gasoline hydrogenation effect is not attenuated, and the combined bed reactor catalyst is proved to have longer service cycle and better cost performance. The catalyst in the combined bed reactor is compared with the catalyst in a trickle bed in the same running time, and the carbon on the surface of the catalyst in the combined bed reactor is slightly less.
The above is only a preferred embodiment of the present invention, and is not intended to limit the present invention, and various modifications and changes will occur to those skilled in the art. Any modification, equivalent replacement, or improvement made within the spirit and principle of the present invention should be included in the protection scope of the present invention.

Claims (10)

1. A gasoline hydrogenation process by performing rectification, gas phase hydrogenation and liquid phase hydrogenation in a combined bed tower apparatus comprising:
mixing raw oil and hydrogen, and then introducing the mixture into the combined bed tower equipment for rectification to obtain a gas-phase component and a liquid-phase component; the gas-phase components go upward in the combined bed tower type equipment to carry out gas-phase hydrogenation reaction, and the liquid-phase components go downward in the combined bed tower type equipment to carry out liquid-phase hydrogenation reaction.
2. The gasoline hydrogenation method according to claim 1, wherein the feed temperature of the mixed raw oil and hydrogen entering the combined bed tower type equipment is 220-350 ℃, preferably 220-340 ℃.
3. The gasoline hydrogenation method of claim 1 or 2, wherein at least 1-3 gas phase hydrogenation catalyst beds are filled in the gas phase hydrogenation reaction zone, and 2-4 liquid phase hydrogenation catalyst beds are filled in the liquid phase hydrogenation reaction zone;
preferably, hydrogen is supplemented between two adjacent catalyst beds in the liquid phase hydrogenation reaction zone.
4. The gasoline hydrogenation method according to claim 3, wherein the operating temperature of the combined bed tower type equipment is controlled to be 240-350 ℃, and the operating pressure is controlled to be 3-10 MPa;
preferably, the hydrogen-oil ratio is 200-1000, and the volume space velocity is 0.5-5.0 h-1。
5. The gasoline hydrogenation method of claim 1, wherein a gas phase hydrogenation product is output from the top of the combined bed tower type equipment, and the gas phase hydrogenation product is subjected to gas-liquid separation to obtain an exhaust gas and a liquid phase product;
preferably, a side line collection is carried out in the region of the liquid phase hydrogenation reaction to collect the discharged gas and output the gas mixed with the discharged gas.
6. The gasoline hydrogenation method according to claim 5, wherein a liquid phase hydrogenation product is output from the bottom of the combined bed tower type equipment, after the liquid phase hydrogenation product exchanges heat with the raw oil, a part of the liquid phase hydrogenation product is mixed with the raw oil and hydrogen to enter the combined bed tower type equipment for rectification, and the rest of the liquid phase hydrogenation product is mixed with the liquid phase product and output.
7. A gasoline hydrogenation device for implementing the gasoline hydrogenation method according to any one of claims 1 to 6, which comprises a combined bed tower type device, wherein the combined bed tower type device sequentially comprises a gas phase hydrogenation reaction section, a rectification section and a liquid phase hydrogenation reaction section from top to bottom, the rectification section is provided with feed inlets for inputting raw oil and hydrogen into the tower, the gas phase hydrogenation reaction section is filled with a gas phase hydrogenation catalyst bed layer, and the liquid phase hydrogenation reaction section is filled with a liquid phase hydrogenation catalyst bed layer.
8. The gasoline hydrogenation device according to claim 7, further comprising a heat exchanger for heating the raw oil, a static mixer for mixing the raw oil and the hydrogen, and a hydrogen conveying pipeline for conveying the hydrogen, wherein a raw oil outlet of the heat exchanger is communicated with a feed inlet of the static mixer, and an outlet of the hydrogen conveying pipeline is respectively communicated with the feed inlet of the static mixer and a hydrogen inlet on the liquid-phase hydrogenation reaction section; and the discharge hole of the static mixer is communicated with the feed inlet on the rectifying section.
9. The gasoline hydrogenation device according to claim 8, wherein the heat exchanger is provided with a heat source inlet and a heat source outlet for flowing in and out of a heat source, the bottom of the liquid phase hydrogenation reaction section is provided with a liquid phase hydrogenation product outlet, the liquid phase hydrogenation product outlet is communicated with the heat source inlet of the heat exchanger, the heat source outlet of the heat exchanger is respectively communicated with a first material conveying pipeline and a second material conveying pipeline so as to divide the liquid phase hydrogenation product into two paths, the outlet of the first material conveying pipeline is communicated with the stripping and fractionating unit, and the outlet of the second material conveying pipeline is communicated with the feed inlet of the static mixer so as to utilize part of the liquid phase hydrogenation product to flow back to the combined bed tower type equipment.
10. The gasoline hydrogenation device according to claim 9, further comprising a separator for gas-liquid separation, wherein a gas-phase hydrogenation product outlet is arranged at the top of the gas-phase hydrogenation reaction section, the gas-phase hydrogenation product outlet is communicated with the feed inlet of the separator, and a liquid-phase outlet at the bottom of the separator is communicated with the first material conveying pipeline.
CN202110954904.8A 2021-08-19 2021-08-19 Gasoline hydrogenation method and gasoline hydrogenation device Pending CN113563923A (en)

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Cited By (1)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN115404100A (en) * 2022-09-28 2022-11-29 石家庄睿途生物科技有限公司 Hydrogenation device and hydrogenation method for coupling fluidized bed and fixed bed

Citations (4)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN1362480A (en) * 2001-01-05 2002-08-07 中国石油化工股份有限公司 Hydrocarbon hydroprocessing technique
EP2199368A1 (en) * 2007-09-27 2010-06-23 Nippon Steel Engineering Co., Ltd. Bubble-column-type hydrocarbon synthesis reactor, and hydrocarbon synthesis reaction system having the reactor
CN107400534A (en) * 2016-05-20 2017-11-28 中石化洛阳工程有限公司 A kind of Clean Fuel Production method
CN109777497A (en) * 2017-11-14 2019-05-21 中国石油化工股份有限公司 Refinery gas's combined hydrogenation method

Patent Citations (4)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN1362480A (en) * 2001-01-05 2002-08-07 中国石油化工股份有限公司 Hydrocarbon hydroprocessing technique
EP2199368A1 (en) * 2007-09-27 2010-06-23 Nippon Steel Engineering Co., Ltd. Bubble-column-type hydrocarbon synthesis reactor, and hydrocarbon synthesis reaction system having the reactor
CN107400534A (en) * 2016-05-20 2017-11-28 中石化洛阳工程有限公司 A kind of Clean Fuel Production method
CN109777497A (en) * 2017-11-14 2019-05-21 中国石油化工股份有限公司 Refinery gas's combined hydrogenation method

Cited By (2)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN115404100A (en) * 2022-09-28 2022-11-29 石家庄睿途生物科技有限公司 Hydrogenation device and hydrogenation method for coupling fluidized bed and fixed bed
CN115404100B (en) * 2022-09-28 2023-10-24 石家庄睿途生物科技有限公司 Fluidized bed and fixed bed coupled hydrogenation device and hydrogenation method

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