Method and system for producing chemical material from diesel oil and / or light wax oil

Abstract

The invention relates to a method and a system for producing a chemical material from diesel oil and / or light wax oil. The method comprises the following steps: introducing raw oil into a reactive distillation tower from a raw oil inlet for contact reaction with hydrogen; the raw oil is diesel oil and / or light wax oil; wherein the reactive distillation tower comprises an upper distillation section and a lower stripping section which are in fluid communication, and the raw oil inlet is positioned between the distillation section and the stripping section; the rectifying section is sequentially provided with a first hydrofining catalyst and a first hydrocracking catalyst from top to bottom; the stripping section is sequentially provided with a second hydrofining catalyst, a second hydrocracking catalyst and a third hydrofining catalyst from top to bottom. According to the method disclosed by the invention, firstly, high-quality chemical materials can be efficiently produced, secondly, secondary cracking of naphtha fractions can be effectively prevented so as to reduce the production of liquefied gas, the normal / isomeric ratio in light naphtha is increased, and an ethylene cracking raw material is optimized.

Description

Technical Field

[0001] This application relates to the field of petroleum processing, and more specifically, to a method and system for producing chemical feedstocks from diesel oil and / or light wax oil. Background Technology

[0002] Straight-run diesel fuel, with its high alkanes content, is a potential high-quality feedstock for ethylene cracking. Diesel hydrocracking, acting as a bridge between refining and chemical processes, can convert aromatics and cycloalkanes from straight-run diesel into naphtha fractions, enriching the alkanes within the diesel fractions and making them high-quality feedstocks for ethylene cracking. Specifically, heavy naphtha fractions, with their high aromatic potential, can serve as excellent feedstocks for catalytic reforming; light naphtha fractions, with their high alkanes content, can also be used as high-quality feedstocks for ethylene cracking. The development of this technology can achieve the goal of "reducing oil consumption and increasing chemical production," realizing the efficient utilization of diesel resources and possessing significant application value.

[0003] Traditional single-stage, one-pass fixed-bed hydrocracking units primarily produce light naphtha, heavy naphtha, diesel oil, unconverted oil, and a small amount of liquefied petroleum gas (LPG). However, due to the characteristics of the hydrocracking process, hydrocarbon isomerization occurs during the reaction, resulting in a high content of isomeric hydrocarbons in light naphtha. Compared to normal hydrocarbons, the proportion of olefins obtained from the thermal cracking of isomeric hydrocarbons into ethylene decreases significantly. Therefore, reducing the proportion of isomeric hydrocarbons in light naphtha to increase olefin yield is one of the problems that needs to be solved. Secondly, traditional diesel hydrocracking units, due to their relatively low pressure ratings, have stringent requirements for the nitrogen mass fraction of straight-run diesel or light wax oil feedstocks (typically not exceeding 200 μg / g), thus limiting the range of feedstocks for chemical production. Increasing the range of nitrogen mass fraction in existing diesel or light wax oil hydrocracking units is another problem that needs to be addressed. Third, the unconverted oil obtained from straight-run diesel or light wax oil has a high content of alkanes and is a high-quality feedstock for ethylene cracking. However, during the conversion of diesel and light wax oil, the unconverted oil will inevitably undergo secondary cracking, which will cause light naphtha and heavy naphtha and other chemical feedstocks to be further cracked into light hydrocarbons. How to control the yield of light hydrocarbons, that is, improve the cracking selectivity, is the third problem that diesel or light wax oil needs to solve.

[0004] Methods for improving the quality of diesel fuel have been disclosed in medium-pressure hydrocracking or medium-pressure hydrotreating, but these methods mainly focus on how to increase the cetane index of diesel fuel, with less emphasis on the production of chemical feedstocks.

[0005] Patent ZL200810166852.2 discloses a low-cost, high-volume hydrocracking method for producing chemical feedstocks. Feedstock oil and hydrogen are mixed and fed into a first hydrocracking reaction zone for hydrocracking. The cracked oil is separated and fractionated, with the diesel fraction reacting in a second hydrocracking zone at a pressure of 1.0 MPa-5.5 MPa. This method can yield chemical feedstocks with a purity of over 98%, including liquefied petroleum gas (LPG), light naphtha, heavy naphtha, and tail oil. The aforementioned patent primarily focuses on the recycling and conversion of diesel fractions.

[0006] US6547956B1 discloses a process for hydrocracking and hydrotreating vacuum fractions and other fractions. The goal of this process is to reduce hydrogen consumption and improve reactor efficiency (by reducing catalyst usage). The process includes a hydrotreating reaction zone where the distillate oil undergoes a certain degree of hydrodenitrification, desulfurization, and aromatics saturation reaction, followed by partial hydrocracking to produce light components, diesel fuel, and unconverted oil. The hydrotreating product oil enters a separation unit, where the heavier unconverted oil does not require further deep hydrotreating to remove sulfur, nitrogen, and saturated aromatics, but instead enters the fractionation system. The diesel fraction and light components obtained from the separation unit then enter a post-processing reaction zone for further hydrotreating to improve product quality.

[0007] In the aforementioned existing technologies, the initial first-stage hydrocracking process only involves partial hydrogenation conversion, while the subsequently separated light components undergo further deep hydrogenation in the post-processing stage to meet the property requirements of the target product. Overall, these technologies are all traditional single-stage or two-stage hydrocracking processes, which improve product properties, reduce energy consumption, and enhance the selectivity of the target product by selecting appropriate catalysts.

[0008] Meanwhile, there is also a lot of research in the field of catalytic distillation, but most of the research is mainly in the fields of aromatic conversion, catalytic desulfurization, and catalytic cracking diesel conversion. There is no relevant content on its use in straight-run or light wax oil hydrocracking to produce high-quality chemical feedstocks. Summary of the Invention

[0009] The purpose of this disclosure is to provide a method and system for producing chemical feedstocks from diesel oil and / or light wax oil. The method of this disclosure can firstly produce high-quality chemical feedstocks efficiently, secondly effectively prevent secondary cracking of naphtha fractions to reduce liquefied petroleum gas production, and thirdly improve the positive-to-negative ratio in light naphtha to optimize ethylene cracking feedstocks.

[0010] To achieve the above objectives, the first aspect of this disclosure provides a method for producing chemical feedstocks from diesel oil and / or light wax oil, the method comprising: introducing feedstock oil from a feedstock oil inlet into a reactive distillation column to react with hydrogen; wherein the feedstock oil is diesel oil and / or light wax oil; The reactive distillation column includes an upper rectification section and a lower stripping section that are fluidly connected, with the feed oil inlet located between the rectification section and the stripping section; the rectification section is provided with a first hydrorefining catalyst and a first hydrocracking catalyst from top to bottom; the stripping section is provided with a second hydrorefining catalyst, a second hydrocracking catalyst and a third hydrorefining catalyst from top to bottom.

[0011] Optionally, the feedstock oil is reacted with hydrogen in a countercurrent manner within the reactive distillation column.

[0012] Optionally, the method further includes: extracting heavy components from the lower part of the rectification section and introducing the heavy components from the bottom of the stripping section into the reactive distillation apparatus.

[0013] Optionally, the initial boiling point of the heavy component is 150-180℃, and the final boiling point is 220-260℃; the circulation rate of the heavy component is 5-30% by weight of the diesel or light wax oil feed rate.

[0014] Optionally, the reactive distillation column has a top pressure of 4-8 MPa, a top reflux ratio of 0.3-3.0, and a hydrogen-to-oil volume ratio of 300-3000 Nm³. 3 / m 3 The bottom temperature of the column is 300-430℃; preferably, the top pressure is 4-8 MPa, the reflux ratio at the top is 0.5-2.0, and the hydrogen-to-oil volume ratio is 600-1800 Nm³. 3 / m 3 The temperature at the bottom of the tower is 350-410℃; The volume hourly space velocity (VHSV) of the first hydrorefining catalyst in the rectification section is 1-20 h⁻¹. -1 Preferably 2-15h -1 The volume hourly space velocity of the first hydrocracking catalyst is 1-20 h⁻¹. -1 Preferably 2-10h -1 ; The volume hourly space velocity (VHSV) of the second hydrorefining catalyst in the stripping section is 0.2-10.0 h⁻¹. -1 Preferably 1.0-4.0h -1 The volume hourly space velocity (VHSV) of the second hydrocracking catalyst is 0.2-10.0 h⁻¹. -1 Preferably 1.5-4.5h -1 The volume hourly space velocity of the third hydrorefining catalyst is 10-20 h⁻¹. -1 Preferably 10-15h -1 .

[0015] Optionally, the method further includes: drawing out the gas phase from the top of the reactive distillation column, cooling and separating the gas and liquid phases, returning a portion of the separated liquid phase to the top of the reactive distillation column, and introducing the hydrogen obtained after gas phase deposition into the reactive distillation column as feedstock hydrogen.

[0016] Optionally, based on the dry weight of the first hydrocracking catalyst, the first hydrocracking catalyst contains 11-50% by weight of a first metal component, calculated as oxides, and the balance being a first support; the first support comprises a first zeolite and an inert support in a weight ratio of 1:(9-1). Preferably, the first metal component is selected from one or more of molybdenum, tungsten, nickel, and cobalt; the total acidity of the first zeolite is 0.02-0.5 mmol / g, and the first zeolite is selected from one or more of octahedral zeolite, mordenite, L-type zeolite, Y-type zeolite, Ω-type zeolite, ZSM-4 zeolite, and Beta zeolite; the inert support is selected from alumina and / or silicon dioxide. Based on the dry weight of the second hydrocracking catalyst, the second hydrocracking catalyst contains 11-60% by weight of a second metal component, calculated as oxides, and the balance being a second support; the second support includes a second zeolite and a silica-alumina support. Preferably, the second metal component is selected from one or more of molybdenum, tungsten, and nickel; the second zeolite is selected from one or more of octahedral zeolite, mordenite, L-type zeolite, Y-type zeolite, Ω-type zeolite, ZSM-4 zeolite, ZSM-5 zeolite, and Beta zeolite.

[0017] Optionally, based on the dry weight of the catalyst, the first hydrorefining catalyst, the second hydrorefining catalyst, and the third hydrorefining catalyst each independently contain 11-60% by weight of a third metal component based on oxides, 0.5-8% by weight of a modifying component based on oxides, and the balance being a third support. Preferably, the third metal component is selected from one or more of nickel, molybdenum and tungsten, the modifying component is phosphorus, and the third carrier is selected from silicon oxide and / or aluminum oxide.

[0018] Optionally, the feed temperature of the raw oil is 320-430℃; The diesel fuel has a density of 0.83 g / mL or higher and a nitrogen content of 600 ppm or lower. Preferably, the density is 0.83-0.85 g / mL and the nitrogen content is 100-300 ppm. The light wax oil has a density of 0.83-0.89 g / mL and a nitrogen content of less than 600 ppm. Preferably, the density is 0.85-0.89 g / mL and the nitrogen content is 300-600 ppm.

[0019] The second aspect of this disclosure provides a catalytic hydrogenation system used in the method provided in the first aspect of this disclosure, the system comprising: a reactive distillation column and a condenser; The reactive distillation column includes a feedstock inlet, a hydrogen inlet, a gaseous product outlet, a liquid product outlet, and a reflux inlet; the reactive distillation column is provided with an upper rectifying section and a lower stripping section, the rectifying section and the stripping section are in fluid communication, the rectifying section is provided with a first catalyst bed, and the stripping section is provided with a second catalyst bed; the feedstock inlet is located between the rectifying section and the stripping section; The condenser includes a refrigerant inlet, a refrigerant outlet, a raw material inlet to be cooled, and a cooled raw material outlet. The gaseous product outlet of the reactive distillation column is in fluid communication with the raw material inlet to be cooled of the condenser, and the cooled raw material outlet of the condenser is in fluid communication with the reflux liquid inlet.

[0020] Optionally, the reactive distillation column further includes a heavy component outlet, which is located on the side wall of the lower part of the rectification section.

[0021] The method disclosed herein has at least the following advantages through the above technical solution: (1) The method disclosed herein couples the reaction and distillation, eliminating the need for an additional distillation column, thus saving on initial investment and reducing operating costs; (2) The method disclosed herein couples the reaction and distillation, which can prevent secondary cracking of naphtha fractions and avoid the generation of light hydrocarbons due to naphtha cracking. It improves cracking selectivity while reducing hydrogen consumption; it reduces the proportion of isoparaffins in light naphtha, and the product obtained is a higher quality ethylene cracking feedstock.

[0022] Other features and advantages of this disclosure will be described in detail in the following detailed description section. Attached Figure Description

[0023] The accompanying drawings are provided to further illustrate the present disclosure and form part of the specification. They are used together with the following detailed description to explain the present disclosure, but do not constitute a limitation thereof. In the drawings: Figure 1 This is a schematic flowchart of one specific embodiment of the system disclosed herein for producing chemical feedstocks from diesel or light wax oil.

[0024] Figure 2 This is a schematic diagram of a specific method for loading catalyst into a reactive distillation column disclosed herein.

[0025] Explanation of reference numerals in the attached figures Detailed Implementation

[0026] The specific embodiments of this disclosure will be described in detail below with reference to the accompanying drawings. It should be understood that the specific embodiments described herein are for illustration and explanation only and are not intended to limit this disclosure.

[0027] In this disclosure, unless otherwise stated, directional terms such as "up," "down," "left," and "right" generally refer to "up," "down," "left," and "right" when the system and apparatus of this disclosure are in normal working order.

[0028] The first aspect of this disclosure provides a method for producing chemical feedstocks from diesel oil and / or light wax oil. The method includes: introducing feedstock oil from a feedstock oil inlet into a reactive distillation column to react with hydrogen; the feedstock oil is diesel oil and / or light wax oil; wherein the reactive distillation column includes an upper rectifying section and a lower stripping section in fluid communication, the feedstock oil inlet being located between the rectifying section and the stripping section; the rectifying section is provided with a first hydrorefining catalyst and a first hydrocracking catalyst arranged sequentially from top to bottom; the stripping section is provided with a second hydrorefining catalyst, a second hydrocracking catalyst, and a third hydrorefining catalyst arranged sequentially from top to bottom.

[0029] The method disclosed herein couples the hydrogenation reaction and distillation within a reactive distillation column, eliminating the need for a separate distillation column, thus saving on initial investment and reducing operating costs. Simultaneously, the reactive distillation column includes a rectification section and a stripping section. Diesel or light wax oil reacts with specific catalysts located in these sections to achieve different hydrogenation effects and improve the adaptability of the feedstock. Furthermore, the reaction products achieve separation of light and heavy components during the rectification and stripping processes. This technical solution enables the efficient production of high-quality chemical feedstocks and effectively avoids secondary cracking of naphtha fractions, reducing the content of isoalkanes in naphtha to enrich alkanes and produce excellent feedstocks for olefin production through cracking.

[0030] In one specific embodiment of this disclosure, the feedstock oil is reacted with hydrogen in a countercurrent manner within the reactive distillation column. In this embodiment, by making the hydrogen and feedstock oil come into countercurrent contact, the inorganic ammonia content in the hydrocracking reaction atmosphere is reduced, which can effectively improve the cracking activity of the hydrocracking catalyst in the stripping section and broaden the adaptability range for the nitrogen mass fraction of the feedstock oil.

[0031] In one specific embodiment of this disclosure, the method further includes: extracting heavy components from the lower part of the rectification section and introducing the heavy components from the bottom of the stripping section into the reactive distillation apparatus. In this embodiment, by circulating the extracted heavy components into the stripping section of the reactive distillation apparatus, the conversion of heavy components can be enhanced, resulting in more high-quality chemical feedstocks.

[0032] In one specific embodiment of this disclosure, the initial boiling point of the heavy component is 150-180°C, and the final boiling point is 220-260°C; the circulation rate of the heavy component can vary within a wide range. In one embodiment, the circulation rate of the heavy component is 5-30% by weight of the diesel or light wax oil feed rate.

[0033] In one specific embodiment of this disclosure, the top pressure of the reactive distillation column is 4-8 MPa, the reflux ratio at the top is 0.3-3.0, and the hydrogen-to-oil volume ratio is 300-3000 Nm³. 3 / m 3 The bottom temperature of the column is 300-430℃; preferably, the top pressure is 4-8 MPa, the reflux ratio at the top is 0.5-2.0, and the hydrogen-to-oil volume ratio is 600-1800 Nm³. 3 / m 3 The temperature at the bottom of the tower is 350-410°C. According to this disclosure, the temperature of the hydrogen can vary over a wide range, for example, it can be 320-430°C, and preferably, the temperature of the hydrogen is 350-410°C.

[0034] In one specific embodiment of this disclosure, the volume hourly space velocity (VHSV) of the first hydrorefining catalyst in the rectification section is 1-20 h⁻¹. -1 Preferably 2-15h -1 The volume hourly space velocity of the first hydrocracking catalyst is 1-20 h⁻¹. -1 Preferably 2-10h -1 .

[0035] In one specific embodiment of this disclosure, the volume hourly space velocity (VHSV) of the second hydrorefining catalyst in the stripping section is 0.2-10.0 h⁻¹. -1 Preferably 1.0-4.0h -1 The volume hourly space velocity (VHSV) of the second hydrocracking catalyst is 0.2-10.0 h⁻¹. -1 Preferably 1.5-4.5h -1 The volume hourly space velocity (VHSV) of the third hydrorefining catalyst is 10-20.0 h⁻¹. -1 Preferably 10-15h -1 .

[0036] In one specific embodiment of this disclosure, the method further includes: drawing out the gas phase from the top of the reactive distillation column, cooling and separating the gas and liquid phases, returning a portion of the separated liquid phase to the top of the reactive distillation column, and returning the hydrogen obtained after gas phase deposition to the reactive distillation column as raw material hydrogen.

[0037] According to this disclosure, the first hydrocracking catalyst is characterized by high metal content and high molecular sieve content. In one specific embodiment, based on the dry weight of the first hydrocracking catalyst, the first hydrocracking catalyst contains 11-50% by weight of a first metal component, calculated as oxides, and the balance being a first support; the first support comprises a first zeolite and an inert support in a weight ratio of 1:(9-1); preferably, the first metal component is selected from one or more of molybdenum, tungsten, nickel, and cobalt; the total acidity of the first zeolite is 0.02-0.5 mmol / g, and the first zeolite is selected from one or more of octahedral zeolite, mordenite, L-type zeolite, Y-type zeolite, Ω-type zeolite, ZSM-4 zeolite, and Beta zeolite, preferably Y-type zeolite; the inert support is selected from alumina and / or silica; In a preferred embodiment, the first metal component contains molybdenum and / or tungsten, and nickel and / or cobalt, with the content of molybdenum and / or tungsten as oxides being 10-35 wt% and the content of nickel and / or cobalt as oxides being 1-15 wt% based on the dry weight of the first hydrocracking catalyst.

[0038] In a preferred embodiment, the first carrier comprises a first zeolite and alumina, wherein the alumina is composed of microporous alumina and macroporous alumina in a weight ratio of (3-1):1, wherein the microporous alumina is alumina in which the pore volume of pores with a diameter of less than 80 angstroms accounts for more than 95% of the total pore volume, and the macroporous alumina is alumina in which the pore volume of pores with a diameter of 60-600 angstroms accounts for more than 70% of the total pore volume.

[0039] According to this disclosure, the second hydrocracking catalyst is characterized by high metal content, high silica-alumina support, and low molecular sieve content. In one specific embodiment, based on the dry weight of the second hydrocracking catalyst, the second hydrocracking catalyst contains 11-60% by weight of a second metal component, calculated as oxides, and the balance being a second support; the second support includes a second zeolite and a silica-alumina support, and the weight ratio of the second zeolite to the silica-alumina support can be selected as needed, which will not be elaborated here; preferably, the second metal component is selected from one or more of molybdenum, tungsten, and nickel; the second zeolite is selected from one or more of octahedral zeolite, mordenite, L-type zeolite, Y-type zeolite, Ω-type zeolite, ZSM-4 zeolite, ZSM-5 zeolite, and Beta zeolite, preferably Y-type zeolite.

[0040] In a preferred embodiment, the second metal component contains molybdenum and / or tungsten, and nickel; based on the dry weight of the first hydrocracking catalyst, the content of molybdenum and / or tungsten, calculated as oxides, is 10-35 wt%, and the content of nickel, calculated as oxides, is 1-15 wt%.

[0041] In a preferred embodiment, the second carrier comprises a second zeolite, alumina, and amorphous aluminum silicate. Preferably, the weight ratio of alumina, amorphous aluminum silicate, and the second zeolite is 1:(0.01-6):(0.00-0.8). The alumina contains microporous alumina and macroporous alumina in a weight ratio of (1-3):1. More preferably, the microporous alumina is alumina with a pore volume of less than 70 angstroms accounting for more than 95% of the total pore volume, and the macroporous alumina is alumina with a pore volume of 70-600 angstroms accounting for more than 70% of the total pore volume.

[0042] According to this disclosure, the first, second, and third hydrorefining catalysts can be hydrorefining catalysts well known to those skilled in the art. In one specific embodiment, based on the dry weight of the catalyst, each of the first, second, and third hydrorefining catalysts independently contains 11-60% by weight of a third metal component (calculated as oxide), 0.5-8% by weight of a modifying component (calculated as oxide), and the balance being a third support; preferably, the third metal component is selected from one or more of nickel, molybdenum, and tungsten, the modifying component is phosphorus, and the third support is selected from silicon oxide and / or aluminum oxide.

[0043] In a preferred embodiment, the third metal component is selected from molybdenum and / or tungsten, and nickel, and the modifying component is phosphorus; based on the dry weight of the catalyst, the content of molybdenum and / or tungsten, calculated as oxides, is 10-50% by weight, the content of nickel, calculated as oxides, is 1-10% by weight, and the content of phosphorus, calculated as oxides, is 0.5-8% by weight.

[0044] In a preferred embodiment, the third carrier is selected from silicon oxide and alumina. Based on the total weight of the third carrier, the silicon oxide content is 2-45% by weight and the alumina content is 55-98% by weight; more preferably, the silicon oxide content is 5-40% by weight and the alumina content is 60-95% by weight. The specific surface area and pore volume of the silicon oxide and alumina can vary within a wide range. Preferably, the specific surface area of ​​each silicon oxide and alumina can be independently 150-350 m². 2 / g, preferably 180-300m 2 / g, the pore volume of silica and alumina can each be independently 0.4-1mLg, preferably 0.5-8mLg.

[0045] According to this disclosure, the feed temperature of the raw oil is 320-430℃; the density of the diesel oil is above 0.83 g / mL, and the nitrogen content is below 600 ppm, preferably 0.83-0.85 g / mL, and the nitrogen content is 100-300 ppm; the density of the light wax oil is 0.83-0.89 g / mL, and the nitrogen content is below 600 ppm, preferably 0.85-0.89 g / mL, and the nitrogen content is 300-600 ppm. In this disclosure, the density of the diesel oil and the density of the light wax oil are both measured at 20℃.

[0046] In one specific embodiment of this disclosure, the content of straight-chain alkanes in diesel oil is 35-45% by weight, and the content of straight-chain alkanes in light wax oil is 25-35% by weight.

[0047] like Figure 1 As shown, the second aspect of this disclosure provides a catalytic hydrogenation system used in the method provided in the first aspect of this disclosure. The system includes: a reactive distillation column 3 and a condenser 4; the reactive distillation column 3 includes a feedstock inlet, a hydrogen inlet, a gaseous product outlet, a liquid product outlet, and a reflux inlet; the reactive distillation column 3 is provided with an upper rectification section 10 and a lower stripping section 11, the rectification section and the stripping section being fluidly connected; a first catalyst bed is provided in the rectification section, and a second catalyst bed is provided in the stripping section; the feedstock inlet is located between the rectification section and the stripping section; the condenser 4 includes a refrigerant inlet, a refrigerant outlet, a feedstock inlet to be cooled, and a cooled feedstock outlet; the gaseous product outlet of the reactive distillation column is fluidly connected to the feedstock inlet to be cooled in the condenser, and the cooled feedstock outlet of the condenser is fluidly connected to the reflux inlet.

[0048] The system disclosed herein has a coupled rectification section and a stripping section in the reactive distillation column, eliminating the need for a subsequent main fractionation column. At the same time, it can make full use of the heat of reaction to achieve the separation of light and heavy components in the reactive distillation process, resulting in lower equipment investment and operating costs.

[0049] In one specific embodiment of this disclosure, the reactive distillation column further includes a heavy component outlet, which is disposed on the side wall of the lower part of the distillation section.

[0050] In one specific embodiment of this disclosure, the system further includes a feedstock oil heating device 2, the outlet of which is in fluid communication with the feedstock oil inlet of the reactive distillation column. As is well known to those skilled in the art, any heating device capable of preheating the feedstock oil to the required inlet temperature is acceptable; for example, it could be a heating furnace or a heat exchanger.

[0051] In one specific embodiment of this disclosure, the system further includes a cold high-pressure separator 5, a cold low-pressure separator 9, and a settling tank 6; the cold high-pressure separator 5 includes a gas phase inlet, a gas phase outlet, and a liquid phase outlet; the cold low-pressure separator 9 includes a liquid phase inlet, a gas phase outlet, and a liquid phase outlet; and the settling tank 6 includes a gas phase inlet, a gas phase outlet, and a liquid phase outlet. The gas phase inlet of the cold high-pressure separator 5 is fluidly connected to the outlet of the condenser 4, the gas phase outlet of the cold high-pressure separator is fluidly connected to the gas phase inlet of the settling tank 6, the liquid phase outlet of the cold high-pressure separator is fluidly connected to the liquid phase inlet of the cold low-pressure separator 9 and the reflux liquid inlet of the reactive distillation column, and the liquid phase inlet of the cold low-pressure separator 9 is also fluidly connected to the liquid phase outlet of the settling tank 6.

[0052] In one specific embodiment of this disclosure, the system further includes a circulating hydrogen compressor 7 and a hydrogen heating device 8. The circulating hydrogen compressor includes an inlet and an outlet, and the hydrogen heating device 8 includes a material inlet to be heated and a material outlet to be heated. The inlet of the circulating hydrogen compressor is in fluid communication with the gas phase outlet of the settling tank 6, the outlet of the circulating hydrogen compressor is in fluid communication with the material inlet to be heated of the hydrogen heating device 8, and the material outlet to be heated of the hydrogen heating device 8 is in fluid communication with the hydrogen inlet of the reactive distillation column.

[0053] The present disclosure will be further illustrated by the following examples, but the present disclosure is not limited thereto.

[0054] The characteristics of the feedstocks used in the following examples and comparative examples are shown in Table 1. The commercial brand name of the second hydrorefining catalyst in the stripping section used in the examples is RN-410B, the commercial brand name of the second hydrocracking catalyst in the stripping section is RHC-131 or RHC-133, the commercial brand name of the first hydrocracking catalyst in the rectifying section is RHC-220, and the commercial brand name of the first and second hydrorefining catalysts is RPT-10. All of these are produced by the Changling Catalyst Branch of Sinopec Corporation.

[0055] Example 1 In such Figure 1 The system shown uses the following method to produce chemical feedstock from feedstock A, and employs the catalyst combination shown in Table 2 (specific loading method as follows). Figure 2 (As shown in Table 3) The product is filled, and under the operating conditions in Table 3, the product distribution and key product properties are listed in Table 3.

[0056] The specific method is as follows: After being preheated by the feedstock oil heating device 2, the feedstock oil 1 enters the feedstock oil inlet of the reactive distillation tower 3. In the stripping section, it undergoes hydrodesulfurization, hydronitrogenation, olefin saturation, and partial aromatic saturation reactions with the upward-flowing hydrogen under the action of the hydropretreatment catalyst (i.e., the second hydrorefining catalyst 22). Simultaneously, the hydrorefined product oil contacts, exchanges heat with, and undergoes mass transfer with the upward-flowing cracked oil gas. The refined heavy components undergo further hydrocracking reactions under the action of the second hydrocracking catalyst 23 in the lower part of the stripping section 11. The light components rise while the heavy components further descend, and then react with the third hydrorefining catalyst 24 in the stripping section 11. Finally, it leaves from the bottom of the reactor as unconverted oil 19.

[0057] In the upper part of the stripping section, the vapor phase of the light components gradually rises to the rectification section 10 and reacts with the first hydrocracking catalyst 21 and the first hydrorefining catalyst 20 for separation. Some of the heavy components gradually descend back to the stripping section for further reaction, while another portion, with an initial boiling point of 170°C and a final boiling point of 230°C, 17 is extracted and recycled to the bottom of the reactive distillation column. After mixing with heated hydrogen gas 18, it enters the stripping section. The recycling rate of the heavy components is 6% by weight of the feed A. The light components in the hydrocracking product oil are further separated in the rectification section. The vapor phase at the top of the reactive distillation column is cooled by the condenser 4 and enters the cold high-pressure separator 5. Part of the liquid phase in the cold high-pressure separator returns to the top of the reactive distillation column as the top reflux 12, and the other part enters the cold low-pressure separator 9. The vapor phase from the cold high-pressure separator enters the settling tank 6 for further settling. A portion of the vapor phase 13 from the settling tank is discharged, while most of it passes through the circulating hydrogen compressor 7, is mixed with supplementary hydrogen, and then further heated by the hydrogen heating device 8. The heated hydrogen 18 enters the bottom of the reactive distillation column as reactive hydrogen. The liquid phase 15 from the settling tank enters the cold low-pressure separator 9. The gas at the top of the cold low-pressure separator 9 is discharged as cold low-pressure fraction 14, while the cold low-pressure fraction oil at the bottom exits the device as crude naphtha fraction 16.

[0058] Examples 2-4 Chemicals were produced from raw material A in the same system as in Example 1, using the same method as in Example 1, except that the operating conditions were different. The specific operating conditions are shown in Table 3, and the product distribution and key properties are listed in Table 3.

[0059] Examples 5-8 Chemicals were produced from raw material B in the same system as in Example 1, using the same method as in Example 1, except that the operating conditions were different. The specific operating conditions are shown in Table 4, and the product distribution and key properties are listed in Table 4.

[0060] Example 9 Chemical feedstock was produced from raw material B in the same system as in Example 5 and in the same method as in Example 1, except that the types of catalysts used were different. The first hydrogenation catalyst used was RHC-210 and the second hydrocracking catalyst was RHC-133.

[0061]

[0062]

[0063]

[0064] As shown in Examples 1-4, from the perspective of liquid hydrocarbon product distribution, the method disclosed herein can control the liquefied petroleum gas (LPG) content to below 1% under the condition that the tail oil ratio is approximately 74-84%, significantly reducing the proportion of secondary cracking; at the same time, the light naphtha yield is also below 4%. In terms of product properties, the light naphtha contains over 90% alkanes, with over 50% being n-alkanes, making it a high-quality feedstock for olefin production via cracking; the heavy naphtha has low sulfur and nitrogen content and an aromatic potential exceeding 65, making it a high-quality feedstock for aromatics production via reforming; and the tail oil has a BMCI value in the range of 14-18, also making it a high-quality feedstock for ethylene production via cracking.

[0065]

[0066] As can be seen from Examples 5-8 in Table 4, in terms of the distribution of liquid hydrocarbon products, the method disclosed herein can control the content of liquefied petroleum gas to be less than 4% under the condition that the proportion of tail oil is about 55-64%, which greatly reduces the proportion of secondary cracking; at the same time, the yield of light naphtha is also less than 9%.

[0067] In terms of product properties, the light naphtha contains over 90% alkanes, with n-alkanes accounting for approximately 50%, making it a high-quality feedstock for olefin production via cracking. The heavy naphtha has low sulfur and nitrogen content and an aromatic potential exceeding 60, making it a high-quality feedstock for aromatics production via reforming. The tail oil has a BMCI value in the range of 12-16, also making it a high-quality feedstock for ethylene production via cracking.

[0068] From the comparison of effects, with the tail oil yield being basically the same, the liquefied gas and light naphtha yields of Example 9 were significantly higher than those of Example 5, indicating that the preferred catalyst is beneficial to improving the selectivity of light hydrocarbons. In addition, in terms of the hydrocarbon composition of the products light naphtha and diesel, the light naphtha of Example 9 had relatively fewer n-hydrocarbons, while the diesel had a higher cetane number and a lower alkane content, and the product composition was worse than that of Example 5.

[0069] The preferred embodiments of this disclosure have been described in detail above with reference to the accompanying drawings. However, this disclosure is not limited to the specific details of the above embodiments. Within the scope of the technical concept of this disclosure, various simple modifications can be made to the technical solutions of this disclosure, and these simple modifications all fall within the protection scope of this disclosure.

[0070] It should also be noted that the various specific technical features described in the above specific embodiments can be combined in any suitable manner without contradiction. In order to avoid unnecessary repetition, this disclosure will not describe the various possible combinations separately.

[0071] Furthermore, various different embodiments of this disclosure can be combined in any way, as long as they do not violate the spirit of this disclosure, they should also be regarded as the content disclosed in this disclosure.

Claims

1. A method for producing chemical feedstock from diesel oil and / or light wax oil, the method comprising: The feedstock oil is introduced from the feedstock oil inlet into the reactive distillation tower to react with hydrogen; The feedstock oil is diesel oil and / or light wax oil; The reactive distillation column includes an upper rectification section and a lower stripping section that are fluidly connected, with the feed oil inlet located between the rectification section and the stripping section; the rectification section is provided with a first hydrorefining catalyst and a first hydrocracking catalyst from top to bottom; the stripping section is provided with a second hydrorefining catalyst, a second hydrocracking catalyst and a third hydrorefining catalyst from top to bottom.

2. The method according to claim 1, wherein, The feedstock oil reacts with hydrogen in a countercurrent manner within the reactive distillation tower.

3. The method according to claim 1, wherein, The method further includes: extracting heavy components from the lower part of the rectification section and introducing the heavy components from the bottom of the stripping section into the reactive distillation apparatus.

4. The method according to claim 3, wherein, The initial boiling point of the heavy component is 150-180℃, and the final boiling point is 220-260℃; the circulation rate of the heavy component is 5-30% by weight of the diesel or light wax oil feed rate.

5. The method according to claim 1, wherein, The reactive distillation column has a top pressure of 4-8 MPa, a top reflux ratio of 0.3-3.0, and a hydrogen-to-oil volume ratio of 300-3000 Nm³. 3 / m 3 The bottom temperature of the column is 300-430℃; preferably, the top pressure is 4-8 MPa, the reflux ratio at the top is 0.5-2.0, and the hydrogen-to-oil volume ratio is 600-1800 Nm³. 3 / m 3 The temperature at the bottom of the tower is 350-410℃; The volume hourly space velocity (VHSV) of the first hydrorefining catalyst in the rectification section is 1-20 h⁻¹. -1 Preferably 2-15h -1 The volume hourly space velocity of the first hydrocracking catalyst is 1-20 h⁻¹. -1 Preferably 2-10h -1 ; The volume hourly space velocity (VHSV) of the second hydrorefining catalyst in the stripping section is 0.2-10.0 h⁻¹. -1 Preferably 1.0-4.0h -1 The volume hourly space velocity (VHSV) of the second hydrocracking catalyst is 0.2-10.0 h⁻¹. -1 Preferably 1.5-4.5h -1 The volume hourly space velocity of the third hydrorefining catalyst is 10-20 h⁻¹. -1 Preferably 10-15h -1 .

6. The method according to claim 1, wherein, The method further includes: drawing out the gas phase from the top of the reactive distillation column, cooling and separating the gas and liquid phases, returning a portion of the separated liquid phase to the top of the reactive distillation column, and returning the hydrogen obtained after gas phase deposition to the reactive distillation column as raw material hydrogen.

7. The method according to claim 1, wherein, Based on the dry weight of the first hydrocracking catalyst, the first hydrocracking catalyst contains 11-50% by weight of a first metal component, calculated as oxides, and the balance being a first support; the first support comprises a first zeolite and an inert support in a weight ratio of 1:(9-1). Preferably, the first metal component is selected from one or more of molybdenum, tungsten, nickel, and cobalt; the total acidity of the first zeolite is 0.02-0.5 mmol / g, and the first zeolite is selected from one or more of octahedral zeolite, mordenite, L-type zeolite, Y-type zeolite, Ω-type zeolite, ZSM-4 zeolite, and Beta zeolite; the inert support is selected from alumina and / or silicon dioxide. Based on the dry weight of the second hydrocracking catalyst, the second hydrocracking catalyst contains 11-60% by weight of a second metal component, calculated as oxides, and the balance being a second support; the second support includes a second zeolite and a silica-alumina support. Preferably, the second metal component is selected from one or more of molybdenum, tungsten, and nickel; the second zeolite is selected from one or more of octahedral zeolite, mordenite, L-type zeolite, Y-type zeolite, Ω-type zeolite, ZSM-4 zeolite, ZSM-5 zeolite, and Beta zeolite.

8. The method according to claim 1, wherein, Based on the dry weight of the catalyst, the first hydrorefining catalyst, the second hydrorefining catalyst and the third hydrorefining catalyst each independently contain 11-60% by weight of a third metal component based on oxides, 0.5-8% by weight of a modifying component based on oxides, and the balance being a third support. Preferably, the third metal component is selected from one or more of nickel, molybdenum and tungsten, the modifying component is phosphorus, and the third carrier is selected from silicon oxide and / or aluminum oxide.

9. The method according to claim 1, wherein, The feed temperature of the raw oil is 320-430℃; The diesel fuel has a density of 0.83 g / mL or higher and a nitrogen content of 600 ppm or lower. Preferably, the density is 0.83-0.85 g / mL and the nitrogen content is 100-300 ppm. The light wax oil has a density of 0.83-0.89 g / mL and a nitrogen content of less than 600 ppm. Preferably, the density is 0.85-0.89 g / mL and the nitrogen content is 300-600 ppm.

10. A catalytic hydrogenation system used in the method of any one of claims 1-9, the system comprising: Reactive distillation column (3) and condenser (4); The reactive distillation column includes a feedstock inlet, a hydrogen inlet, a gaseous product outlet, a liquid product outlet, and a reflux inlet; the reactive distillation column is provided with an upper rectifying section and a lower stripping section, the rectifying section and the stripping section are in fluid communication, the rectifying section is provided with a first catalyst bed, and the stripping section is provided with a second catalyst bed; the feedstock inlet is located between the rectifying section and the stripping section; The condenser includes a refrigerant inlet, a refrigerant outlet, a raw material inlet to be cooled, and a cooled raw material outlet. The gaseous product outlet of the reactive distillation column is in fluid communication with the raw material inlet to be cooled of the condenser, and the cooled raw material outlet of the condenser is in fluid communication with the reflux liquid inlet.

11. The catalytic hydrogenation system according to claim 10, wherein, The reactive distillation column also includes a heavy component outlet, which is located on the side wall of the lower part of the distillation section.

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