A heat integration process for coal diesel hydrocracking system

Abstract

The application discloses a coal diesel oil hydrocracking system heat integration process and relates to the technical field of petroleum chemical industry.The system comprises a hydrocracking reactor, a H2S stripping tower and a main fractionating tower which are sequentially connected.The middle reflux of the main fractionating tower is used for heating cold low fraction oil to obtain secondary cold low fraction oil.The diesel oil at the bottom outlet of the main fractionating tower is used for heating the secondary cold low fraction oil to obtain tertiary cold low fraction oil.The reaction product of the hydrocracking after one-time heat exchange is used for heating the tertiary cold low fraction oil to obtain quaternary cold low fraction oil.The quaternary cold low fraction oil is used as the reboiling heat source of the H2S stripping tower.The application changes the cold low fraction oil heat exchange process, fully realizes the waste heat utilization of the middle reflux of the main fractionating tower, the hydrocracking product, the diesel oil at the bottom outlet of the main fractionating tower and the cold low fraction oil, reduces the cooling load and the heat load, simultaneously closes and stops the reboiling heating furnace of the original H2S stripping tower and reduces the energy consumption.

Description

Technical Field

[0001] This invention relates to the field of petrochemical technology, and in particular to a thermal integration process for a coal and diesel hydrocracking system. Background Technology

[0002] Coal and diesel fraction hydrocracking is a common petroleum processing technology, consisting of two main parts: reaction and fractionation. The reaction part mainly includes feedstock heating and hydrocracking, hot high-pressure separation (hereinafter referred to as "hot high-pressure fraction"), hot low-pressure separation (hereinafter referred to as "hot low-pressure fraction"), cold high-pressure separation (hereinafter referred to as "cold high-pressure fraction"), cold low-pressure separation (hereinafter referred to as "cold low-pressure fraction"), and circulating hydrogen compression and fresh hydrogen replenishment. The fractionation part mainly includes desulfurization of low-grade oil, absorption and desorption, and product separation.

[0003] Hydrocracking is an exothermic reaction, resulting in excess heat in the reaction section. However, the feed gas pressure is high (approximately 12-15 MPa), which is directly transferred to the fractionation section (pressure approximately 0.1-0.9 MPa). This presents challenges such as large pressure differentials and difficulties in heat exchange. Therefore, conventional designs often utilize air or water cooling to cool the waste heat from the high-temperature fractionation gas. However, this process leads to a high cooling load on the reaction section, prevents the upgrading and utilization of waste heat, and results in high energy consumption and high process costs in the fractionation section.

[0004] Therefore, in the petroleum refining process, the reaction section of the coal and diesel hydrocracking unit has excess heat and a large cooling load, while the fractionation section lacks heat and requires heating furnaces, resulting in unbalanced energy utilization. There is an urgent need to develop a process that can reduce the cooling load of the reaction section, realize the upgrading and utilization of waste heat, and reduce the energy consumption of the fractionation section. Summary of the Invention

[0005] The present invention aims to at least solve one of the aforementioned technical problems existing in the prior art. Therefore, the objective of the present invention is to provide a thermal integration process for a coal-fired diesel hydrocracking system. By changing the heat exchange process of the cold low-grade oil, the reflux, diesel fuel, and hydrocracking reaction products in the main fractionation tower are exchanged with the cold low-grade oil. The heat-exchanged cold low-grade oil is then used as the reboiling heat source for the H2S stripping tower, thereby shutting down the original H2S stripping tower reboiler, saving the effective heat load of the H2S stripping tower reboiler, and achieving energy saving and efficiency improvement.

[0006] To achieve the above objectives, the technical solution adopted by the present invention is as follows:

[0007] This invention provides a thermal integration process for a coal and diesel hydrocracking system. The system includes a hydrocracking reactor, a H2S stripping tower, and a main fractionation tower connected in sequence. In the main fractionation tower, reflux is used to heat cold low-grade oil to obtain secondary cold low-grade oil.

[0008] The diesel oil at the bottom outlet of the main fractionation tower is used to heat the cold low-grade oil twice, and then to obtain the cold low-grade oil three times.

[0009] The reaction products of the hydrocracking are used to heat the cold low-grade oil three times after one heat exchange, resulting in four times cold low-grade oil.

[0010] The cold low-grade oil is used four times as the reboiling heat source for the H2S stripping tower, and after heat exchange, cold low-grade oil is obtained five times.

[0011] In existing coal and diesel hydrocracking processes, the cold low-pressure oil after exiting the cold low-pressure separator (cold low-pressure separator) undergoes a single heat exchange with the hot high-pressure separator (hot high-pressure separator). This invention adjusts this process to involve heat exchange with the primary reflux of the main fractionation tower. On one hand, this effectively increases the temperature of the cold low-pressure oil, ensuring that heat can be fully transferred to it. On the other hand, it recovers the heat lost during cooling of the condenser at the top of the main fractionation tower, thus upgrading and utilizing the waste heat from the primary reflux of the main fractionation tower.

[0012] In existing coal-diesel hydrocracking processes, the heat exchange process for diesel oil at the bottom of the main fractionation tower involves a primary heat exchange between diesel and the bottom stream of the H2S stripping tower, followed by a secondary heat exchange between diesel and the bottom stream of the jet fuel (i.e., aviation kerosene) stripping tower. This invention utilizes the higher thermal energy of the diesel oil at the main fractionation tower outlet, and performs a second heat exchange on the cold low-grade oil, resulting in a third heat exchange for the cold low-grade oil. In other words, the process of this invention is adjusted so that after the diesel oil is distilled from the bottom of the self-fractionation tower, it enters the primary heat exchanger between diesel and the bottom stream of the H2S stripping tower, and the secondary heat exchanger between diesel and the cold low-grade oil.

[0013] In existing coal-fired diesel hydrocracking processes, the hot high-grade gas undergoes two heat exchange processes. The first heat exchange occurs between the hot high-grade gas and the cold low-grade oil. This invention advances the first heat exchange, changing the process to a heat exchange between the hydrocracking reaction products and the cold low-grade oil. The outlet reaction oil from the hydrocracking reactor enters the primary heat exchanger for reaction oil-circulating hydrogen mixing and the secondary heat exchanger for reaction oil-cold low-grade oil. This allows the hydrocracking reaction products to provide a significant amount of heat to the cold low-grade oil, thus transferring thermal energy from the reaction section to the fractionation section.

[0014] This invention eliminates the need for a reboiler at the bottom of the H2S stripping tower. Instead, it delivers the cold, low-grade oil to the H2S stripping tower four times as a heat source at the bottom. In other words, it uses the cold, low-grade oil, heated three times and at a high temperature, as the reboiler heat source for the H2S stripping tower, thus effectively saving the effective heat load of the reboiler. This process requires the addition of a small heat exchanger.

[0015] This invention modifies the heat exchange process of the cold low-grade oil by exchanging heat with the reflux, diesel, and hydrocracking reaction products in the first fractionation tower. The cold low-grade oil after heat exchange is then used as the reboiling heat source for the H2S stripping tower, thereby shutting down the original H2S stripping tower reboiler, saving the effective heat load of the H2S stripping tower reboiler, and achieving energy saving and efficiency improvement.

[0016] In this invention, "substance primary" refers to a substance that has not undergone heat exchange, "substance secondary" refers to a substance that has undergone heat exchange once, and so on. For example, "cold low-temperature oil primary" refers to cold low-temperature oil that has not undergone heat exchange, "cold low-temperature oil secondary" refers to cold low-temperature oil that has undergone heat exchange once, "cold low-temperature oil tertiary" refers to cold low-temperature oil that has undergone heat exchange twice, and so on.

[0017] In some embodiments of the present invention, the reflux extraction rate in the main fractionation tower is increased to 200-250 t / h.

[0018] By reducing the reflux flow rate at the top of the main fractionation column and increasing the reflux extraction rate in the first stage of the main fractionation column, the heat at the top of the main fractionation column is transferred to the first reflux point, so that it can be fully supplied to the cold low-grade oil for heat exchange.

[0019] In some embodiments of the present invention, the reflux flow rate at the top of the main fractionation column is 40~50 t / h.

[0020] In some embodiments of the present invention, a reboiler is connected to the bottom of the H2S stripping tower.

[0021] In some embodiments of the present invention, the pressure difference between the bottom pressure of the H2S stripping tower and the pressure difference between the cold low-temperature oil stream at the bottom of the H2S stripping tower is 1~3 MPaG.

[0022] In this invention, the pressure difference between the cold low-grade oil flow rate in the reboiler at the bottom of the H2S stripping tower and the pressure at the bottom of the H2S stripping tower is 1~3 MPaG. This means that the pressure difference between the flow rate entering the fractionation section from the reaction section and the pressure difference between the fractionation section and the reaction section is small, which further ensures that the heat exchanger has no leakage risk and avoids the problems of large pressure difference and difficult heat exchange.

[0023] In some embodiments of the present invention, the bottom pressure of the H2S stripping tower is 0.02 MPaG lower than the top pressure of the H2S stripping tower.

[0024] In some embodiments of the present invention, the system further includes a naphtha stabilization tower and a low-pressure steam generator connected to the main fractionation tower, wherein the cold low-grade oil after heat exchange with the bottom product of the H2S stripping tower is used as the heat source for the bottom product of the naphtha stabilization tower.

[0025] In the prior art, diesel products are used three times as the reboiling heat source at the bottom of the naphtha stabilizer. This invention adjusts this process by leading the cold low-grade oil, after exchanging heat with the bottom product of the H2S stripping tower (i.e., the feed liquid of the main fractionation tower), to the naphtha stabilizer in the fractionation section of the system as the bottom heat source of the naphtha stabilizer, where it exchanges heat with the bottom product of the naphtha stabilizer. The heat exchanger used in this process is the old heat exchanger at the bottom of the naphtha stabilizer.

[0026] In some embodiments of the present invention, the cold low-grade oil fraction 6 after heat exchange with the naphtha stabilizer bottom product is used to heat the deoxygenated water in the low-pressure steam generator.

[0027] The cold low-grade oil after heat exchange with the bottom product of the naphtha stabilizer is led to a low-pressure steam generator to heat the deoxygenated water at a low temperature to generate low-pressure steam. The low-pressure steam can be used by other units. The steam generator is an existing unit.

[0028] In some embodiments of the present invention, the cold low-grade oil is used as feed for the H2S stripping tower after exchanging heat four times with the bottom product of the naphtha stabilizer and the deoxygenated water of the low-pressure steam generator.

[0029] The cold low-grade oil, after heat exchange with the deoxygenated water at a low temperature, is led to the H2S stripping tower as feed for the H2S stripping tower. The feed location remains unchanged compared to the existing coal and diesel hydrocracking system process.

[0030] In some embodiments of the present invention, the bottom product of the H2S stripping tower is directly fed into the main fractionation tower after four heat exchanges with the cold low-grade oil.

[0031] In existing technology, the bottom product of the H2S stripping tower first exchanges heat with the bottom product of the main fractionation tower before entering the main fractionation tower. This invention adjusts this process by leading the cold low-separation oil four times to exchange heat with the bottom product of the H2S stripping tower, after which the bottom product directly enters the main fractionation tower. This increases the outlet temperature of the bottom product from the H2S stripping tower, reduces the reboiling load at the bottom of the subsequent main fractionation tower, and effectively utilizes the heat energy at high temperatures. This process uses the existing heat exchanger for the bottom product of the H2S stripping tower.

[0032] In some embodiments of the present invention, the raw materials in the system first pass through a raw material buffer tank during feeding, and no heat exchangers are installed before the raw material buffer tank for diesel and raw materials, or for aviation kerosene and raw materials; the diesel is the bottom product of the main fractionation tower; and the aviation kerosene is the bottom product of the aviation kerosene stripping tower.

[0033] In some embodiments of the present invention, the heat of the diesel oil is used as the heating source for the secondary heating of the cold low-separation oil; and the heat of the jet fuel is used as the reboiling heat source for the absorption tower.

[0034] This invention eliminates the diesel-raw material primary and aviation kerosene-raw material secondary heat exchangers before the raw materials enter the atmospheric pressure buffer tank in the prior art, so as to free up the heat of diesel and aviation kerosene to heat the cold low-grade oil and serve as the reboiling heat source for the absorption tower.

[0035] In some embodiments of the present invention, the system further includes a hot high-pressure separator connected to a hydrocracking reactor, wherein the reaction products of the hydrocracking are used to heat circulating hydrogen to obtain secondary circulating hydrogen.

[0036] In some embodiments of the present invention, the hot high-pressure separated gas obtained by the hot high-pressure separator is used as the heating source for the secondary circulating hydrogen mixing.

[0037] In some embodiments of the present invention, the hot high-pressure separated gas after heating and circulating hydrogen mixing is used as a secondary heating source for the high-pressure raw material.

[0038] In the existing technology, the process of hot high-pressure gas separation is to heat the hot high-pressure gas once to cool the low-pressure oil, heat the hot high-pressure gas a second time to circulate and mix with hydrogen, and then let the hot high-pressure gas enter the air cooler (air cooling) a third time. The present invention adjusts this process to heat the hot high-pressure gas once to circulate and mix with hydrogen, and heat the raw material (i.e., high-pressure raw material) a second time, thereby significantly reducing the air cooling load of the hot high-pressure gas.

[0039] In some embodiments of the present invention, the system further includes a jet fuel stripping tower, a naphtha stripping tower, and an absorption and desorption tower connected in sequence to the main fractionation tower, wherein the product of the jet fuel stripping tower serves as the heat source at the bottom of the absorption and desorption tower.

[0040] In the prior art, the heat source at the bottom of the absorption and desorption tower is the reflux in the first stage of the main fractionation tower. The present invention adjusts this to heat the product of the side-stream jet stripping tower (i.e., jet stripping tower) in one stage, while still using the old heat exchanger at the bottom of the absorption and desorption tower.

[0041] In some embodiments of the present invention, the secondary reflux of the main fractionation column serves as the heat source at the bottom of the naphtha fractionation column.

[0042] In the prior art, the heat source at the bottom of the naphtha fractionation tower is the diesel fuel at the bottom of the main fractionation tower, which undergoes three heat exchanges. This invention adjusts the heat source to a secondary reflux in the main fractionation tower, while still using the old heat exchanger at the bottom of the naphtha fractionation tower.

[0043] In some embodiments of the present invention, a water cooler is provided before the absorption and desorption tower to cool the heavy naphtha circulating absorbent; the water in the circulating water cooler is heated once and then used as a cold source to cool the secondary extracted stream of the absorption and desorption tower.

[0044] This invention adds a water cooler, which first cools the heavy naphtha circulating absorbent with cold water. After the heavy naphtha circulating absorbent is cooled by the cold water, i.e., heated once, it is used as a cold source to cool the secondary extracted stream of the absorption and desorption tower. The heat exchanger is an existing heat exchanger.

[0045] In some embodiments of the present invention, the water temperature in the water cooler before the absorption and desorption tower is 5~15°C.

[0046] In some embodiments of the present invention, the water temperature in the water cooler before the absorption and desorption tower is 6~8°C.

[0047] In some examples of the present invention, the water temperature in the water cooler before the absorption and desorption tower is 7°C.

[0048] In some embodiments of the present invention, the high-pressure feed oil pressurized by the reaction feed pump is shut off to exchange heat with the post-reaction feed oil, and the valve at the bypass point is opened so that the high-pressure feed oil no longer passes through the bypass.

[0049] In the original unit, the high-pressure feed oil pressurized by the reaction feed pump exchanges heat with the reacted feed oil. A bypass was opened, and there was a valve at the bypass. Now the original valve is closed so that the high-pressure feed oil no longer passes through the bypass.

[0050] This invention utilizes cold low-grade oil as a heat transfer medium, exchanging heat with the reflux in the first main fractionation tower, diesel fuel, and hydrocracking reaction products to raise the temperature of the cold low-grade oil. This high-temperature cold low-grade oil then supplies heat to the reboiler at the bottom of the desulfurization tower, the feed liquid in the main fractionation tower, the naphtha stabilizer, and the low-pressure steam generator, achieving rational utilization of the overall system's thermal energy. On the other hand, this invention directs the hot high-grade gas before it enters the air cooler to exchange heat with the pressurized high-pressure feedstock, thereby recovering a large amount of thermal energy from the hot high-grade gas. The essence of this invention is to optimize the heat exchange network of a traditional hydrocracking unit, improving the system's thermal efficiency. Compared to existing processes, the heat exchange process design of this invention eliminates the need for the reboiler at the bottom of the H2S stripping tower, saving fuel gas usage. Simultaneously, because a large amount of thermal energy from the hot high-grade gas and the reflux in the first main fractionation tower is recovered and utilized, the system's cooling load is reduced, the number of air coolers required is decreased, and the available space in the system is expanded.

[0051] In some embodiments of the present invention, the temperature of the fourth cold low-temperature oil separation process is ≥300°C.

[0052] In some embodiments of the present invention, the temperature of the cold low-separation oil entering the H2S stripping tower is 170~190°C.

[0053] In some embodiments of the present invention, the temperature of the cold low-separation oil entering the H2S stripping tower is 170~180°C.

[0054] In some embodiments of the present invention, after the hot high-pressure separated gas is reheated with the high-pressure raw material, the temperature of the hot high-pressure separated gas is 140~200℃.

[0055] In some embodiments of the present invention, after the hot high-pressure separated gas reheats the high-pressure raw material, the temperature of the raw material is 150~180°C.

[0056] In some embodiments of the present invention, after the reaction product undergoes a second and third heat exchange with the cold low-part oil, the temperature of the reaction product is 250~300°C.

[0057] In some embodiments of the present invention, the temperature of the second extracted stream from the absorption and desorption tower is 25~30°C after cooling.

[0058] In some embodiments of the present invention, the top pressure of the H2S stripping tower is 0.4~0.6 MPaG.

[0059] In some embodiments of the present invention, the top pressure of the H2S stripping tower is 0.5~0.6 MPaG.

[0060] In some specific embodiments of the present invention, the top pressure of the H2S stripping tower is 0.55~0.6 MPaG.

[0061] In some examples of the present invention, the top pressure of the H2S stripping tower is 0.56~0.58 MPaG.

[0062] Lowering the pressure of the H2S stripping tower increases the relative volatility within the tower, reduces the difficulty of separating the gas and liquid phases, further lowers the bottom product temperature, and reduces the bottom reboiling load.

[0063] In some embodiments of the present invention, the pressure at the top of the absorption and desorption tower is 0.4~0.6 MPaG, and the pressure of the absorption and desorption tower is lower than the pressure of the H2S stripping tower.

[0064] In some embodiments of the present invention, the pressure at the top of the absorption and desorption tower is 0.4~0.5 MPaG.

[0065] In some specific embodiments of the present invention, the pressure at the top of the absorption and desorption tower is 0.45~0.47 MPaG.

[0066] Lowering the pressure of the absorption and desorption tower in the fractionation section allows the absorbed moisture to enter the absorption and desorption tower from the top of the H2S stripping tower under its own pressure.

[0067] In some embodiments of the present invention, the top pressure of the main fractionation column is 0.028~0.078 MPaG.

[0068] In some embodiments of the present invention, the top pressure of the main fractionation column is 0.070~0.078 MPaG.

[0069] In some embodiments of the present invention, the pressure at the top of the naphtha stabilizer is 0.7~0.9 MPaG.

[0070] In some embodiments of the present invention, the pressure at the top of the naphtha stabilizer is 0.8~0.9 MPaG.

[0071] In some specific embodiments of the present invention, the pressure at the top of the naphtha stabilizer is 0.85~0.87 MPaG.

[0072] In some embodiments of the present invention, the top pressure of the naphtha fractionation tower is 0.1~0.2 MPaG.

[0073] In some embodiments of the present invention, the top pressure of the naphtha fractionation tower is 0.15~0.16 MPaG.

[0074] In some embodiments of the present invention, the liquid separation rate (i.e. flow rate) of the heavy naphtha circulating absorbent is increased to 30~50 t / h, while the tower pressure of the absorption and desorption tower is controlled to 0.4~0.6 MPa, thereby increasing the absorption effect of the absorption and desorption tower and compensating for the loss of absorption effect caused by the pressure reduction of the absorption and desorption tower.

[0075] In some embodiments of the present invention, the flow rate of the heavy naphtha circulating absorbent is 45~50 t / h.

[0076] In some embodiments of the present invention, the theoretical number of plates in the H2S stripping tower is 25 to 35.

[0077] In some embodiments of the present invention, the main fractionation column has 34 to 42 theoretical plates.

[0078] In some embodiments of the present invention, the naphtha fractionation tower has a theoretical number of 25 to 35 plates.

[0079] Compared with the prior art, the beneficial effects of the present invention are:

[0080] (1) By changing the heat exchange process of cold low fraction oil, this invention fully realizes the utilization of waste heat from the reflux in the first stage of the main fractionation tower, hydrocracking products, diesel oil at the bottom outlet of the main fractionation tower, and cold low fraction oil, thereby reducing the cooling load and heat load. At the same time, it shuts down the original H2S stripping tower reboiler, reducing energy consumption.

[0081] (2) The present invention utilizes the reaction products of hydrocracking and the waste heat of hot high-separation gas to upgrade and utilize the feedstock and cold low-separation oil, which greatly reduces the external energy supply of the system. The process provided by the present invention can shut down 1 heating furnace and 10 to 16 air coolers, including 4 to 6 air coolers for hot high-separation gas, 2 to 4 air coolers for the top of the main fractionation tower, and 4 to 6 air coolers for the top of the H2S stripping tower, which reduces the fuel consumption of the heating furnace by more than 35% and the corresponding cooling load of the unit by more than 16%.

[0082] (3) This invention only involves the adjustment of the heat exchange process and does not change the reaction system and the oil-gas separation system. Therefore, it is easy to implement, requires little investment, has no safety hazards, and is suitable for the construction of new equipment and the renovation of old equipment. Attached Figure Description

[0083] Figure 1 This is a process flow diagram of Comparative Example 1 of the present invention.

[0084] Figure 2 This is a process flow diagram of Embodiment 1 of the present invention.

[0085] The numbering in the attached figures is explained as follows:

[0086] 1-Low-pressure feed pump; 2-Primary feed to quaternary diesel heat exchanger; 3-Secondary feed to jet fuel heat exchanger; 4-High-pressure feed pump inlet buffer tank; 5-High-pressure feed pump; 6-Primary high-pressure feed to secondary reaction product heat exchanger; 7-Reaction feed heater; 8-Tertiary hydrogen mixing to primary reaction product heat exchanger; 9-Hydrocracking reactor; 10-Hydrorefining reactor; 11-Tertiary hydrogen mixing to tertiary reaction product heat exchanger; 12-Hot high-pressure separator; 13-Hydraulic turbine; 14-Hot low-pressure separator; 15-Primary cold low-pressure oil to primary hot high-pressure gas heat exchanger; 16-Primary hydrogen mixing to secondary hot high-pressure gas heat exchanger; 17-Hot high-pressure gas air cooler; 18-Hot low-pressure gas air cooler; 19-Hot low-pressure separator; 20-Hot high-pressure separator 21-Separator; 22-Circulating hydrogen compressor feed buffer tank; 23-Circulating hydrogen compressor-turbine unit; 24-Condensing turbine condenser-air cooler; 25-New hydrogen compressor; 26-H2S stripping tower; 27-H2S stripping tower overhead air cooler; 28-H2S stripping tower overhead water cooler; 29-H2S stripping tower bottom circulating pump; 30-H2S stripping tower bottom reboiler; 31-H2S stripping tower bottom oil-diesel primary heat exchanger; 32-Main fractionation tower; 33-Main fractionation tower bottom circulating pump; 34-Main fractionation tower bottom reboiler; 35-Main fractionation tower secondary reflux pump; 36-Naphtha stabilizer tower bottom reboiler; 37-Low-pressure steam generator. 38-Reflux pump in the main fractionation tower; 39-Reboiler at the bottom of the absorption and desorption tower; 40-Air cooler for the overhead gas of the main fractionation tower; 41-Reflux tank at the top of the main fractionation tower; 42-Jet kerosene stripping tower; 43-Secondary heat exchanger between naphtha stabilizer feed and jet kerosene; 44-Air cooler for jet kerosene product; 45-Circulating water cooler for jet kerosene product; 46-Diesel pump for the bottom product of the main fractionation tower; 47-Reboiler at the bottom of the jet kerosene side-stream stripping tower; 48-Reboiler at the bottom of the naphtha fractionation tower; 49-Air cooler for diesel product; 50-Heat exchanger between crude naphtha and heavy naphtha; 51-Naphtha fractionation tower; 52-Air cooler for the overhead gas of the naphtha fractionation tower; 53-Circulating water cooler for the overhead gas of the naphtha fractionation tower; 54-Reflux tank at the top of the naphtha fractionation tower; 5 5-Heavy naphtha pump; 56-Heavy naphtha air cooler; 57-Heavy naphtha circulating water cooler; 58-Absorption desorption tower; 59-Absorption desorption tower secondary reflux circulating water cooler; 60-Absorption desorption tower secondary reflux pump; 61-Absorption desorption tower primary reflux circulating cooler; 62-Absorption desorption tower primary reflux pump; 63-Absorption desorption tower bottom oil pump; 64-Naphtha stabilizer tower; 65-Naphtha stabilizer tower top gas air cooler; 66-Naphtha stabilizer tower top gas circulating water cooler; 67-Naphtha stabilizer tower top reflux tank; 68-High pressure feed primary to high heat fraction secondary heat exchanger; 69-H2S stripping tower bottom reboiler; 70-Deoxygenated water to diesel heat exchanger; 71-Low temperature cold water to heavy naphtha heat exchanger. Detailed Implementation

[0087] The present invention will be further described in detail below through specific embodiments. Unless otherwise specified, the raw materials, reagents, or apparatus used in the embodiments and comparative examples are all available from conventional commercial sources or can be obtained by existing technical methods. Unless otherwise specified, the experimental or testing methods are conventional methods in the art.

[0088] Figure 1 This is a process flow diagram of Comparative Example 1 of the present invention. In the diagram, the reaction section flows as follows: the mixed raw materials are pressurized by a low-pressure raw material pump 1, and sequentially pass through a primary-to-fourth-stage heat exchanger 2 (raw material primary to diesel quaternary) and a secondary-to-jet fuel heat exchanger 3 for heat exchange with the product diesel and jet fuel. Then, the mixture is pressurized again by a high-pressure raw material pump inlet buffer tank 4 and a high-pressure feed pump 5, and heated to a specified temperature by a secondary heat exchanger for the reaction products and a reaction feed heater 7. It is then mixed with mixed hydrogen and fed into a hydrorefining reactor 10 and a hydrocracking reactor 9. The reaction products and residual hydrogen flow out from the outlet of the hydrocracking reactor 9, sequentially passing through a tertiary-to-first-stage heat exchanger for mixed hydrogen and reaction products 8, a primary-to-second-stage heat exchanger for high-pressure raw materials and reaction products 6, and a secondary-to-tertiary-stage heat exchanger for mixed hydrogen and reaction products 11 for cooling before entering a hot high-pressure separator 12. Inside the hot high-pressure separator 12, the hot stream is separated into a vapor and liquid phases. The liquid hydrocarbons pass through the hydraulic turbine 13 and are then separated in the hot low-pressure separator 14. The liquid phase in the hot low-pressure separator 14 enters the H2S stripping tower 25 of the fractionation section. The vapor phase at the top of the hot low-pressure separator 14 is cooled by the hot low-pressure gas air cooler 18 and enters the cold low-pressure separator 19 together with the oil from the cold high-pressure separator 20.

[0089] The gas phase in the hot high-pressure separator 12 is cooled and enters the cold high-pressure separator 20 after passing through the cold low-pressure oil primary to hot high-pressure gas primary heat exchanger 15, the mixed hydrogen primary to hot high-pressure gas secondary heat exchanger 16, and the hot high-pressure gas air cooler 17. The liquid hydrocarbons in the cold high-pressure separator 20 are sent to the cold low-pressure separator 19 for separation. After heat exchange in the cold low-pressure separator 19 through the cold low-pressure oil primary to hot high-pressure gas primary heat exchanger 15, the liquid hydrocarbons enter the H2S stripping tower 25 of the fractionation section. That is, the cold low-pressure separator 19 is connected in sequence to the cold low-pressure oil primary to hot high-pressure gas primary heat exchanger 15 and the H2S stripping tower.

[0090] The gas phase in the cold high-pressure separator 20 enters the feed buffer tank 21 of the circulating hydrogen compressor for circulating hydrogen separation. The separated gas is pressurized by the circulating hydrogen compressor-turbine unit 22 and returned to the reaction section. The hydrogen produced after the circulating hydrogen compressor-turbine unit 22 is separated and returned to the inlet of the hot high-pressure gas air cooler 17 under the control of a valve. After the valve, the hydrogen is mixed with the supplementary hydrogen and then heat-exchanged through the hydrogen mixing primary to hot high-pressure gas secondary heat exchanger 16 and the hydrogen mixing secondary to reaction product tertiary heat exchanger 11. Then it is divided into two paths: one path is used as pre-furnace hydrogen mixing, which is mixed with the raw materials and enters the high-pressure raw material primary to reaction product secondary heat exchanger 6; the other path is used as post-furnace hydrogen mixing, which is heat-exchanged through the hydrogen mixing tertiary to reaction product primary heat exchanger 8 and mixed with the material at the outlet of the reaction feed heater 7 and enters the hydrorefining reactor 10.

[0091] The fractionation process is as follows: the product oil from the reaction section enters the H2S stripping tower 25. The C4 and above components in the gas phase at the top of the tower are cooled by the air cooler 26 and the circulating water cooler 27 at the top of the H2S stripping tower before entering the reflux tank 28 at the top of the H2S stripping tower. The sulfur-containing gas at the top of the tower enters the absorption and desorption tower 58. The liquid hydrocarbons in the reflux tank 28 at the top of the H2S stripping tower enter the tower as feed via the secondary reflux in the absorption and desorption tower 58; the other path enters the top of the H2S stripping tower 25. The bottom material is pressurized by the bottom circulation pump 29, heated in the reboiler 30 at the bottom of the H2S stripping tower, and then returned to the bottom of the H2S stripping tower 25. The other route is pressure-pressed out from the H2S stripping tower 25, and after heat exchange in the bottom oil-diesel primary heat exchanger 31 of the H2S stripping tower, it is divided into two routes as feed for the main fractionation tower 32.

[0092] The naphtha component at the top of the main fractionation column 32 is cooled by the main fractionation column top air cooler 40 and enters the main fractionation column top reflux tank 41. One path of the naphtha component in the reflux tank returns to the top of the main fractionation column; the other path passes through the crude naphtha-heavy naphtha heat exchanger 50, where it is mixed with the circulating naphtha from the naphtha stabilizer 64 and used as feed to the naphtha fractionation column 51.

[0093] The main fractionation tower has primary and secondary reflux lines and a jet fuel extraction line in its middle section. Another route is pumped by the primary reflux pump 38, pressurized, and returned to the main fractionation tower 32 after heat exchange via the reboiler 39 at the bottom of the absorption and desorption tower. The secondary reflux is extracted from 12 trays, pressurized by the secondary reflux pump 35, and returned to the main fractionation tower 32 after heat exchange via the reboiler 36 at the bottom of the naphtha stabilizer and the low-pressure steam generator 37.

[0094] The jet fuel at the bottom of the first tray of the jet fuel stripping tower 42, after heat exchange in the naphtha fractionation tower reboiler 48, returns to the bottom of the tower, providing heat to the bottom. The jet fuel product at the bottom of the tower undergoes heat exchange in the feed-jet fuel heat exchanger 3 and the naphtha stabilizer feed-jet fuel secondary heat exchanger 43, and is then cooled by the jet fuel product air cooler 44 and the jet fuel product circulating water cooler 45 before being sent out of the unit.

[0095] The diesel fuel at the bottom of the main fractionation tower 32 is divided into two routes: one route is pressurized by the main fractionation tower bottom circulation pump 33, enters the main fractionation tower bottom reboiler 34, and returns to the bottom of the tower. The other route is pressurized by the main fractionation tower bottom product diesel fuel pump 46, passes through the H2S stripping tower bottom oil to diesel primary heat exchanger 31, the jet fuel side stream stripping tower bottom reboiler 47, the naphtha fractionation tower bottom reboiler 48, and the feedstock primary to diesel quaternary heat exchanger 2. After heat exchange, it is cooled by the diesel product air cooler 49 and then sent out of the unit.

[0096] The light naphtha component at the top of naphtha fractionation column 51 is cooled by the overhead air cooler 52 and the overhead circulating water cooler 53 before entering the overhead reflux tank 54. The light naphtha component in the reflux tank is either returned to the top of the column or discharged from the unit. The bottom of naphtha fractionation column 51 is equipped with a reboiler 48. The heavy naphtha at the bottom is pressurized by the heavy naphtha pump 55, and after heat exchange in the crude naphtha-heavy naphtha heat exchanger 50, it is divided into two paths: one path is sent to the reforming workshop. The other path is cooled by the heavy naphtha air cooler 56 and the heavy naphtha circulating water cooler 57 before being split into another path that enters the top of the absorption and desorption column 58 as an absorbent.

[0097] The dry gas from the top of the absorption and desorption tower 58 is sent out for external desulfurization. Two reflux systems are located in the middle of the tower. The first reflux is drawn from tray 43, cooled by the second reflux circulating water cooler 59, and then pressurized by the second reflux pump 60 before returning to the tower. The second reflux is drawn from tray 35, cooled by the first reflux circulating cooler 61, and then pressurized by the first reflux pump 62 before mixing with light hydrocarbons from the top reflux tank of the hydrogen sulfide desulfurization tower and entering tray 35. A reboiler is located at the bottom of the absorption and desorption tower 58. The mixture at the bottom of the tower is pressurized by the bottom oil pump 63, passes through the naphtha stabilizer feed-jet kerosene secondary heat exchanger 43, and after heat exchange, is divided into three streams and enters the naphtha stabilizer tower 64.

[0098] The liquid hydrocarbons at the top of naphtha stabilizer 64 are cooled by the naphtha stabilizer top gas air cooler 65 and the naphtha stabilizer top gas circulating water cooler 66 before entering the naphtha stabilizer top reflux tank 67. The dry gas at the top of the naphtha stabilizer top reflux tank 67 is sent to the external desulfurization unit.

[0099] Figure 2The diagram below shows the process flow of Embodiment 1 of the present invention. As can be seen, Embodiment 1 of the present invention has made the following adjustments compared to Comparative Example 1:

[0100] 1. Change in gas phase heat exchange process in hot high pressure separator 12: The process changes so that the hot high pressure gas first passes through the hydrogen mixing primary ~ hot high pressure gas secondary heat exchanger 16 and the high pressure raw material primary ~ hot high pressure gas secondary heat exchanger 68, and then enters the hot high pressure gas air cooler 17 for cooling, thereby recovering a large amount of waste heat in the hot high pressure gas.

[0101] 2. Changes to the cold low-pressure oil heat exchange process: A new sixth heat exchange for cold low-pressure oil is added. 1) The oil just exiting the cold low-pressure separator enters the heat exchanger through the secondary-jet fuel heat exchanger 3, where it exchanges heat with the reflux in the main fractionation tower 1; 2) The cold low-pressure oil enters twice through the primary-diesel quadruple heat exchanger 2, where it exchanges heat with the diesel for the second time; 3) The position of the primary-hot high-pressure gas heat exchanger 15 is moved forward, allowing the cold low-pressure oil to enter three times, where it exchanges heat with the reaction products for the second time. Now, the position of the primary-hot high-pressure gas heat exchanger 15 is after the secondary-reaction product heat exchanger 8; 4) The reboiler 30 at the bottom of the H2S stripping tower is removed, and the newly added reboiler 69 at the bottom of the H2S stripping tower is used as a heat exchanger. The reboiler at the bottom of H2S stripping tower 25 allows the cold low-grade oil to pass through the reboiler four times as the reboiler heat source for the bottom of H2S stripping tower 25; 5) The cold low-grade oil enters five times using the original H2S stripping tower bottom oil-diesel primary heat exchanger 31, exchanging heat with the feed liquid of the main fractionation tower 32 (i.e., the bottom product of the H2S stripping tower); 6) The cold low-grade oil passes six times through the original naphtha stabilizer bottom reboiler 36 as the bottom reboiler heat source; 7) The cold low-grade oil exchanges heat with deoxygenated water seven times using the original low-pressure steam generator 37 to generate low-pressure steam; the cold low-pressure separator 19 is connected to the feed secondary-aviation kerosene heat exchanger 3, the feed primary-diesel quadruple heat exchanger 2, the cold low-grade oil primary-hot high-grade gas primary heat exchanger 15, and the H2S stripping tower... The stripping tower bottom reboiler 69, the H2S stripping tower bottom oil to diesel primary heat exchanger 31, the naphtha stabilizer tower bottom reboiler 36, the low-pressure steam generator 37, and the H2S stripping tower are connected in sequence.

[0102] 3. After the diesel product is produced at the bottom of the self-fractionation tower 32, it no longer enters the H2S stripping tower bottom oil to diesel primary heat exchanger 31, but directly enters the reboiler 47 at the bottom of the jet fuel side stripping tower to obtain secondary diesel.

[0103] 4. After the diesel product is heated by the reboiler 48 at the bottom of the naphtha fractionation tower and the primary to secondary diesel heat exchanger 2, it flows into the newly added deoxygenated water to diesel heat exchanger 70 to exchange heat with low temperature water.

[0104] 5. The secondary reflux of the main fractionation column 32 is heated by the reboiler at the bottom of the naphtha fractionation column 51.

[0105] 6. After exiting the jet fuel stripping tower 42, the jet fuel product is heated by the reboiler 39 at the bottom of the absorption and desorption tower.

[0106] 7. A new cooling process is added at point 58 of the absorption and desorption tower: a new low-temperature cold water-heavy naphtha heat exchanger 71 is installed to cool the heavy naphtha absorbent with cold water; then this cold water is led to the original absorption and desorption tower 2's return circulating water cooler 59 to exchange heat with the heat extracted from the absorption and desorption tower 2, thereby reducing the return temperature of the heat extracted from the tower 2. The purpose of this process is to enhance the absorption effect of the absorption and desorption tower.

[0107] Specific comparative examples and embodiments are provided below.

[0108] Comparative Example 1

[0109] This comparative example provides a process for the hydrocracking of coal and diesel fuel, using a 3.6 million tons / year coal and diesel fuel hydrocracking unit. The mixed feedstock contains 2170 mg / g of sulfur, has a flow rate of 402.7 t / h, and its distillation range is shown in Table 1.

[0110] Table 1 Feedstock Distillation Range of a Coal-Diesel Hydrocracking Unit

[0111]

[0112] Table 2 shows the main operating parameters of the tower system.

[0113] Table 2 Main operating parameters of the tower system

[0114]

[0115] The theoretical number of plates for each tower is as follows: H2S stripping tower 30 plates, main fractionation tower 38 plates, jet fuel stripping tower 10 plates, naphtha fractionation tower 30 plates, absorption and desorption tower 52 plates, and naphtha stabilization tower 46 plates.

[0116] Comparative Example 1 corresponds to a total effective heat load of 2444.54 × 10⁻⁶ for the heating furnace. 4 kcal / h, of which the H2S stripping tower reboiler is 1232.9 × 10 kcal / h. 4 kcal / h, main fractionation tower reboiler 1036.24×10 4 kcal / h, reaction feed furnace 175.4×10 4 kcal / h. Total effective cooling load: 9007.17 × 10 4 kcal / h, of which the hot high-temperature air cooler has a capacity of 3707.61 × 10 kcal / h. 4 kcal / h.

[0117] Table 3 shows the distillation range of the liquid products.

[0118] Table 3. Distillation range of liquid products in Comparative Example 1

[0119]

[0120] Table 4 shows the composition of the gas products.

[0121] Table 4 Comparative Gas Product Composition

[0122]

[0123] Example 1

[0124] This embodiment provides a thermal integration process for a coal and diesel hydrocracking system, which, compared with Comparative Example 1, has the following adjustments:

[0125] 1. Reduce the reflux flow rate at the top of the main fractionation column to 45 t / h, and use the reflux from the first stage of the main fractionation column (178.6℃, 250 t / h) to heat the cold low-grade oil (52.0℃, 175.8 t / h), raising its temperature to 116℃, absorbing 604.36 × 10⁻⁶ heat. 4 kcal / h;

[0126] 2. The cold low-part oil (117℃, 175.8t / h) was heated twice with diesel oil (255.2℃, 181.9t / h) to raise its temperature to 196.9℃, absorbing 914.73×10⁻⁶ heat. 4 kcal / h;

[0127] 3. The cold low-part oil was heated three times (196.9℃, 175.8t / h) with the reaction product (358.3℃, 517t / h) to raise its temperature to 307℃, absorbing 1830.57×10⁻⁶ heat. 4 kcal / h;

[0128] 4. After three heating cycles, the cold low-grade oil (307℃, 175.8t / h) flows sequentially through the reboiler 69 at the bottom of the H2S stripping tower, the primary heat exchanger 31 between the bottom oil and diesel fuel in the H2S stripping tower, the reboiler 36 at the bottom of the naphtha stabilizer tower, and the low-pressure steam generator 37, releasing 749 × 10⁻⁶ heat each time. 4 kcal / h, 640.04×10 4 kcal / h, 446.89×10 4 kcal / h and 258.62×10 4 kcal / h, cooled to 175℃ and fed into the 11th plate of the H2S stripping tower. At this time, the pressure of the cold low-grade oil stream is 2.8MPaG, and the pressure difference between the cold low-grade oil stream and the bottom pressure of the H2S stripping tower is about 2.3MPaG.

[0129] 5. The high-pressure feedstock (219.8℃, 287.7t / h) is heated from 48℃ to 168.4℃ by the high-pressure feedstock primary heat exchanger (68), which is then used to heat the high-pressure feedstock primary heat exchanger (381.5t / h) via the high-pressure feedstock primary heat exchanger to the high-pressure feedstock secondary heat exchanger. This process absorbs 2342.52 × 10⁻⁶ heat. 4 kcal / h; the hot high-pressure gas after heat exchange enters the hot high-pressure gas air cooler 17, reducing the cooling load by 866.11 × 10 compared to the comparative example. 4 kcal / h;

[0130] 6. The reflux (200℃, 270t / h) in the second main fractionation column is used as the heat source for the reboiler 48 at the bottom of the naphtha fractionation column, releasing 476.96×10⁻⁶ heat. 4 kcal / h;

[0131] 7. Using product jet fuel (199℃, 81.6t / h) as the heat source for the reboiler 39 at the bottom of the absorption and desorption tower, 167.07×10⁻⁶ heat is released. 4 kcal / h;

[0132] 8. Reduce the top pressure of H2S stripping tower 25 to 0.56 MPaG and the top pressure of absorption stripping tower 58 to 0.45 MPaG, corresponding to an increase in the flow rate of heavy naphtha circulating absorbent to 48 t / h.

[0133] 9. Low-temperature chilled water (7℃, 90t / h) generated by the factory's absorption-type lithium bromide hot water chiller unit is used to cool the heavy naphtha circulating absorbent (35.6℃, 48t / h) and the second absorption-desorption tower (38.8℃, 56.7t / h), respectively, reducing their temperatures to 20℃ and 26.5℃, corresponding to heat exchange loads of 34.8×10⁻⁶. 4 kcal / h and 34.21×10 4 kcal / h.

[0134] Table 5 shows the main operating parameters of the tower system in Example 1.

[0135] Table 5 Main operating parameters of the tower system in Example 1

[0136]

[0137] Example 1 corresponds to a total effective heat load of 1586 × 10⁻⁶ for the heating furnace. 4 kcal / h, of which the reboiler of the main fractionation column is 1096 × 10 4 kcal / h, reaction feed furnace 490×10 4 kcal / h, H2S stripping tower reboiler shut down; total effective cooling load 7488.89×10 4 kcal / h, of which the hot high-temperature air cooler has a capacity of 2841.5 × 10 kcal / h.4 kcal / h.

[0138] Table 6 shows the distillation range of liquid products.

[0139] Table 6 Distillation range of liquid product in Example 1

[0140]

[0141] Table 7 shows the composition of the gas products.

[0142] Table 7 Composition of the gas product in Example 1

[0143]

[0144] Comparing Example 1 and Comparative Example 1, it can be seen that:

[0145] 1. The material balance and product quality of Example 1 were unchanged compared to Comparative Example 1;

[0146] 2. The effective heat load of the heating furnace in Example 1 was reduced by 858.54 × 10⁻⁶ compared to the comparative example. 4 kcal / h, a decrease of 35.1%; the H2S stripping tower reboiler was shut down, thus saving 1232.9 × 10 kcal / h of effective heat load compared to Comparative Example 1. 4 kcal / h;

[0147] 3. In Example 1, the total effective cooling load of air cooling and water cooling was reduced by 1518.28 × 10⁻⁶ compared to the comparative example. 4 kcal / h, a decrease of 16.9%, and the hotter high-separation air cooling load decreased from 3707.16×10 4 kcal / h decreased to 2841.5×10 4 kcal / h, a reduction of 866.11 × 10 4 kcal / h, a decrease of 23.4%;

[0148] 4. Based on the same unit operating for 8400 hours per year, standard fuel oil at 2805 yuan / ton, and refrigeration utility unit price at 0.2 yuan / 10 tons. 4 Based on kcal and a heating furnace thermal efficiency of 90%, Example 1 shows an efficiency increase of RMB 27.445 million per year compared to Comparative Example 1.

[0149] Example 2

[0150] This embodiment provides a thermal integration process for a coal and diesel hydrocracking system. In embodiment 2, based on embodiment 1, the pressure of the H2S stripping tower is increased from 0.56 MPaG to 0.58 MPaG, the pressure of the absorption stripping tower is increased from 0.45 MPaG to 0.47 MPaG, the flow rate of the heavy naphtha circulating absorbent is reduced from 48 t / h to 46 t / h, and the flow rate of the first main fractionation tower is increased from 250 t / h to 260 t / h, while other operating parameters remain unchanged.

[0151] Table 8 shows the main operating parameters of the tower system in Example 2.

[0152] Table 8 Main operating parameters of the tower system in Example 2

[0153]

[0154] The corresponding total effective heat load of the heating furnace is 1572×10 4 kcal / h, of which the H2S stripping tower reboiler is 0×10 4 kcal / h, main fractionation column reboiler 1088×10 4 kcal / h, reaction feed furnace 484×10 4 kcal / h. Total cooling load: 7359.02 × 10 4 kcal / h, of which the hot high-temperature air cooler has a capacity of 2841.2 × 10 kcal / h. 4 kcal / h.

[0155] Table 9 shows the distillation range for liquid products.

[0156] Table 9. Distillation range of liquid product in Example 2

[0157]

[0158] Table 10 shows the composition of the gas products.

[0159] Table 10 Composition of the gas product in Example 2

[0160]

[0161] It should be noted that %v represents volume percentage.

[0162] Comparing Example 2 and Comparative Example 1, it can be seen that:

[0163] 1. The material balance and product quality of Example 2 are unchanged compared to the comparative example;

[0164] 2. The heating furnace heat load in Example 2 was reduced by 872.54 × 10⁻⁶ compared to the comparative example. 4kcal / h, a decrease of 35.7%; the H2S stripping tower reboiler was shut down, thus saving 1232.9 × 10 kcal / h of heat load compared to Comparative Example 1. 4 kcal / h;

[0165] 3. In Example 2, the total cooling load of air cooling and water cooling was reduced by 1648.15 × 10⁻⁶ compared to the comparative example. 4 kcal / h, a decrease of 18.3%, and the hotter high-separation air cooling load decreased from 3707.16×10 4 kcal / h decreased to 2841.2×10 4 kcal / h, a reduction of 866.41 × 10 4 kcal / h, a decrease of 23.4%;

[0166] 4. Based on an annual operating time of 8400 hours for the unit, an electricity price of 0.55 yuan / kWh, a standard fuel oil price of 2805 yuan / ton, and a refrigeration utility unit price of 0.2 yuan / 10 4 kcal, Example 1 shows an efficiency improvement of 27.962 million yuan / year.

[0167] By comparing the above comparative examples and embodiments, it can be seen that:

[0168] (1) In this embodiment of the invention, the primary heat exchanger for diesel fuel and the secondary heat exchanger for aviation kerosene before the diesel fuel enters the atmospheric pressure buffer tank are eliminated, and the heat of diesel fuel and aviation kerosene is used to heat the cold low-temperature oil and to serve as the reboiling heat source for the absorption tower, respectively.

[0169] (2) In the comparative example, the process of hot high-temperature gas is to heat the hot high-temperature gas once to cool the low-temperature oil, heat the hot high-temperature gas twice to circulate and mix with hydrogen, and heat the hot high-temperature gas three times to enter the air cooler. In this embodiment of the invention, it is adjusted to heat the hot high-temperature gas once to circulate and mix with hydrogen, and heat the hot high-temperature gas twice to heat the raw material (i.e. high-pressure raw material) once, thereby fully reducing the air cooler load of the hot high-temperature gas.

[0170] (3) In the comparative example, the process of the cold low-pressure oil after exiting the cold low-pressure separator is to exchange heat with the hot high-pressure gas once. In this embodiment of the invention, it is adjusted to exchange heat with the first reflux of the main fractionation tower. There are two advantages to this adjustment: first, it increases the temperature of the cold low-pressure oil to ensure that the heat can be fully transferred to the cold low-pressure oil; second, it recovers the heat lost by the cooling of the condenser at the top of the main fractionation tower. At the same time, compared with the comparative example, this embodiment of the invention reduces the reflux flow rate at the top of the main fractionation tower and increases the reflux extraction rate at the first reflux of the main fractionation tower, so that the heat at the top of the main fractionation tower is transferred to the first reflux point, so that it can be fully supplied to the cold low-pressure oil for heat exchange.

[0171] (4) In the comparative example, the heat exchange process of diesel at the bottom of the main fractionation tower is diesel to H2S stripping tower bottom stream once, and diesel to jet fuel stripping tower bottom reboil twice. In this embodiment of the invention, the higher heat energy of the diesel at the outlet is used to exchange heat with the cold low-grade oil after reflux heat exchange in the main fractionation tower for a second time. Therefore, the process of this embodiment of the invention is adjusted to diesel distilled out of the bottom of the self-fractionation tower and then enters the diesel to H2S stripping tower bottom stream once and diesel to cold low-grade oil twice in sequence.

[0172] (5) The hot high-separation gas in the comparative example has two heat exchange processes. The first heat exchange is between the hot high-separation gas and the cold low-separation oil. In this embodiment of the invention, the first heat exchange is brought forward and the heat exchange process is changed to the heat exchange between the reaction oil at the outlet of the hydrocracking reactor and the cold low-separation oil. That is, the reaction oil at the outlet of the hydrocracking reactor enters the reaction oil-circulating hydrogen mixing primary heat exchanger and the reaction oil-cold low-separation oil secondary heat exchanger in sequence, so as to provide a large amount of heat to the cold low-separation oil, thereby realizing the transfer of heat energy from the reaction part to the fractionation part.

[0173] (6) In this embodiment of the invention, the heating furnace at the bottom of the H2S stripping tower is suspended. The cold low-temperature oil that has been heated by three heat exchanges is transported to the H2S stripping tower as the bottom heat source. That is, the cold low-temperature oil that has been heated by three heat exchanges is used as the bottom reboiling heat source of the H2S stripping tower. The heating furnace is no longer used. In this process, a small heat exchanger is added as the bottom reboiler of the H2S stripping tower, but a heating furnace can be saved.

[0174] (7) In the comparative example, the bottom product of the H2S stripping tower first exchanges heat with the bottom product of the main fractionation tower and then enters the main fractionation tower. In this embodiment of the invention, the cold low fraction oil after heat exchange at the bottom of the H2S stripping tower is led to exchange heat with the bottom product of the H2S stripping tower. Then the bottom product of the H2S stripping tower directly enters the main fractionation tower. This method increases the outlet temperature of the bottom product of the H2S stripping tower, reduces the reboiling load at the bottom of the subsequent main fractionation tower, and effectively utilizes the heat energy at the high temperature. This process uses the old heat exchanger of the bottom product of the H2S stripping tower.

[0175] (8) In the comparative example, the diesel product was used three times as the reboiling heat source at the bottom of the naphtha stabilizer. In this embodiment of the invention, the cold low fraction oil after heat exchange with the product at the bottom of the H2S stripping tower was led to the naphtha stabilizer in the fractionation section of the system and used as the heat source at the bottom of the naphtha stabilizer. The heat exchanger used in this process is the old heat exchanger at the bottom of the naphtha stabilizer.

[0176] (9) In this embodiment of the invention, the cold low-temperature oil after heat exchange with the bottom of the naphtha stabilizer in (8) is led to a low-pressure steam generator to heat the deoxygenated water at a low temperature to generate low-pressure steam. The low-pressure steam can be used by other equipment in the plant. At the same time, the final temperature of the cold low-temperature oil is controlled at about 175°C. Then, the cold low-temperature oil after heat exchange is led to the H2S stripping tower as the feed for the H2S stripping tower. The feed position remains unchanged. The steam generator is an existing device.

[0177] (10) Compared with the comparative example, the embodiments of the present invention reduce the pressure of the H2S stripping tower, thereby increasing the relative volatility in the H2S stripping tower, reducing the difficulty of separating the gas and liquid phases in the H2S stripping tower, reducing the product temperature at the bottom of the tower, and reducing the reboiling load at the bottom of the tower; reduce the pressure of the absorption stripping tower, so that the absorbed moisture can enter the absorption stripping tower from the top of the H2S stripping tower by self-pressure; increase the flow rate of the heavy naphtha circulating absorbent (the absorbent flow rate is increased to 48t / h), and adjust it in conjunction with the reduction of the absorption stripping tower pressure, so as to further increase the absorption effect of the absorption stripping tower and make up for the loss of absorption effect caused by the pressure reduction of the absorption stripping tower;

[0178] (11) In the comparative example, the heat source at the bottom of the absorption and desorption tower is the primary reflux of the main distillation tower. In this embodiment of the invention, it is adjusted to the primary heat source of the side-stream jet fuel stripping tower product, and the old heat exchanger at the bottom of the absorption and desorption tower is still used. In the comparative example, the heat source at the bottom of the naphtha distillation tower is the diesel tertiary heat exchanger. In this embodiment of the invention, it is adjusted to the secondary reflux of the main distillation tower, and the old heat exchanger at the bottom of the naphtha distillation tower is still used.

[0179] (12) In this embodiment of the invention, a water cooler is added. Cold water is used to cool the heavy naphtha circulating absorbent first. After the cold water is heated once, it is used as a cold source to cool the second extracted stream of the absorption desorption tower. The heat exchanger is an existing heat exchanger.

[0180] (13) In the comparative example, the high-pressure feed pump 5 pressurizes the high-pressure raw material oil and exchanges heat with the reacted raw material oil. A bypass is opened, and there is a valve at the bypass. In this embodiment of the invention, the original valve is closed so that the high-pressure raw material oil no longer passes through the bypass.

[0181] In summary, this invention, by modifying the heat exchange process of the cold low-grade oil, fully utilizes the waste heat from the reflux in the main fractionation tower, hydrocracking products, diesel oil from the bottom outlet of the main fractionation tower, and the cold low-grade oil, reducing cooling and heat loads. Simultaneously, it shuts down the original H2S stripping tower reboiler, further reducing energy consumption. Furthermore, by upgrading the utilization of the waste heat from the hydrocracking reaction products and hot high-grade gas to heat the feedstock and cold low-grade oil, the external energy supply to the system is significantly reduced, resulting in a reduction of over 35% in heater fuel consumption and a corresponding reduction of over 16% in the unit's cooling load.

[0182] The above embodiments are preferred embodiments of the present invention, but the embodiments of the present invention are not limited to the above embodiments. Any changes, modifications, substitutions, combinations, or simplifications made without departing from the spirit and principle of the present invention shall be considered equivalent substitutions and shall be included within the protection scope of the present invention.

Claims

1. A thermal integration process for a coal-fired diesel hydrocracking system, the system comprising a hydrocracking reactor, a H2S stripping tower, and a main fractionation tower connected in sequence, characterized in that, The reflux in the first main fractionation tower is used to heat the cold low-grade oil, resulting in a secondary cold low-grade oil. The diesel oil at the bottom outlet of the main fractionation tower is used to heat the cold low-grade oil twice, and then to obtain the cold low-grade oil three times. The reaction products of the hydrocracking are used to heat the cold low-grade oil three times after one heat exchange, resulting in four times cold low-grade oil. The cold low-grade oil is used four times as the reboiling heat source for the H2S stripping tower, and after heat exchange, cold low-grade oil is obtained five times.

2. The process according to claim 1, characterized in that, The bottom of the H2S stripping tower is connected to a reboiler. And / or, the pressure difference between the bottom pressure of the H2S stripping tower and the pressure difference between the cold low-temperature oil stream at the bottom of the H2S stripping tower is 1~3 MPaG.

3. The process according to claim 1, wherein the system further comprises a naphtha stabilization tower and a low-pressure steam generator connected to the main fractionation tower, characterized in that, The cold low-grade oil fraction 5 after heat exchange with the bottom product of the H2S stripping tower is used as the heat source for the bottom product of the naphtha stabilizer tower. And / or, the bottom product of the H2S stripping tower is directly fed into the main fractionation tower after four heat exchanges with the cold low-grade oil.

4. The process according to claim 3, characterized in that, The cold low-grade oil, after being heated six times by the naphtha stabilizer bottom product, is used seven times to heat the deoxygenated water in the low-pressure steam generator.

5. The process according to claim 3, characterized in that, The cold, low-grade oil is then exchanged four times with the bottom product of the naphtha stabilizer and the deoxygenated water from the low-pressure steam generator before being used as feed for the H2S stripping tower.

6. The process according to claim 1, wherein the system further comprises a hot high-pressure separator connected to the hydrocracking reactor, characterized in that, The reaction products of the hydrocracking are used to heat and circulate hydrogen to obtain secondary circulate hydrogen; And / or, the hot high-pressure separated gas obtained by the hot high-pressure separator is used as the heating source for the secondary cycle of hydrogen mixing. And / or, the hot high-pressure separated gas after heating and circulating hydrogen mixing is used as a secondary heating source for the high-pressure raw material.

7. The process according to claim 1, wherein the system further comprises a jet fuel stripping tower, a naphtha fractionation tower, and an absorption and desorption tower connected in sequence to the main fractionation tower, characterized in that, The jet fuel produced by the jet fuel stripping tower is used as the heat source at the bottom of the absorption and desorption tower. And / or, the secondary reflux of the main fractionation column serves as the heat source at the bottom of the naphtha fractionation column; And / or, a water cooler is installed before the absorption and desorption tower to cool the heavy naphtha circulating absorbent; the water in the water cooler is heated once and then used as a cold source to cool the second extracted stream of the absorption and desorption tower.

8. The process according to claim 1 or 2, characterized in that, The temperature of the four cold low-temperature oil separation processes is ≥300℃; And / or, the temperature of the cold low-separation oil entering the H2S stripping tower is 170~190℃.

9. The process according to claim 6, characterized in that, After the high-pressure raw material is reheated, the temperature of the high-pressure separated gas is 140~200℃. And / or, after the hot high-pressure separation gas reheats the high-pressure raw material, the temperature of the raw material is 150~180℃.

10. The process according to claim 1, characterized in that, After the reaction product undergoes a second and third heat exchange with the cold low-grade oil, the temperature of the reaction product is 250~300℃.

11. The process according to claim 7, characterized in that, The water temperature in the water cooler is 5~15℃; And / or, the temperature of the second extracted stream from the absorption and desorption tower is 25~30℃ after cooling.

12. The process according to claim 7, characterized in that, The top pressure of the H2S stripping tower is 0.4~0.6 MPaG; And / or, the top pressure of the absorption and desorption tower is 0.4~0.6 MPaG, and the top pressure of the absorption and desorption tower is lower than the top pressure of the H2S stripping tower.

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