Wax oil hydrocracking unit absorption stabilization system and separation method

Abstract

The invention relates to an absorption stabilization system and a separation method of a wax oil hydrocracking unit, which are characterized in that an original naphtha separation tower is changed into a heavy naphtha tower, a feed stream is adjusted into a mixture of top oil of a steam stripping tower, top oil of a pre-fractionating tower and top oil of a main fractionating tower, and heavy naphtha is produced at the bottom of the tower; tower top distillate and tower bottom oil of the deethanizer are mixed and then enter a debutanizer, liquefied gas is obtained at the tower top after separation, light naphtha is obtained at the tower bottom, and part of the light naphtha serves as an absorbent to be recycled. By changing the top oil feeding position of the main fractionating tower, reducing the processing amount and the energy consumption of the debutanizer, and adopting pure light naphtha as an absorbent in the absorption tower, the adsorption effect is better, and the composition of the bottom of the debutanizer is changed, so that a reboiling heat source is changed from middle-section reflux of the main fractionating tower to product diesel oil, and the heat consumption level is reduced. The method does not change the reaction and main fractionating tower system process, has strong applicability in the aspects of new device construction and old device transformation, effectively solves the problem of high energy consumption in the traditional process, and improves the overall benefits of the device.

Description

Technical Field

[0001] The present invention relates to the technical field of petrochemical engineering, and particularly to an absorption and stabilization system and a separation method for a wax oil hydrocracking unit. Background Art

[0002] Wax oil hydrocracking is a core heavy fraction lightening process in petroleum processing. It mainly consists of five parts: hydrocracking reaction, separation of reaction product oil and gas, reaction product stripping column, main fractionating column, and absorption and stabilization. Among them, the task of the absorption and stabilization unit is to process the gas and light fraction oil from the stripping column and the main fractionating column system. Through a series of processes such as absorption, desorption, and stabilization, it is separated into dry gas, liquefied gas, light naphtha, and heavy naphtha products. Currently, the typical operation mode of this process is as follows: The gas from the top of the stripping column and the gas from the top of the main fractionating column enter the deethanizer together. Driven by the reboiling heat at the bottom of the column, gas is separated at the top of the column, and deethanized oil is fractionated at the bottom. These gases enter the absorber by self-pressure. Under the action of the ≥C5 fraction oil from the bottom of the debutanizer, ≥C3 components are removed to obtain dry gas products. The bottom oil of the absorber is pumped back to the deethanizer for recycling after being saturated with light hydrocarbons. The deethanized oil is mixed with the top oil of the main fractionating column and then enters the debutanizer. After separation, C3-C4 liquefied gas is produced at the top of the column, and debutanized oil is obtained at the bottom. Part of the debutanized oil is sent to the absorber as an absorbent, and the other part enters the light and heavy naphtha separation column. Finally, light naphtha products are obtained from the top of the column, and heavy naphtha products are obtained from the bottom of the column.

[0003] However, there are many problems to be solved urgently in this traditional process. On the one hand, the bottom oil of the debutanizer, which is the circulating naphtha and serves as the absorbent, has a large flow rate, resulting in a large processing volume of the system and thus high energy consumption. On the other hand, due to the high pressure at the top of the deethanizer (about 1000 KPag), the bottom temperature of the column also rises (about 120 °C). This makes it impossible to use the hot water generated by recovering waste heat as a heat source, and only low-pressure steam of about 0.5 MPa can be relied on, greatly increasing the production cost. In addition, the reboiling load at the bottom of the debutanizer is extremely large. Taking a wax oil hydrocracking unit with a scale of 260×10 4 tons / year as an example, its reboiling load is about 1500×10 4 kcal / h, which is equivalent to the heat of about 25 - 28 t / h of medium-pressure steam, further exacerbating energy consumption. Therefore, how to develop a new absorption and stabilization system for a wax oil hydrocracking unit that can effectively reduce separation energy consumption is an urgent problem to be solved currently. Summary of the Invention

[0004] In view of the technical problems existing in the prior art, the first object of the present invention is to provide an absorption and stabilization system for a wax oil hydrocracking unit, which changes the original naphtha separation tower into a heavy naphtha tower and adjusts the feed materials. The heavy naphtha tower separates heavy naphtha in advance to avoid subsequent repeated rectification, improves the separation efficiency, optimizes the reboiling heat sources of the deethanizer and the debutanizer, and changes the path of the top oil of the main fractionator, thereby reducing the energy consumption.

[0005] The second object of the present invention is to adopt a separation method for the absorption and stabilization system of the wax oil hydrocracking unit to solve the problems of large circulating naphtha flow rate in the absorption and stabilization system of the current wax oil hydrocracking unit, high bottom temperature of the deethanizer, and large reboiling load of the debutanizer.

[0006] To achieve the above objects, the present invention adopts the following technical solutions:

[0007] An absorption and stabilization system for a wax oil hydrocracking unit, comprising:

[0008] A heavy naphtha tower, the feed end of the heavy naphtha tower receives a liquid-phase mixture obtained by mixing, heating and flashing the stripping tower top oil, pre-fractionation tower top oil and main fractionation tower top oil from the upstream wax oil hydrocracking unit. Adjust the rectification operating conditions of the heavy naphtha tower, and a mixture of liquefied gas and light naphtha is obtained at the top of the tower, and a heavy naphtha product is obtained at the bottom of the tower.

[0009] A deethanizer, the feed end of the deethanizer receives a mixture of the gas phase after flashing and the stripping tower top gas. Adjust the rectification operating conditions of the deethanizer, and dry gas is obtained at the top of the tower and sent to the absorption tower, and a light naphtha mixture of saturated light hydrocarbons is obtained at the bottom of the tower. The liquid phase part at the bottom of the deethanizer passes through a reboiler, and after vaporizing 15-25%, it returns to the heavy naphtha tower, and the remaining liquid phase is sent to the debutanizer as a mixture of liquefied gas and light stone through a pump.

[0010] An absorption tower, which is used to remove the heavy components in the dry gas. Adjust the rectification operating conditions of the absorption tower, and dry gas products are obtained at the top of the tower and sent out of the device for desulfurization, and a light naphtha mixture of saturated light hydrocarbons is obtained at the bottom of the tower and sent back to the deethanizer for recycling. Among them, the absorption tower adopts total reflux operation. The gas at the top of the absorption tower is cooled to 45 °C by a circulating water cooler and enters a separation tank, and all the liquid is returned to the top of the absorption tower as reflux.

[0011] A debutanizer, the feed end of the debutanizer receives the mixture of liquefied gas and light naphtha from the top of the heavy naphtha tower and the bottom of the deethanizer. Adjust the rectification operating conditions of the debutanizer, and liquefied gas products are obtained at the top of the tower, and light naphtha products are obtained at the bottom of the tower.

[0012] Further, the top gas of the heavy naphtha tower is completely cooled to a liquid at 45°C by an air cooler and then enters a liquid separation tank. Part of the liquid is returned as reflux to the top of the heavy naphtha tower, and the remaining liquid is sent as a mixture of liquefied gas and light naphtha to the debutanizer through a pump.

[0013] Further, a part of the liquid phase at the bottom of the heavy naphtha tower passes through a reboiler, is vaporized by 15 - 25% and then returned to the heavy naphtha tower. The remaining liquid phase is sent out of the unit as a heavy naphtha product by self - pressure or through a pump. Part of it is cooled to 40°C by an air cooler and a circulating water cooler and then sent out of the unit, and the other part is sent directly to the continuous reforming unit as raw material after recovering heat.

[0014] Further, the flash evaporation is selected from a flash tank. The gas phase from the flash tank is mixed with the top gas of the stripper and then sent to the deethanizer for separation. Dry gas is obtained from the top and sent to an absorption tower to remove heavy components, and a light - stone mixture of saturated light hydrocarbons is obtained from the bottom.

[0015] Further, the deethanizer operates under total reflux. The top gas of the deethanizer is cooled to 45°C by a circulating water cooler and then enters a liquid separation tank. All the liquid is returned as reflux to the top of the deethanizer and is heated to 81°C by the gas - phase product of the flash tank and light naphtha.

[0016] Further, the mixtures of liquefied gas and light naphtha from the top of the heavy naphtha tower and the bottom of the deethanizer are combined, heated to a higher temperature by the jet fuel product, and then sent to the debutanizer by self - pressure or through a pump for separation. Liquefied gas is obtained from the top and light naphtha is obtained from the bottom.

[0017] Further, the top gas of the debutanizer is completely cooled to a liquid at 45°C by an air cooler and then enters a liquid separation tank; part of the liquid is returned as reflux to the top of the debutanizer, and the remaining liquid is sent as product liquefied gas out of the unit by self - pressure or through a pump for external desulfurization.

[0018] Further, a part of the liquid phase at the bottom of the debutanizer passes through a reboiler, is vaporized by 15 - 25% and then returned to the debutanizer. The remaining liquid phase, as the product light naphtha, is cooled by heat exchange, and then cooled to 40°C by self - pressure or through a pump by an air cooler and a circulating water cooler and is divided into two parts. One part is sent to the absorption tower as an absorbent, and the other part is sent out of the unit as a gasoline blending component.

[0019] Further, the liquid mixture is directly obtained by successively mixing, heating, and flash - evaporating the top oil of the stripper, the top oil of the pre - fractionator, and the top oil of the main fractionator.

[0020] Further, the liquid mixture is obtained by first mixing and heating - flash - evaporating the top oil of the stripper and the top oil of the pre - fractionator, and then mixing the liquid phase after flash - evaporation with the top oil of the main fractionator and heating it for the second time to raise the temperature.

[0021] Furthermore, the number of theoretical plates of the heavy naphtha column is 30, the feed end is located in the middle of the heavy naphtha column, and the feed is from the 13th tray. The operating pressure is 120 - 300 kPag, the reflux ratio is 1.3 - 1.8 (reflux ratio = reflux flow rate / distillate flow rate), the reboiler heat source at the bottom of the column is the middle-section stream of the main fractionation column, and the temperature of the middle-section reflux of the main fractionation column is 275°C. Since the feed temperature decreases and the fractionation difficulty increases, the middle-section reflux of the main fractionation column is used instead of the product diesel to provide heat for the heavy naphtha column.

[0022] Furthermore, the number of theoretical plates of the deethanizer column is 25, the feed end is located at the top of the deethanizer column, and the feed is specifically from the top liquid separation tank. The operating pressure is 900 - 1000 kPag, and the reboiler heat source at the bottom of the column is hot water, which is 130°C circulating hot water. Since the heavy naphtha is separated in advance, the system becomes lighter and the corresponding temperature level also drops. Therefore, there is no need to provide heat with additional steam, and all 130°C circulating hot water can be used as the reboiler heat source.

[0023] Furthermore, the number of theoretical plates of the absorber column is 10. The dry gas is fed from the bottom of the column, and the light naphtha is fed from the top of the column as the absorbent. The operating pressure is 900 - 1000 kPag.

[0024] Furthermore, the number of theoretical plates of the debutanizer column is 30, the feed end is located in the middle of the debutanizer column, and the feed is from the 14th tray. The operating pressure is 800 - 1200 kPag, and the reflux ratio is 1.5 - 3.0. The reboiler heat source at the bottom of the column is the diesel product obtained from the main fractionation column, and the temperature is 240°C. Since the heavy naphtha is separated in advance, the system becomes lighter and the corresponding temperature level also drops, and the product diesel is used instead of the middle-section reflux of the main fractionation column to provide the reboiler heat.

[0025] Furthermore, the arithmetic mean temperature difference of the reboiler heat exchanger at the bottom of the heavy naphtha column is ≥90°C. The flow rate on the hot stream side is 350 - 480 t / h, the temperature after heat exchange is 220 - 250°C, the flow rate of the cold-side stream is 650 - 780 t / h, and the temperature returning to the column is 155 - 170°C; specifically, the flow rate on the hot stream side is 415 t / h, the temperature after heat exchange is 235.7°C; the flow rate of the cold-side stream is 710 t / h, and the temperature returning to the column is 162.3°C.

[0026] Furthermore, the arithmetic mean temperature difference of the reboiler heat exchanger at the bottom of the deethanizer column is ≥20°C. The flow rate on the hot stream side is 200 - 260 t / h, the temperature after heat exchange is 110 - 130°C, the flow rate of the cold-side stream is 65 - 85 t / h, and the temperature returning to the column is 100 - 110°C; specifically, the flow rate on the hot stream side is 232 t / h, the temperature after heat exchange is 119°C; the flow rate of the cold-side stream is 75 t / h, and the temperature returning to the column is 103.5°C.

[0027] Further, the arithmetic mean temperature difference of the reboiler heat exchanger at the bottom of the debutanizer is ≥ 40°C, the flow rate on the hot stream side is 60 - 90 t / h, the temperature after heat exchange is 130 - 145°C, the flow rate of the cold side stream is 350 - 450 t / h, and the temperature returning to the column is 140 - 150°C; specifically, the flow rate on the hot stream side is 76 t / h, the temperature after heat exchange is 136°C; the flow rate of the cold side stream is 405 t / h, and the temperature returning to the column is 143.8°C. Among them, the calculation formula for the arithmetic mean heat transfer temperature difference is ((hot stream inlet temperature - cold stream outlet temperature) + (hot stream outlet temperature - cold stream inlet temperature)) / 2.

[0028] A separation method for an absorption and stabilization system of a wax oil hydrocracking unit, comprising the following steps: mixing, heating up, and flash distilling the top oil of the stripping column, the top oil of the pre-fractionating column, and the top oil of the main fractionating column to obtain a liquid-phase mixture and entering a heavy naphtha column for fractionation; mixing the gas phase generated by the flash tank with the top gas of the stripping column and then entering a deethanizer for separation; the dry gas at the top of the deethanizer enters an absorber to remove heavy components; the liquefied gas and light naphtha mixture at the top of the heavy naphtha column and the bottom of the deethanizer enter a debutanizer for separation.

[0029] Further, the top oil of the stripping column, the top oil of the pre-fractionating column, and the top oil of the main fractionating column are mixed, heated up, and flash distilled in sequence; or, first mix the top oil of the stripping column and the top oil of the pre-fractionating column, heat up and flash distill, and then mix the liquid phase after flash distillation with the top oil of the main fractionating column and heat up and raise the temperature for the second time.

[0030] The present invention has the following advantages:

[0031] 1. By separating the heavy naphtha component in advance through the heavy naphtha column, the present invention avoids its repeated rectification in subsequent columns, optimizes the separation process, and improves the separation efficiency. The original naphtha separation column is changed to a heavy naphtha column, and the feed composition is adjusted to the liquid-phase mixture of the top oil of the main fractionating column, the top oil of the stripping column, and the top oil of the pre-fractionating column, enabling each component to be separated in a more suitable column, reducing the separation steps and energy consumption, and enhancing the overall separation effect.

[0032] 2. For the deethanizer and debutanizer of the present invention, since the bottom product components become lighter and the temperature decreases, the reboiling heat source of the deethanizer is changed from steam to circulating hot water at 130°C, and the reboiling heat source of the debutanizer is changed from the intermediate section stream of the main fractionating column to diesel products, reducing the heat consumption level. The top oil of the main fractionating column no longer enters the debutanizer first, but directly enters the heavy naphtha column, reducing the processing volume of the debutanizer, thereby reducing energy consumption. Compared with Comparative Example 1, the total energy consumption cost in Example 1 is reduced by 12.502 million yuan / year, with a reduction rate of 24.1%; the total energy consumption cost in Example 2 is reduced by 11.451 million yuan / year, with a reduction rate of 22.1%.

[0033] 3. The absorbent of the absorption tower of the present invention changes from the heavy and light naphtha mixture at the bottom of the debutanizer to pure light naphtha. The fraction becomes lighter, the molecular weight is closer to the liquefied gas components, and the absorption effect is better. It can more effectively remove the heavy components in the dry gas, improve the quality of the dry gas product, and the dry gas, liquefied gas, light naphtha and heavy naphtha products all meet the quality requirements.

[0034] 4. Multiple devices of the present invention can utilize existing equipment, such as air coolers, circulating water coolers, liquid separation tanks, some heat exchangers, etc., reducing the investment in newly built equipment. Due to the reduction of the bottom temperature of the deethanizer, the steam reboiler can be shut down, and the hot water heat exchanger can utilize existing equipment; the reboiler heat exchanger of the debutanizer can also utilize existing equipment, further saving equipment costs. Brief Description of the Drawings

[0035] Figure 1 is a process flow diagram of Comparative Example 1 of the absorption and stabilization system of the existing wax oil hydrocracking unit.

[0036] Among them, 1 is the circulating water cooler for the deethanizer top gas, 2 is the gas-liquid separation tank for the deethanizer top gas, 3 is the product reflux pump for the deethanizer top, 4 is the heat exchanger between the deethanizer top reflux and the debutanizer bottom distillate, 5 is the heat exchanger between the deethanizer top reflux and the jet fuel, 6 is the deethanizer, 7 is the hot water reboiler at the bottom of the deethanizer, 8 is the steam reboiler at the bottom of the deethanizer, 9 is the product pump at the bottom of the deethanizer, 10 is the circulating water cooler for the absorber top gas, 11 is the gas-liquid separation tank for the absorber top gas, 12 is the absorber top reflux pump, 13 is the absorber, 14 is the absorber bottom pump, 15 is the heat exchanger between the deethanizer bottom distillate and the debutanizer bottom distillate, 16 is the heat exchanger between the deethanizer bottom distillate and the jet fuel, 17 is the air cooler for the debutanizer top gas, 18 is the gas-liquid separation tank for the debutanizer top gas, 19 is the product reflux pump for the debutanizer top, 20 is the debutanizer, 21 is the reboiler at the bottom of the debutanizer, 22 is the product pump at the bottom of the debutanizer, 23 is the air cooler for the naphtha separation tower top gas, 24 is the gas-liquid separation tank for the naphtha separation tower top gas, 25 is the product reflux pump for the naphtha separation tower top, 26 is the naphtha separation tower, 27 is the reboiler at the bottom of the naphtha separation tower, 28 is the product pump at the bottom of the naphtha separation tower.

[0037] Figure 2 is a process flow diagram of Embodiment 1 of the absorption and stabilization system of the wax oil hydrocracking unit of the present invention.

[0038] Among them, 1 is the de-ethanizer top gas circulating water cooler, 2 is the de-ethanizer top gas-liquid separation tank, 3 is the de-ethanizer top product reflux pump, 4 is the de-ethanizer top reflux / flashed gas stream heat exchanger, 5 is the de-ethanizer top reflux / light naphtha heat exchanger, 6 is the de-ethanizer column, 7 is the de-ethanizer bottom hot water reboiler, 8 is the de-ethanizer bottom steam reboiler, 9 is the de-ethanizer bottom product pump, 10 is the absorber top gas circulating water cooler, 11 is the absorber top gas-liquid separation tank, 12 is the absorber top reflux pump, 13 is the absorber column, 14 is the absorber bottom pump, 15 is the de-ethanizer bottom distillate / product jet fuel heat exchanger, 16 is the debutanizer top gas air cooler, 17 is the debutanizer top gas-liquid separation tank, 18 is the debutanizer top product reflux pump, 19 is the debutanizer column, 20 is the debutanizer bottom reboiler, 21 is the debutanizer bottom product pump, 22 is the absorbent circulating water cooler, 23 is the mixed oil feed pump, 24 is the hot water / feed heat exchanger, 25 is the heavy naphtha / feed heat exchanger, 26 is the flash tank, 27 is the heavy naphtha top gas air cooler, 28 is the heavy naphtha top gas-liquid separation tank, 29 is the heavy naphtha top product reflux pump, 30 is the heavy naphtha column, 31 is the heavy naphtha bottom reboiler, 32 is the heavy naphtha bottom product pump.

[0039] Figure 3 is the process flow diagram of Embodiment 2 of the absorption and stabilization system of the wax oil hydrocracking unit of the present invention. Among them, Figure 3 the markings of Figure 2 are the same as Detailed Description of the Invention

[0040] The present invention will be further described in detail below with reference to the accompanying drawings and specific embodiments.

[0041] Unless otherwise specified, the connection modes of the equipment, devices or systems in the following examples and comparative examples are all pipeline connections, which belong to the conventional technologies in the art. Arranging metering and control components such as valves, flow meters, thermometers, and pressure gauges on the pipelines is also a conventional technology in the art.

[0042] Taking the absorption and stabilization system of a 2.6 million tons / year wax oil hydrocracking unit as an example, the main operating conditions of Comparative Example 1, Example 1 and Example 2 are described under the same separation requirements.

[0043] Comparative Example 1

[0044] As Figure 1As shown in the figure, the system mainly consists of a deethanizer 6, an absorber 13, a debutanizer 20 and a heavy naphtha column 30. Its raw materials include the stripper top gas (3998 kg / h), the stripper top oil (77000 kg / h), the pre-fractionator top oil (25300 kg / h) and the main fractionator top oil (95440 kg / h). The target products are dry gas (content of components ≥ C3 ≤ 1 mol%), liquefied petroleum gas (content of C3 / C4 components ≥ 97 mol%), light naphtha (Engler distillation range 24°C - 76°C) and heavy naphtha (Engler distillation range 83°C - 166°C).

[0045] The processing flow of the raw materials is as follows: After the stripper top gas, the stripper top oil, the pre-fractionator top oil are mixed with the bottom oil of the absorber, they enter the system through the pipeline at the top of the deethanizer. The mixed material first flows through the deethanizer top gas circulating water cooler 1 for cooling, and then enters the deethanizer top gas-liquid separation tank 2 for equilibrium separation. The separated gas phase is sent to the absorber 13; the liquid phase is boosted by the deethanizer top product reflux pump 3, and successively undergoes heat exchange with the bottom oil of the debutanizer 20 through the heat exchanger 4, and then undergoes secondary heat exchange with the jet fuel product through the heat exchanger 5. After being heated to about 81°C, it enters the deethanizer 6. The deethanizer 6 operates under total reflux, and there is no product output at the top of the column. The bottom distillate of the deethanizer is boosted by the deethanizer bottom product pump 9, mixed with the main fractionator top oil, and then successively passes through the heat exchanger 15 for heat exchange with the bottom liquid of the debutanizer 20 and the heat exchanger 16 for heat exchange with the jet fuel product. After being heated to about 128°C, it enters the debutanizer 20.

[0046] A deethanizer bottom hot water reboiler 7 and a deethanizer bottom steam reboiler 8 are provided at the bottom of the deethanizer 6. Due to the top pressure of about 960 KPag, the bottom temperature is as high as about 117°C, so the hot water reboiler 7 is shut down, and only the 0.5 MPa steam heat source reboiler 8 is started.

[0047] The operation process of the absorber 13 is as follows: The gas phase after cooling and separation at the top of the deethanizer 6 enters the absorber 13. The overhead distillate of the absorber 13 is cooled by the absorber top gas circulating water cooler 10 and then enters the absorber top gas-liquid separation tank 11 for equilibrium separation. The separated gas phase is discharged as a dry gas product, and the liquid phase is boosted by the absorber top reflux pump 12 and then returned to the absorber 13. In addition, the material at the bottom of the absorber 13 is refluxed to the upstream of the deethanizer top gas circulating water cooler 1 by the absorber bottom pump 14.

[0048] In debutanizer 20, liquefied petroleum gas (LPG) products are produced at the top of the column. After the overhead distillate is cooled by debutanizer overhead gas air cooler 17, it enters debutanizer overhead gas-liquid separator 18 for equilibrium separation. The gas phase is discharged as LPG products, and the liquid phase is returned to debutanizer 20 after being boosted by debutanizer overhead product reflux pump 19. The bottom liquid is boosted by debutanizer bottom product pump 22, first heated by heat exchanger 15 using its own feed, and then divided into two streams: one stream is sent to absorber 13 as the absorbent to ensure the qualified composition of the product dry gas; the other stream is sent to naphtha splitter 26. Since the top pressure of debutanizer 20 is relatively high (about 1150 KPag), the bottom reboiler 21 of the debutanizer needs to use the middle reflux of the main fractionator (extraction temperature about 275°C) as the heat source. Also, because part of the bottom liquid is recycled for absorber 13, the processing capacity of debutanizer 20 is large, the heat load of the reboiler is high, and the energy consumption is relatively large.

[0049] In naphtha splitter 26, light naphtha products are produced at the top of the column. After the overhead distillate is cooled by naphtha splitter overhead gas air cooler 23, it enters naphtha splitter overhead gas-liquid separator 24 for equilibrium separation. Part of the liquid phase is returned to naphtha splitter 26 after being boosted by naphtha splitter overhead product reflux pump 25, and the other part is used as light naphtha products. Heavy naphtha products are produced at the bottom of the column and are discharged by naphtha splitter bottom product pump 28. The bottom reboiler 27 of the naphtha splitter usually uses diesel products as the heat source.

[0050] Table 1 shows the material balance of Comparative Example 1.

[0051]

[0052] Table 2 shows the composition of the product dry gas and LPG in Comparative Example 1.

[0053]

[0054]

[0055] Table 3 shows the density and distillation range of the product light naphtha and heavy naphtha in Comparative Example 1.

[0056] No. Item Unit Light naphtha Heavy naphtha 1 Density at 20℃ <![CDATA[kg / m 3 > 645 707 2 Distillation range 1% ℃ 24.4 83.4 3 5％ ℃ 25.5 94.4 4 10％ ℃ 26.9 99.4 5 20％ ℃ / 103.1 6 30％ ℃ 32.7 106.8 7 40％ ℃ / 111.1 8 50％ ℃ 35.4 116.2 9 60％ ℃ / 122.1 10 70％ ℃ 45.8 128.8 11 80％ ℃ / 136.1 12 90％ ℃ 63.6 145.8 13 95％ ℃ 71.2 160.5 14 98％ ℃ 76.5 166.4

[0057] It can be seen that the products all meet the quality requirements.

[0058] Table 4 shows the main parameters of each column in the absorption and stabilization system of Comparative Example 1.

[0059]

[0060]

[0061] It can be seen from Table 4 that the total cooling load of each column in Comparative Example 1 is 1230.6×104 kcal / h, the total reboiler load is 2235.5×10 4 kcal / h. Based on the standard oil price of 3500 yuan / ton, the 0.5MPag steam price of 180 yuan / ton, the 1.0MPag steam price of 200 yuan / ton, the 3.5MPag steam price of 250 yuan / ton, and the 110 - 130°C hot water price of 0.7 yuan / 10 4 kcal, and the device operates 8400 hours a year (the same below). Without considering the power consumption cost of pumps, the total energy consumption of Comparative Example 1 converted to the cost of steam is 51.853 million yuan / year.

[0062] Example 1

[0063] In this example, the raw materials used are the same as those in the previous section, including the gas from the stripping tower top, the oil from the stripping tower top, the oil from the pre-fractionation tower top, and the oil from the main fractionation tower top. Compared with Comparative Example 1, this example has been improved in many aspects. The naphtha separation tower is changed to a heavy naphtha tower, the structural sequence with the debutanizer is re-planned, and the feed treatment method is changed. The specific process is as follows:

[0064] Refer to Figure 2 , the oil from the stripping tower top, the oil from the pre-fractionation tower top, and the oil from the main fractionation tower top are first boosted by the mixed oil feed pump 23, and then heated to 98°C successively through the hot water / feed heat exchanger 24 and the heavy naphtha / feed heat exchanger 25, and then enter the flash tank 26. This process is because there are some volatile light components dissolved in the feed. To ensure the normal operation of the heavy naphtha tower top cooler 27, such treatment is required. After flashing, the liquid phase enters the heavy naphtha tower 30 and feeds from the 13th tray. After fractionation, the heavy naphtha product is obtained at the bottom of the tower, and a mixture of light naphtha and liquefied gas is obtained at the top of the tower; the gas phase is injected into the top gas phase extraction line of the depropanizer 6 and enters the deethanizer top gas circulating water cooler 1 together with the gas from the stripping tower top. Among them, the product quality of the heavy naphtha can be adjusted according to parameters such as tower pressure, extraction volume, and reflux ratio to meet the separation requirements. Two new heat exchangers 24, 25 and a flash tank 26 are added to increase the feed temperature of the heavy naphtha tower 30 and recover the heat of the heavy naphtha, and at the same time remove some light components.

[0065] The cooled mixed gas enters the deethanizer top gas-liquid separation tank 2. After equilibrium separation, the gas phase is sent to the absorption tower 13, and the liquid phase is boosted by the deethanizer top product reflux pump 3. Then, the liquid phase exchanges heat with the gas phase product of the flash tank 26 through the heat exchanger 4 and exchanges heat with the light naphtha product through the heat exchanger 5, and is heated to about 63°C and then enters the deethanizer 6. The deethanizer 6 operates under total reflux, and there is no product output at the top of the tower. The bottom distillate of the tower is boosted by the deethanizer bottom product pump 9, mixed with the heavy naphtha tower top oil, and then exchanges heat with the jet fuel product through the heat exchanger 15, and is heated to about 120°C and then enters the debutanizer 19.

[0066] At the bottom of the deethanizer 6, there are a hot water reboiler 7 and a steam reboiler 8 for the bottom of the deethanizer. Since the separated material system becomes lighter and the bottom temperature drops to about 98°C, the 0.5 MPa steam heat source reboiler 8 can be shut down, and only the hot water heat source reboiler 7 at 130°C is turned on. The top gas of the deethanizer 6 is cooled to about 45°C by water cooling and then enters the liquid separation tank 2. The gas-phase product distills out from the top of the tank and enters the absorption tower 13, and all the liquid phase returns to the tower as reflux. And to reduce the reboiling load at the bottom of the tower, the top reflux is heated to about 80°C and then returns to the tower. The vaporization rate of the circulating heated material in the bottom reboiler 7 of the deethanizer is controlled between 15% and 25%. The bottom product is pumped by the pump 9 and sent to the debutanizer 19.

[0067] The working process of the absorption tower 13 is as follows: The gas phase after cooling and separation at the top of the deethanizer 6 enters from the bottom of the absorption tower 13. Part of the light naphtha from the bottom of the debutanizer is injected into the gas-phase product line at the top of the tower as an absorbent. After absorption, the dry gas product is obtained at the top of the tower and sent to the fuel system, and the rich absorbent oil at the bottom of the tower goes to the deethanizer top for recycling. The operating pressure of the absorption tower is 900 - 1000 KPag. The top distillate of the absorption tower 13 is cooled by the absorption tower top gas circulating water cooler 10 and then enters the absorption tower top gas-liquid separation tank 11 for equilibrium separation. The gas phase is discharged as the dry gas product, and the liquid phase is returned to the absorption tower 13 after being boosted by the absorption tower top reflux pump 12. In addition, the material at the bottom of the absorption tower 13 is refluxed to the upstream of the deethanizer top gas circulating water cooler 1 by the absorption tower bottom pump 14.

[0068] In the debutanizer 19, the liquefied petroleum gas (LPG) product is produced at the top of the tower, and the light naphtha product is obtained at the bottom of the tower. The top distillate is cooled by the debutanizer top gas air cooler 16 and then enters the debutanizer top gas-liquid separation tank 17 for equilibrium separation. The gas phase is discharged as the liquefied gas product, and the liquid phase is returned to the debutanizer 19 after being boosted by the debutanizer top product reflux pump 18. The bottom liquid is boosted by the debutanizer bottom product pump 21, first undergoes heat exchange through the heat exchanger 5, and then is divided into two streams: one stream is sent to the absorption tower 13 as an absorbent through the absorbent circulating water cooler 22 to ensure the qualified composition of the product dry gas; the other stream is discharged as the light naphtha product. The operating pressure of the debutanizer 19 is 800 - 1200 KPag. The debutanizer top gas is cooled to 40°C by air cooling and enters the liquid separation tank. After being pumped, part of the liquid phase returns as reflux to the top of the debutanizer, and the other part is sent to the tank farm. The vaporization rate of the circulating heated material in the debutanizer bottom reboiler is controlled between 15% and 25%. Since the bottom product does not contain heavy naphtha and the components become lighter, the bottom temperature drops, and the product diesel at 239°C is used as the reboiling heat source. The bottom distillate is pumped and divided into two paths. One path is cooled to 40°C by water cooling and then sent to the top of the absorption tower as an absorbent, and the other path leaves the device as a raw material for hydrogen production or gasoline blending.

[0069] After the distillate from the top of the heavy naphtha tower 30 is cooled to 45°C by the air cooler 27 for the overhead gas of the heavy naphtha tower, it enters the gas-liquid separation tank 28 at the top of the heavy naphtha tower for equilibrium separation. A part of the liquid phase is returned to the heavy naphtha tower 30 after being boosted by the reflux pump 29 for the products at the top of the heavy naphtha tower, and the other part converges with the bottom distillate of the deethanizer 6. The operating pressure of the heavy naphtha tower is 250 KPag, and the operating pressure can be adjusted according to the specific operating conditions of the unit. By adjusting the air cooling and water cooling operations at the top of the tower, the operating temperature of the liquid separation tank 28 at the top of the tower is controlled between 38°C and 50°C (maintaining this range in summer to ensure stable operation, for example, the lowest temperature is 42°C, and it can be appropriately reduced in winter, for example, 38°C). The vaporization rate of the circulating heated material in the reboiler 31 at the bottom of the heavy naphtha tower is controlled between 15% and 25%. After the heavy naphtha product at the bottom of the tower is pressurized by the pump 32, one stream is cooled to 40°C - 42°C by air cooling and water cooling and enters the tank farm, and the other stream is directly sent to the continuous reforming unit as raw material. Due to the increase in the feed flow rate and the heavier material system in the heavy naphtha tower, the reboiling load at the bottom of the tower has increased significantly, so the middle reflux of the main fractionator (the extraction temperature is about 275°C) is used as the heat source for the reboiler 31.

[0070] Table 5 shows the compositions of the dry gas and liquefied gas of the product in Example 1.

[0071] No. Component Unit Dry gas Liquefied gas No. Component Unit Dry gas Liquefied gas 1 <![CDATA[H2]]> ％v 53.96 / 12 n-Butane ％v 0.73 38.54 2 <![CDATA[O2]]> ％v / / 13 Isobutane ％v 3.11 48.11 3 <![CDATA[N2]]> ％v / / 14 Trans-butene ％v / 0.06 4 CO ％v / / 15 n-Butene ％v / 0.03 5 <![CDATA[CO2]]> ％v / / 16 Isobutene ％v / 0.05 6 <![CDATA[H2S]]> ％v 4.77 0.05 17 Cis-butene ％v / 0.05 7 Methane ％v 8.99 / 18 Isopentane ％v 1.03 0.48 8 Ethane ％v 15.75 / 19 n-Pentane ％v 0.48 0.01 9 Ethylene ％v 0.02 / 20 <![CDATA[≥C6]]> ％v 0.67 / 10 Propane ％v 10.92 12.51 21 Total ％v 100 100 11 Propylene ％v 0.12 0.10

[0072] Table 6 shows the densities and distillation ranges of the light naphtha and heavy naphtha of the product in Example 1.

[0073]

[0074]

[0075] It can be seen that the products all meet the quality requirements.

[0076] Table 7 shows the main parameters of each tower in the absorption and stabilization system of Example 1.

[0077] No. Component Unit Deethanizer Absorber Debutanizer Heavy naphtha column 1 Top pressure kPag 980 970 1150 250 2 Bottom pressure kPag 990 980 1170 270 3 Reflux ratio kg / kg Total reflux Total reflux 2.74 1.5 4 Top temperature ℃ 76.4 49.7 77.8 63.1 5 Bottom temperature ℃ 97.1 55 141.8 157.8 6 Reboiler return temperature ℃ 103.5 / 200.5 162.3 7 Top cooling load <![CDATA[10 4 kcal]]> 98 17 143.8 675 8 Bottom reboiler load <![CDATA[10 4 kcal]]> 117 / 476.6 1067.3 9 Theoretical number of plates 25 10 30 30 10 Feed plate location 1 10 14 13

[0078] It can be seen from Table 7 that the total cooling load of each tower in the comparative example is 1645×10 4 kcal / h, and the total reboiling load is 1661×10 4 kcal / h. The heat load of the newly added hot water heat exchanger in Example 1 is 630×10 4 kcal / h. Calculated at the unit price of standard oil of 3500 yuan / ton, the unit price of 0.5 MPag steam of 180 yuan / ton, the unit price of 1.0 MPag steam of 200 yuan / ton, the unit price of 3.5 MPag steam of 250 yuan / ton, and the unit price of hot water at 110 - 130°C of 0.7 yuan / 10 4kcal. Without considering the electricity consumption cost of pumps, the total energy consumption cost of Example 1 is 39.351 million yuan per year, which is 12.502 million yuan per year lower than that of the comparative example, with a decrease of 24.1%.

[0079] Example 2

[0080] As Figure 3 shown, based on Example 1, the process of this example is improved. Considering that there are more dissolved light components in the overhead oil of the stripper and the overhead oil of the pre-fractionator, and less dissolved light components in the crude gasoline obtained from the main fractionator overhead. Therefore, first, the overhead oil of the stripper and the overhead oil of the pre-fractionator are mixed, heated by the hot water heat exchanger 24, and then enter the flash tank 26. After the liquid phase at the bottom of the flash tank is mixed with the crude gasoline obtained from the main fractionator overhead, it is heated and raised in temperature by the heavy naphtha heat exchanger 25 and then enters the heavy naphtha tower 30. Other processes are the same as those in Example 1.

[0081] Table 8 shows the compositions of the dry gas and liquefied gas of the product in Example 2.

[0082] No. Component Unit Dry gas Liquefied gas No. Component Unit Dry gas Liquefied gas 1 <![CDATA[H2]]> ％v 52.64 / 12 n-Butane ％v 1.45 35.55 2 <![CDATA[O2]]> ％v / / 13 Isobutane ％v 3.71 48.66 3 <![CDATA[N2]]> ％v / / 14 Trans-butene ％v / 0.05 4 CO ％v / / 15 n-Butene ％v / 0.03 5 <![CDATA[CO2]]> ％v / / 16 Isobutene ％v / 0.05 6 <![CDATA[H2S]]> ％v 4.47 0.21 17 Cis-butene ％v / 0.04 7 Methane ％v 9.2 / 18 Isopentane ％v 1.08 0.84 8 Ethane ％v 15.11 / 19 n-Pentane ％v 0.4 0.03 9 Ethylene ％v 0.01 / 20 <![CDATA[≥C6]]> ％v 0.3 / 10 Propane ％v 11.56 14.42 21 Total ％v 100.00 100.00 11 Propylene ％v 0.07 0.12

[0083] Table 9 shows the densities and distillation ranges of the light naphtha and heavy naphtha of the product in Example 2.

[0084]

[0085]

[0086] Table 10 shows the main parameters of each tower in the absorption and stabilization system of Example 2.

[0087] No. Component Unit Deethanizer Absorber Debutanizer Heavy naphtha column 1 Top pressure kPag 980 970 1150 250 2 Bottom pressure kPag 990 980 1170 270 3 Reflux ratio kg / kg Total reflux Total reflux 2.72 1.45 4 Top temperature ℃ 75.7 49.7 78 68.2 5 Bottom temperature ℃ 94.3 55 139.3 161.2 6 Reboiler return temperature ℃ 103.5 / 140.8 165 7 Top cooling load <![CDATA[10 4 kcal]]> 99.8 13 857 676 8 Bottom reboiler load <![CDATA[10 4 kcal]]> 122.8 / 452 1168 9 Theoretical number of plates 25 10 30 30 10 Feed plate location 1 10 14 13

[0088] It can be seen from Table 10 that the total cooling load of each tower in Example 2 is 1646×10 4 kcal / h, and the total reboiling load is 1743×10 4 kcal / h. The heat load of the newly added hot water heat exchanger in Example 2 is 523×10 4 kcal / h. Calculated at a standard oil price of 3500 yuan per ton, a 0.5MPag steam price of 180 yuan per ton, a 1.0MPag steam price of 200 yuan per ton, a 3.5MPag steam price of 250 yuan per ton, and a 110 - 130℃ hot water price of 0.7 yuan per 10 4 kcal. Without considering the electricity consumption cost of pumps, the total energy consumption cost of Example 1 is 40.402 million yuan per year, which is 11.451 million yuan per year lower than that of Comparative Example 1, with a decrease of 22.1%.

[0089] The above embodiments are the preferred embodiments of the present invention, but the embodiments of the present invention are not limited to the above embodiments. Any other changes, modifications, substitutions, combinations, and simplifications made without departing from the spirit and principle of the present invention shall be equivalent replacement methods and are all included in the protection scope of the present invention.

Claims

1. A gas oil hydrocracking unit absorption and stabilization system, characterized in that, Comprising: A heavy naphtha tower, the feed end of which receives a liquid-phase mixture obtained by mixing, heating up, and flash-vaporizing the stripped top oil, pre-fractionation top oil, and main fractionation top oil from an upstream wax oil hydrocracking unit. By adjusting the rectification operating conditions of the heavy naphtha tower, a liquefied gas and light naphtha mixture is obtained at the top of the tower, and a heavy naphtha product is obtained at the bottom of the tower; A de-ethanizer tower, the feed end of which receives a mixture of the flash-vaporized gas phase and the stripped top gas. By adjusting the rectification operating conditions of the de-ethanizer tower, dry gas is obtained at the top of the tower and sent to an absorber tower, and a light naphtha mixture of saturated light hydrocarbons is obtained at the bottom of the tower; An absorber tower, which is used to remove heavy components from the dry gas. By adjusting the rectification operating conditions of the absorber tower, dry gas product is obtained at the top of the tower, and a light naphtha mixture of saturated light hydrocarbons at the bottom of the tower is sent back to the de-ethanizer tower for recycling; A de-butane tower, the feed end of which receives the liquefied gas and the light naphtha mixture from the top of the heavy naphtha tower and the bottom of the de-ethanizer tower. By adjusting the rectification operating conditions of the de-butane tower, a liquefied gas product is obtained at the top of the tower, and a light naphtha product is obtained at the bottom of the tower.

2. The absorption and stabilization system of a wax oil hydrocracking unit according to claim 1, characterized in that The liquid-phase mixture is directly prepared by successively mixing, heating up, and flash-vaporizing the stripped top oil, the pre-fractionation top oil, and the main fractionation top oil.

3. A absorption and stabilization system of a wax oil hydrocracking unit according to claim 1, characterized in that, The liquid-phase mixture is prepared by first mixing and heating up and flash-vaporizing the stripped top oil and the pre-fractionation top oil, and then mixing the flash-vaporized liquid phase with the main fractionation top oil and heating up again for the second time.

4. A absorption and stabilization system of a wax oil hydrocracking unit according to claim 1, characterized in that, The number of theoretical plates of the heavy naphtha tower is 30, the feed end is located in the middle of the heavy naphtha tower, the operating pressure is 120 - 300 kPag, the reflux ratio is 1.3 - 1.8, and the reboiler heat source at the bottom of the tower is the middle section stream of the main fractionation tower.

5. The absorption and stabilization system of a wax oil hydrocracking unit according to claim 1, characterized in that The number of theoretical plates of the de-ethanizer tower is 25, the feed end is located at the top of the de-ethanizer tower, the operating pressure is 900 - 1000 kPag, and the reboiler heat source at the bottom of the tower is hot water.

6. A absorption and stabilization system of a wax oil hydrocracking unit according to claim 1, characterized in that, The number of theoretical plates of the absorber tower is 10, the dry gas is fed from the bottom of the tower, and the light naphtha is fed from the top of the tower as an absorbent, and the operating pressure is 900 - 1000 kPag.

7. A absorption and stabilization system of a wax oil hydrocracking unit according to claim 1, characterized in that, The number of theoretical plates of the de-butane tower is 30, the feed end is located in the middle of the de-butane tower, the operating pressure is 800 - 1200 kPag, the reflux ratio is 1.5 - 3.0, and the reboiler heat source at the bottom of the tower is the diesel product obtained from the main fractionation tower.

8. A absorption and stabilization system of a wax oil hydrocracking unit according to claim 1, characterized in that, The arithmetic mean temperature difference of the bottoms reboiler heat exchanger of the heavy naphtha column is ≥90 °C, the flow rate on the hot stream side is 350 - 480 t / h, the temperature after heat exchange is 220 - 250 °C, the flow rate of the cold side stream is 650 - 780 t / h, and the temperature returning to the column is 155 - 170 °C; the arithmetic mean temperature difference of the bottoms reboiler heat exchanger of the deethanizer column is ≥20 °C, the flow rate on the hot stream side is 200 - 260 t / h, the temperature after heat exchange is 110 - 130 °C, the flow rate of the cold side stream is 65 - 85 t / h, and the temperature returning to the column is 100 - 110 °C; the arithmetic mean temperature difference of the bottoms reboiler heat exchanger of the debutanizer column is ≥40 °C, the flow rate on the hot stream side is 60 - 90 t / h, the temperature after heat exchange is 130 - 145 °C, the flow rate of the cold side stream is 350 - 450 t / h, and the temperature returning to the column is 140 - 150 °C.

9. The separation method of the absorption and stabilization system of the wax oil hydrocracking unit according to any one of claims 1-8, characterized in that, Comprising the following steps: Mixing, heating up and flash vaporizing the top oil of the stripper column, the top oil of the pre-fractionator column and the top oil of the main fractionator column to obtain a liquid-phase mixture which enters the heavy naphtha column for fractionation; The gas phase generated by the flash drum is mixed with the top gas of the stripper column and then enters the deethanizer column for separation; The dry gas at the top of the deethanizer column enters the absorber to remove heavy components; The liquefied gas and light naphtha mixture at the top of the heavy naphtha column and the bottom of the deethanizer column enter the debutanizer column for separation.

10. The separation method according to claim 9, wherein The top oil of the stripper column, the top oil of the pre-fractionator column and the top oil of the main fractionator column are mixed, heated up and flash vaporized in sequence; or, first the top oil of the stripper column and the top oil of the pre-fractionator column are mixed, heated up and flash vaporized, and then the liquid phase after flash vaporization is mixed with the top oil of the main fractionator column and heated up for the second time.

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