A process for the recovery of coking heavy benzene constituents

Abstract

This invention discloses a process for recovering and utilizing heavy components of crude benzene from coking plants. It relates to the field of crude benzene refining technology, where the heavy benzene residue discharged during the hydrorefining process of crude benzene from coking plants is fed into the middle of a distillation column. The column operates at a gauge pressure of 0~0.05MPa, and the bottom material temperature is maintained at 220~240℃. After being extracted from the bottom, the material is fed into the tube side of a reboiler, while high-pressure steam is introduced into the shell side. The steam indirectly exchanges heat with the bottom material through the tube wall and is condensed and discharged. The material after heat exchange returns to the bottom. The saturation temperature of the steam at its pressure is higher than the bottom material temperature but not exceeding it by more than 25℃, and the reboiler heat flux density does not exceed 6kW / m². Components with boiling points lower than tetrahydronaphthalene are collected from the top of the column; after condensation, part of this is refluxed, and part is collected as crude benzene. Tetrahydronaphthalene and heavier components are collected from the bottom as wash oil, without side streams. This process sets an upper limit for the tube wall temperature based on the steam saturation temperature and reduces wall superheat, which helps to inhibit coking and does not introduce water into the system.

Description

Technical Field

[0001] This invention relates to the field of coking crude benzene refining technology, specifically to a process for recovering and utilizing heavy components of coking crude benzene. Background Technology

[0002] Coking crude benzene can be hydrorefined to obtain pure benzene, toluene, xylene, and other aromatic products. During the feedstock vaporization pretreatment stage of hydrorefining, the non-vaporizable residue in the feedstock is discharged from the bottom of the automatic backwash filter, pre-evaporator, staged evaporator, and pre-reactor. After flash evaporation in the residual oil flash tank, it is collected from the bottom, which is the heavy benzene residue oil, accounting for 5% to 7% of the crude benzene content. The heavy benzene residue oil is dark brown, has a strong pungent odor, a high content of unsaturated compounds, and poor fluidity. It contains both recoverable light components such as benzene, toluene, xylene, and trimethylbenzene, as well as heavy components such as tetrahydronaphthalene, naphthalene, and methylnaphthalene. Most companies blend the heavy benzene residue oil as crude fuel oil, recover a small amount of benzene, xylene, and trimethylbenzene through simple distillation, and then sell it at a low price. A few use it for the synthesis of coumarone-indene resin, but the light components are not recycled into the crude benzene refining unit. The article "Development and Application of Heavy Benzene Distillation Technology" published in the 29th issue of "Grassroots Construction" in 2019 has the same record of the above situation, and points out that some enterprises have built heavy benzene residual oil primary processing units, which utilize the differences in boiling points of each component and extract industrial naphthalene, methylnaphthalene, etc. by changing the operating pressure, while the remaining components are processed into light oil, wash oil, heavy oil, asphalt or solvent oil.

[0003] When recovering light components from heavy benzene residue oil by cutting it according to boiling point using a distillation column, the bottom of the column must maintain a high operating temperature and obtain corresponding heat due to the high boiling point of the heavy components. In the coking industry, the bottom heating of such columns has long been done using tubular furnaces: the bottom material is drawn out by a circulating pump, sent to the tubular furnace for heating, and then returned to the bottom of the column. Chinese invention patent application CN104403704A describes in its background section that the atmospheric pressure crude benzene distillation process is mainly a steam stripping distillation process that uses a tubular furnace to heat the wash oil. It operates at a high temperature of 85~185℃, and each ton of crude benzene or light benzene requires approximately 1000~1500 kg of steam and 600~900 Nm³ of gas, resulting in high operating costs. The bottom of industrial naphthalene distillation columns also uses the same configuration, that is, the bottom residual oil is drawn out by a hot oil circulating pump, pumped into a tubular furnace for heating, and then returned to the column for heating.

[0004] Tubular furnaces have the following limitations when processing materials such as heavy benzene residue. The furnace tubes are radiantly heated by the flame side, resulting in a significantly higher temperature on the inner wall of the tubes than the bulk temperature of the material inside. Heavy benzene residue has a high content of unsaturated compounds and poor fluidity, leading to a long residence time near the high-temperature tube wall. This makes it prone to polymerization and condensation on the inner wall, forming coke. Once coking occurs, heat transfer deteriorates, requiring the furnace temperature to be increased to maintain the outlet temperature, further raising the tube wall temperature and exacerbating coking. Detached coke can also clog pipes and the bottom extraction system, forcing the unit to shut down for cleaning. Furthermore, the combustion exhaust gas from tubular furnaces contains nitrogen oxides.

[0005] To reduce the operating temperature required at the bottom of the distillation column, the art has generally turned to negative pressure distillation. The aforementioned CN104403704A employs a negative pressure benzene removal column and states that the amount of steam used for distillation is reduced by approximately 50% under a certain vacuum. The article "Development and Application of Heavy Benzene Distillation Technology" also advocates for negative pressure distillation, arguing that "the boiling points of all components in heavy benzene are relatively high, direct distillation consumes a lot of energy, and the separation is incomplete." However, negative pressure operation requires a vacuum pump or ejector and its associated piping, thus increasing the complexity of the equipment.

[0006] CN104403704A further replaces tubular furnace heating with steam heating. This involves steam-blowing the oil in the regenerator and then using oil-containing steam or fresh superheated steam from the regenerator to blow away benzene compounds in the benzene removal tower. In this method, steam enters the system and comes into direct contact with the material. The material escaping from the top of the tower therefore contains water vapor, which must be condensed and separated into gas-liquid and oil-water components to obtain the product. The separated water contains phenols and benzenes and requires further treatment. If the resulting fraction is intended to be returned to the hydrorefining unit as feedstock, the water must first be removed.

[0007] In summary, for heavy benzene residue oil with high unsaturated compound content, poor fluidity, and a tendency to coke at high temperatures, how to provide sufficient heat to the bottom of the column while simultaneously suppressing coking on the heat exchange surface and making the distillate obtained at the top of the column easy to reuse directly remains a problem to be solved in this field. Summary of the Invention

[0008] The purpose of this invention is to provide a recycling process for coking crude benzene heavy fractions that are high in unsaturated compounds, have poor fluidity, and are prone to coking at high temperatures. This process aims to suppress coking on the heat exchange surface while providing the heat required for distillation to the bottom of the column, and to make the fraction obtained at the top of the column easy to reuse directly.

[0009] The overall concept of this invention is to shift the focus of suppressing coking not by lowering the bulk temperature of the material, but by the superheat of the heat exchange surface wall. When steam condenses and releases heat, its temperature remains constant at the saturation temperature under that pressure. Therefore, an upper limit, uniquely determined by the steam pressure, can be set for the wall temperature of the heat exchange tubes. Below this upper limit, another upper limit is set for the heat flux density (heat passing through a unit heat exchange area), further constraining the superheat on the heat exchange surface. The required heat exchange area is not an independent variable defined by these two limits, but is obtained by dividing the heat load by the heat flux density. Since this upper limit of the wall temperature does not change with the heat load or the degree of coking, the path of "coking leading to poor heat transfer, thus requiring an increase in the heat source temperature, and consequently, a rise in wall temperature" no longer exists. Simultaneously, heat is transferred indirectly, steam does not enter the system, and the top fraction is therefore water-free.

[0010] The technical solution of this invention is as follows: Heavy benzene residue discharged during the coking crude benzene hydrogenation refining process is fed into the distillation column from the middle section; the distillation column is operated at a gauge pressure of 0~0.05MPa, and the temperature of the bottom material is maintained at 220~240℃; the bottom material is drawn from the bottom of the column and fed into the tube side of the reboiler; high-pressure steam is introduced into the shell side of the reboiler; the high-pressure steam indirectly exchanges heat with the bottom material in the tube side through the heat exchange tube wall and condenses before being discharged; the bottom material after heat exchange returns to the bottom of the column; the high-pressure steam, under its pressure... The saturation temperature is higher than the bottom material temperature by no more than 25°C, and the heat flux density of the reboiler is no higher than 6 kW / m². The vapor phase of the component with a boiling point lower than tetrahydronaphthalene is collected from the top of the column, and after condensation, part of it is refluxed to the top of the distillation column, and the other part is collected as crude benzene. The tetrahydronaphthalene and the component with a boiling point higher than tetrahydronaphthalene are collected from the bottom of the column, and after cooling, they are collected as wash oil. The streams collected from the distillation column are only the above-mentioned top stream and bottom stream, and no side stream fractions are collected from the distillation column during the distillation process.

[0011] In the above scheme, the effects of each constraint are as follows: Steam flows through the shell side, and the bottom material flows through the tube side, separated by the tube walls. This ensures that the wall temperature has an upper limit uniquely determined by the steam pressure, and prevents steam from entering the system. The upper limit of the difference between the saturation temperature and the bottom material temperature limits the wall superheat of the heat exchange surface; the upper limit of the heat flux density limits the heat transferred per unit heat exchange area. These two are related to the overall heat transfer coefficient of the reboiler—the heat flux density equals the product of the overall heat transfer coefficient and the temperature difference. The overall heat transfer coefficient depends on the reboiler structure and boiling conditions and is not a constant; therefore, limiting only the temperature difference is insufficient to limit the heat flux density: structures with a higher overall heat transfer coefficient provide a greater heat flux density at the same temperature difference. Thus, the two constraints each have an independent effect—the upper limit of the temperature difference limits the extent to which the wall temperature can exceed the bulk material temperature, while the upper limit of the heat flux density limits the product of the overall heat transfer coefficient and the temperature difference. The required heat exchange area is not an independent variable of these two constraints, but is obtained by dividing the heat load by the heat flux density. Atmospheric pressure operation keeps the bottom material temperature within a range that can be heated by steam at that pressure level with a small temperature difference, eliminating the need for a vacuum system. The cutting plane is positioned between tetrahydronaphthalene and heavier components with boiling points below tetrahydronaphthalene, and without side streams, allowing the entire column to complete the separation of key components with large boiling point differences only once.

[0012] Furthermore, the saturation temperature of the high-pressure steam at its pressure is 15-20°C higher than the temperature of the material at the bottom of the tower. Within this range, the heat transfer driving force is sufficient to complete the required heat transfer with a reasonable heat exchange area, while the wall superheat further decreases. When this range is below 15°C, the required heat exchange area increases sharply, and when it is above 20°C, the margin for wall superheat decreases accordingly. After this range narrows, the heat flux density also falls into a correspondingly narrow range given the overall heat transfer coefficient. Therefore, the value of heat flux density is not given separately in this section. For cases where the overall heat transfer coefficient is too high, the upper limit of the aforementioned heat flux density still provides a safety net.

[0013] Furthermore, the gauge pressure of the high-pressure steam is 3.0~4.2MPa, and the temperature of the bottom material is maintained at 225~235℃. The pressure level of the steam determines its saturation temperature, and thus determines the upper limit of the wall temperature. The values ​​within this pressure range must be selected in pairs with the bottom material temperature, so that the saturation temperature is higher than the bottom material temperature, but the difference does not exceed the aforementioned limit. The endpoints of the two ranges cannot be freely combined—there are combinations within the ranges that do not meet this condition. For example, if the bottom material temperature is 225℃ and the steam gauge pressure is 4.2MPa, the saturation temperature is 29.8℃ higher, exceeding the aforementioned limit. Such combinations should be excluded, and the judgment and selection order are detailed in the specific implementation method. In addition, under the order of magnitude of the total heat transfer coefficient in this embodiment of the invention, the limitation that the heat flux density is not higher than 6kW / m² means that this range actually does not exceed about 24℃. If a reboiler structure with a total heat transfer coefficient not higher than 240W / (m²·K) is used, this range can be taken up to 25℃ while the heat flux density still does not exceed the upper limit.

[0014] Furthermore, the heavy benzene residue oil is preheated in two stages before entering the tower. First, it is heated to 80-100°C by the condensate from the reboiler, and then heated to 150-170°C by the component collected from the bottom of the tower. The component collected from the bottom of the tower is cooled to 70-80°C after exchanging heat with the heavy benzene residue oil. The two-stage preheating is matched in reverse order of heat source temperature from high to low and material temperature from low to high, so that the condensate with a high temperature is used in the low-temperature section and the bottom product with a lower temperature and larger flow rate is used in the high-temperature section. If the order is reversed, the bottom product with a temperature of 230°C must be used to directly heat the cold material with a temperature of 40°C. The near-wall viscosity on the wash oil side of the heat exchange surface will increase significantly, which will deteriorate the heat transfer. Moreover, the temperature of the wash oil leaving the heat exchanger will be higher, and the amount of circulating water required for its subsequent cooling will increase accordingly.

[0015] Furthermore, the bottom material is drawn from the bottom of the column by a circulating pump and forcibly fed into the tube side of the reboiler. The circulation rate is 30 to 80 times the bottom product, and the circulating pump operates continuously throughout the entire process of introducing high-pressure steam into the reboiler. Forced circulation causes the flow velocity inside the tubes to be higher than that under natural convection conditions, thinning the boundary layer and shortening the residence time of the material near the high-temperature wall. If the pump is stopped while the reboiler is still under heat load, the flow velocity inside the tubes drops to the level of natural convection, and this residence time is significantly extended.

[0016] Furthermore, the distillation column has 10-20 trays, and 2-3 feed inlets are opened in the middle of the column. One of the feed inlets is selected according to the composition of the heavy benzene residue oil to feed the heavy benzene residue oil into the column. The number of trays is commensurate with the boiling point difference of the key component pairs corresponding to the aforementioned cut surface. By setting multiple feed inlets and selecting one according to composition, the composition of the feed is made as close as possible to the liquid phase composition on the tray at the selected feed inlet, thereby reducing the sudden change in concentration in the column caused by the feed.

[0017] Furthermore, the distillation column has 14 trays. This number of trays is matched with the boiling point difference of the aforementioned key component pairs and the range of values ​​for the ratio of reflux flow to crude benzene production, achieving a balance between the column height and the operating utility consumption.

[0018] Furthermore, the temperature of the vapor phase extracted from the top of the column is 110~130℃. This vapor phase, after condensation, enters the reflux tank, where the temperature is 50~60℃. The ratio of the reflux flow rate to the crude benzene output is 5~12. The temperature of the vapor phase at the top of the column and this ratio together determine the position of the cutting surface within the column; the temperature in the reflux tank keeps the condensate in the liquid phase and enables pumping. As this ratio increases, the improvement in separation tends to saturate, while the heat load on the reboiler and condenser increases accordingly; therefore, excessively high values ​​should not be used.

[0019] Furthermore, the total content of benzene, toluene, xylene, and trimethylbenzene in the heavy benzene residue is 30%~50%, and the total content of tetrahydronaphthalene and naphthalene is 20%~40%, with the sum of these two contents not exceeding 80%, all by mass. This limit defines the range of raw materials applicable to this process. Variations in the proportion of light and heavy components within this range can be accommodated by adjusting the feed inlet position and reflux ratio. An upper limit is set for the sum of the two contents because heavy benzene residue also contains components such as coumarone, indene, and methylnaphthalene. Raw materials with excessively high sums of these two contents are not suitable for this process.

[0020] Furthermore, the crude benzene collected from the top of the column is returned to the coking crude benzene hydrorefining unit as feedstock. This allows the recovered light components to re-enter the crude benzene refining process, and the fact that the top fraction is water-free is a condition for its direct reuse.

[0021] Beneficial effects Compared with the prior art, the beneficial effects of the present invention are as follows.

[0022] The bottom of the tower is heated indirectly by high-pressure steam via a reboiler. When the steam condenses and releases heat in the shell side of the reboiler, its temperature remains constant at the saturation temperature under that pressure. Therefore, the wall temperature of the heat exchange tubes does not exceed this saturation temperature, and this upper limit is uniquely determined by the steam pressure, independent of heat load fluctuations and the degree of coking on the heat exchange surface. In a tubular furnace, the furnace tube wall temperature is determined by radiation from the flame side. Coking of the furnace tubes reduces heat transfer, requiring an increase in furnace temperature to maintain the outlet temperature, leading to a rise in wall temperature. This invention avoids this path, which is beneficial for suppressing the polymerization and condensation of unsaturated compounds on the heat exchange surface.

[0023] The steam saturation temperature must be higher than the bottom material temperature by no more than 25°C, and the reboiler heat flux density must be controlled below 6 kW / m². Wall superheat is the direct driving force for the polymerization of unsaturated compounds on the heat exchange surface. The former limitation restricts the allowable difference in wall temperature relative to the bulk material temperature; the latter limitation restricts the heat transfer per unit heat exchange area. Heat flux density equals the product of the reboiler's overall heat transfer coefficient and this temperature difference. These two limitations are independent and indispensable: the overall heat transfer coefficient depends on the reboiler's structure and boiling conditions and is not a constant. Therefore, limiting only the temperature difference will still result in a higher heat flux density for structures with a higher overall heat transfer coefficient; conversely, limiting only the heat flux density will allow for a larger temperature difference and a higher upper limit for the wall temperature for structures with a lower overall heat transfer coefficient. Only by limiting both can the upper limit of the wall temperature and the heat transfer per unit area be controlled simultaneously. The required heat exchange area is not determined by these two variables, but by dividing the heat load by the heat flux density—the lower the temperature difference, the larger the area required.

[0024] The steam and the bottom material are separated by the walls of the heat exchange tubes, and no water is introduced into the system throughout the process. When heating by direct steam blowing, steam enters the system, and the top material contains water vapor. After condensation, it must undergo oil-water separation to obtain the product. The separated water contains phenols and benzenes and requires further treatment. If the obtained fraction is to be returned to the hydrorefining unit, it must be dehydrated first. In this invention, the top fraction is water-free, no oil-water stratification occurs in the reflux tank, no oil-water separation is required, and no separation wastewater is generated. The crude benzene obtained at the top can be directly returned to the coking crude benzene hydrorefining unit as raw material.

[0025] The distillation column operates at atmospheric pressure, and the apparatus does not include a vacuum pump or ejector and its associated piping. In the art, to reduce the operating temperature at the bottom of the column, negative pressure distillation is commonly employed, citing the reduction in steam consumption as a rationale. This invention operates at atmospheric pressure, thus eliminating the need for a complete vacuum system. Furthermore, since the wall temperature difference is limited to within 25°C, when the bottom material temperature is 230°C, only steam with a saturation temperature not exceeding 255°C is required. In this embodiment, a saturation temperature of 247°C is used, resulting in a lower pressure rating than that used in the comparative example (see Example 4). A lower pressure rating results in a higher latent heat of vaporization for the steam; therefore, with only a 3.8% difference in reboiler heat load between the two schemes, the mass consumption of heating steam is reduced by 12.5%. This reduction is not entirely attributable to the pressure rating; a breakdown is provided in Example 4.

[0026] The column collects components with boiling points lower than tetrahydronaphthalene from the top, and tetrahydronaphthalene and components with boiling points higher than tetrahydronaphthalene from the bottom, without collecting side stream fractions from the upper middle part of the column. This cut-off plane falls within the boiling point gap between indene and tetrahydronaphthalene, which is the widest gap in the boiling range from 1,3,5-trimethylbenzene to naphthalene. The critical component pair for this column is therefore indene and tetrahydronaphthalene, whose boiling point difference is greater than the boiling point difference between tetrahydronaphthalene and naphthalene when side stream fractions are used, thus reducing the number of trays required to achieve the same separation requirements.

[0027] The condensate from the reboiler is used as the heat source for the first stage of preheating, and the high-temperature wash oil from the bottom of the tower is used as the heat source for the second stage of preheating. When the condensate is discharged from the reboiler, it is still a saturated liquid at that pressure, and its sensible heat will be lost if it is directly discharged. The two stages of preheating are matched in reverse order, with the heat source temperature from high to low and the material temperature from low to high, so that the two streams of waste heat with different temperature levels are used in the heating section commensurate with their respective temperature levels.

[0028] The material at the bottom of the tower is pumped out by the circulating pump and forced into the tube side of the reboiler. The flow velocity inside the tube is higher than that under natural convection conditions, the boundary layer is thinned, the residence time of the material near the high-temperature wall is shortened, and the material near the wall is constantly replaced by the main flow. Together with the aforementioned upper limit of wall temperature and the limitation of heat flux density, this inhibits coking.

[0029] The device does not contain any open flame heating equipment, so it does not produce combustion exhaust gas or nitrogen oxides found in combustion exhaust gas. Attached Figure Description

[0030] Figure 1 This is a schematic diagram of the process for recovering and utilizing coking crude benzene components provided in an embodiment of the present invention. Detailed Implementation

[0031] The present invention will be further described below with reference to the accompanying drawings and embodiments.

[0032] In the various embodiments of this invention, heavy benzene residual oil refers to the non-vaporizable residue discharged from the bottom of the automatic backwash filter, pre-evaporator, segmented evaporator, and pre-reactor in the raw material vaporization pretreatment section during the coking crude benzene hydrorefining process. This residue is a heavy component collected from the bottom after flash evaporation in the residual oil flash tank, and its production amount is 5% to 7% of the crude benzene content. Crude benzene refers to the fraction obtained after condensation of components with boiling points lower than tetrahydronaphthalene collected from the top of the distillation column; wash oil refers to tetrahydronaphthalene and components with boiling points higher than tetrahydronaphthalene collected from the bottom of the distillation column. The above names have the same meaning in all embodiments.

[0033] In various embodiments of this invention, the wall temperature difference refers to the difference between the saturation temperature of the heating steam at its pressure and the temperature of the material at the bottom of the column. This difference is positive, meaning the saturation temperature must be higher than the temperature of the material at the bottom of the column. The saturation temperature is used here instead of the actual temperature of the superheated steam because the steam releases heat during condensation in the shell side of the reboiler, and the temperature of the condensation section is constant at the saturation temperature at that pressure. The superheated section occupies a very small proportion of the heat exchange surface, and the upper limit of the wall temperature of the heat exchange tubes is determined by the saturation temperature. In contrast, wall superheat refers to the difference between the wall temperature of the heat exchange tubes and the bulk temperature of the material at the bottom of the column. Since the wall temperature does not exceed the saturation temperature of the steam, the wall superheat does not exceed the wall temperature difference. Therefore, this invention uses the wall temperature difference as the upper limit of the wall superheat. Heat flux density refers to the heat load of the reboiler divided by its heat exchange area; according to heat transfer relationships, this quantity is equal to the product of the total heat transfer coefficient of the reboiler and the wall temperature difference. The definitions of the above quantities are consistent in all embodiments.

[0034] In the various embodiments of this invention, the aforementioned wall temperature difference limit applies to the reboiler. The material at the bottom of the column is at the highest temperature throughout the entire process on the reboiler heat exchange surface and boils on this surface, making it the most likely location for high-temperature polymerization. Although the two-stage heat exchangers for raw material preheating are also on the heating side, their heat sources are the condensate steam discharged from the reboiler and the material collected from the bottom of the column, respectively. The temperatures of both do not exceed the saturation temperature of the steam used in the reboiler. Therefore, the upper limit of the wall temperature of their heat exchange surfaces is constrained by the same saturation temperature as the reboiler. The main body temperature of the material does not exceed 160°C, far lower than the bottom of the column, and the degree of heating of the material near the high-temperature wall surface is also lower than that of the reboiler. Therefore, these two-stage heat exchangers are not separately checked against the wall temperature difference limit. The heat exchange surfaces of the condenser and residual oil cooler are on the cooling side, and the material does not undergo high-temperature polymerization on them; therefore, they are also not subject to this limit.

[0035] The wall temperature difference not exceeding 25℃ and the heat flux density not exceeding 6kW·m⁻² are the upper limits shared by Examples 1 to 3: the former limits the range by which the wall temperature can exceed the bulk material temperature, and the latter limits the amount of heat passing through a unit heat exchange area; under the same heat transfer capacity, the lower the wall temperature difference, the larger the heat exchange area required to achieve that heat transfer capacity. The following are representative configurations for Example 1: bottom material temperature 230℃, heating steam gauge pressure 3.7MPa, reboiler heat exchange area 82m², number of trays 14, and reflux flow rate to crude benzene output ratio 6. In Examples 2 and 3, if different values ​​are used, those values ​​shall prevail. Example 4 is a comparative example; the parameters listed are taken from the design document of the pressure reduction scheme and are used for comparison with this invention.

[0036] In the embodiments of this invention, high-pressure steam refers to water vapor with a gauge pressure of not less than 2.0 MPa. This lower limit is determined according to the temperature range of the bottom material in this process: the lowest temperature of the bottom material is 220°C, while the saturation temperature of the steam must be higher than the temperature of the bottom material, and the corresponding gauge pressure is already above 2.2 MPa. Therefore, all steam that meets the temperature difference limit of this process falls within this definition; the gauge pressure used in the embodiments of this invention is 3.0~4.2 MPa.

[0037] In various embodiments of the present invention, the middle section of the distillation column refers to the section between the rectification section and the stripping section, that is, the height section where the feed inlet is located; when feed is fed from this section, the upper tray constitutes the rectification section and the lower tray constitutes the stripping section.

[0038] In all embodiments of the present invention, the pressure values ​​are gauge pressures; the reflux ratio refers to the ratio of the reflux flow rate to the crude benzene production rate.

[0039] Example 1 The heavy benzene residue oil processed in this embodiment is a dark brown liquid with a pungent odor. It has a high content of unsaturated compounds, poor fluidity, and is prone to coking at high temperatures. Its composition by mass is as follows: benzene 24.243%, toluene 8.0425%, ethylbenzene 0.2453%, m-xylene and p-xylene combined 3.4659%, o-xylene 2.3477%, 1,3,5-trimethylbenzene 2.2158%, 1,2,4-trimethylbenzene 0.9809%, tetrahydronaphthalene 8.5276%, naphthalene 19.5604%, and the remaining 30.3709% consists of coumarone, n-decane, indene, indene, 1-methylnaphthalene, 2-methylnaphthalene, and small amounts of styrene and phenolic substances. The temperature of the heavy benzene residual oil when it is sent out from the storage tank in the storage area is 40℃ and at normal pressure. The transmission pipeline is equipped with heat tracing to maintain this temperature. The volumetric flow rate is 0.88 m³·h⁻¹, which is converted to a mass flow rate of 0.85 t·h⁻¹ based on its density of 0.9659 t·m⁻³. The annual operating time of the unit is 7200 h, and the annual processing capacity is 6120 t.

[0040] like Figure 1 As shown, the heavy benzene residue is first pressurized by the feed pump. The outlet pressure of this pump should be able to overcome the resistance drop between the subsequent two heat exchangers and the tower inlet, and meet the pressure requirements inside the tower. The pressurized heavy benzene residue enters the shell side of the feed preheater, which has a heat exchange area of ​​4.5 m². The tube side is fed with condensate from the reboiler, with a condensate temperature of 247°C. Here, the heavy benzene residue rises from 40°C to 90°C. The heat load of this stage of heat exchange is 22.4 kW. After releasing heat here, the condensate temperature drops to approximately 222°C (calculated as a decrease of approximately 25°C based on 700 kg·h⁻¹ and the specific heat of saturated water in this temperature range of 4.6 kJ·kg⁻¹·K⁻¹). After being depressurized by a steam trap, it is sent to the condensate recovery system. The tube side of the preheater is maintained at the pressure on the steam side to prevent flash evaporation of the condensate inside the tubes. Subsequently, the heavy benzene residue oil enters the bottom residual oil heat exchanger, where it exchanges heat with the 230°C wash oil drawn from the bottom of the tower. The temperature of the heavy benzene residue oil rises from 90°C to 160°C, while the temperature of the wash oil drops from 230°C to 116°C. The heat load of this stage of heat exchange is 31.4 kW. The specific heat of liquid in this temperature range is 1.9 kJ·kg⁻¹·K⁻¹, and that of the wash oil is 2.1 kJ·kg⁻¹·K⁻¹. The heat exchange area of ​​this heat exchanger is determined by the aforementioned heat load and the logarithmic mean temperature difference between the two ends, which is a conventional selection for liquid-liquid heat exchange. The two-stage preheating is matched in reverse order according to the heat source temperature from high to low and the material temperature from low to high: the low-temperature section is heated by steam condensate with a temperature as high as 247°C, and the high-temperature section is heated by the bottom wash oil with a lower temperature but a larger flow rate. The order cannot be reversed; if the bottom wash oil at 230°C is used to heat the cold material at 40°C first, the near-wall viscosity of the wash oil side on the heat exchange surface will increase significantly, which will deteriorate the heat transfer. Furthermore, the temperature of the wash oil when it leaves the heat exchanger will be too high, and the amount of circulating water required for its subsequent cooling will increase accordingly.

[0041] Preheated heavy benzene residue oil to 160°C is fed into the distillation column from the middle section. Exemplarily, the column has a diameter of 820 mm and a height of 18615 mm; it contains 14 valve trays. The column operates at a gauge pressure of 0.008 MPa, and the bottom material temperature is maintained at 230°C. Three feed inlets are located along the height of the column; in this embodiment, one of the inlets is positioned centrally. Other conventional trays, such as sieve trays, can also be used, selected based on the gas-liquid flow rate and allowable pressure drop within the column.

[0042] The bottom heating of the tower is carried out as follows: The bottom material is drawn from the bottom of the tower by a bottom circulation pump with a rated flow rate of 25 m³·h⁻¹; the drawn material is forcibly fed into the tube side of the reboiler, and the heat exchange area of ​​the reboiler is 82 m². Superheated steam with a gauge pressure of 3.7 MPa and a temperature of 270 °C is introduced into the shell side of the reboiler. The saturation temperature of the steam at this pressure is 247 °C; after the steam condenses and releases heat in the shell side, it is discharged as condensate at 247 °C and sent to the tube side of the aforementioned raw material preheater. The steam flow rate is 700 kg·h⁻¹. This steam cools from a superheated state of 270℃ to a saturated state of 247℃ and condenses into a isothermal condensate. The heat release per unit mass is 1795 kJ·kg⁻¹ (of which approximately 67 kJ·kg⁻¹ is from the superheated section, calculated based on the superheat of 2.9 kJ·kg⁻¹·K⁻¹ and a superheat of 23 K; the remainder is the latent heat of vaporization at this pressure). Based on this, the reboiler heat load is calculated to be 349 kW. Dividing this by the heat exchange area of ​​82 m², the heat flux density is 4.3 kW·m⁻². The steam saturation temperature of 247℃ is 17℃ higher than the bottom material temperature of 230℃, which is the wall temperature difference. The overall heat transfer coefficient of the reboiler, calculated from the heat load, heat exchange area, and wall temperature difference, is approximately 250 W / (m²·K) = 349 kW ÷ (82 m² × 17 °C). The precise value in this example is 250.4 W / (m²·K). The heat flux density is equal to the product of this overall heat transfer coefficient and the wall temperature difference. Therefore, for a given overall heat transfer coefficient, it is uniquely determined by the wall temperature difference and is independent of the heat exchange area. The heat exchange area is obtained by dividing the heat load by the heat flux density, which in this example is approximately 82 m² = 349 kW ÷ 4.256 kW·m⁻². The bottom material after heat exchange is returned to the bottom of the column. The circulation rate of the bottom circulation pump is 53 times the bottom output rate of 0.47 m³·h⁻¹.

[0043] The aforementioned bottom heating method has a triple effect. First, steam flows through the shell side, while the bottom material flows through the tube side, separated by the tube walls, preventing steam from contacting the material inside the tower. When steam condenses and releases heat, its temperature remains constant at the saturation temperature under that pressure. Therefore, the wall temperature of the heat exchange tubes will not exceed 247°C, and this upper limit is uniquely determined by the steam pressure, independent of heat load fluctuations or the degree of coking on the heat exchange surface. This differs from using an open-flame furnace as the bottom heat source: the furnace tube wall temperature of an open-flame furnace is determined by radiation from the flame side. Once the furnace tubes coke and heat transfer deteriorates, the furnace temperature must be increased to maintain the outlet temperature, further raising the wall temperature and creating a positive feedback loop that exacerbates coking. In this embodiment, the upper limit for the wall temperature is set at the steam saturation temperature, thus interrupting this positive feedback chain and helping to suppress coking on the heat exchange surface. Secondly, the product of the wall temperature difference and the overall heat transfer coefficient is the heat flux density, which in this embodiment is 4.3 kW·m⁻², not exceeding the upper limit of 6 kW·m⁻². Therefore, the heat exchange area is determined to be the heat load of 349 kW divided by this heat flux density, approximately 82 m². If the heat flux density is taken to the upper limit of 6 kW·m⁻², the heat exchange area corresponding to the same heat load is 58.2 m², which is the minimum heat exchange area under the conditions of this embodiment. The limits for wall temperature difference and heat flux density cannot be chosen arbitrarily: the overall heat transfer coefficient depends on the reboiler structure and boiling conditions, and is not a fixed value. When only the wall temperature difference is limited, structures with a higher overall heat transfer coefficient will still exhibit a larger heat flux density; when only the heat flux density is limited, structures with a lower overall heat transfer coefficient allow for a larger wall temperature difference and a higher upper limit for the wall temperature. Third, the bottom material is forcibly fed into the reboiler tubes by a circulating pump. The flow velocity inside the tubes is higher than that under natural convection conditions, resulting in a thinner boundary layer and a shorter residence time of the material near the high-temperature wall. The main flow continuously replaces the material near the wall. This circulating pump should operate continuously throughout the entire process of the reboiler being under heat load. If the pump is stopped while steam is still being supplied to the reboiler, the flow velocity inside the tubes will drop to the level of natural convection, and the residence time of the material near the high-temperature wall will be significantly extended.

[0044] Because the steam and the bottom material are separated by the pipe wall, no water is introduced into the system throughout the entire process in this embodiment. The column and the top fraction contain no water, and no oil-water stratification occurs in the reflux tank. Therefore, there is no need for oil-water separation, and no separation wastewater containing phenols or benzenes is generated. In this embodiment, the crude benzene collected from the top of the column is returned to the crude benzene tank in the crude benzene refining unit of the plant, serving as feedstock for the crude benzene hydrorefining unit.

[0045] The components collected from the top of the tower were determined according to the following criteria. The boiling points of the components of the heavy benzene residue oil processed in this example at atmospheric pressure are as follows: benzene 80.1℃, toluene 110.6℃, ethylbenzene 136.2℃, p-xylene 138.4℃, m-xylene 139.1℃, o-xylene 144.4℃, styrene 145.2℃, 1,3,5-trimethylbenzene 164.7℃, 1,2,4-trimethylbenzene 169.4℃, coumarone 174.0℃, n-decane 174.1℃, indene 176.5℃, indene 182.2℃, tetrahydronaphthalene 207.6℃, naphthalene 217.9℃, 2-methylnaphthalene 241.1℃, and 1-methylnaphthalene 244.7℃. Within the boiling point range between 1,3,5-trimethylbenzene and naphthalene, the interval between indene and tetrahydronaphthalene is 25.4°C, which is the widest point in this range. In this embodiment, the cutting surface is set within this interval, meaning that the component with a boiling point lower than tetrahydronaphthalene is collected from the top of the column, and tetrahydronaphthalene and the component with a boiling point higher than tetrahydronaphthalene are collected from the bottom of the column. Therefore, the key component pair of this column is indene and tetrahydronaphthalene, whose boiling point difference is 25.4°C, resulting in relatively high volatility. 14 trays are sufficient to meet the separation requirements. If a naphthalene oil side stream is set at the top of the column, the key component pair becomes tetrahydronaphthalene and naphthalene, whose boiling point difference is only 10.3°C, significantly increasing the required number of trays. In this embodiment, no naphthalene oil side stream is set; the streams collected from the column are only the top stream and the bottom stream, with the heavy components at the bottom being collected as wash oil.

[0046] The vapor phase at the top of the column is drawn off at 120°C and enters the tube side of the condenser. The heat exchange area of ​​the condenser is 51 m², and circulating water flows through the shell side. The supply temperature of the circulating water is 33°C, and the return temperature does not exceed 43°C. The condensed liquid enters the reflux tank at 55°C and is then pumped out by the reflux pump, which has a rated flow rate of 5 m³·h⁻¹. One stream of the liquid pumped out returns to the top of the column for reflux, and the other stream is collected as crude benzene. In this embodiment, the ratio of reflux flow rate to crude benzene collection rate is 6. The crude benzene collection rate is 0.43 m³·h⁻¹, with an outlet temperature of 55°C and a pressure of 0.35 MPa. Based on its density of 0.8779 t·m⁻³, the mass flow rate is 0.3775 t·h⁻¹, and the annual production is 2718 t. By mass, the top product accounts for 44.412% of the feed. In this embodiment, the total named components with boiling points lower than tetrahydronaphthalene in the heavy benzene residue amount to 41.541%, a difference of 2.871 percentage points. This difference comes from the remaining 30.3709% of the product, which contains components with boiling points lower than tetrahydronaphthalene, namely styrene, coumarone, n-decane, indene, indene, and phenols. The majority of the remaining 30.3709% consists of 1-methylnaphthalene and 2-methylnaphthalene. The above division is based on completely separating the key component pairs. In actual operation, there is a small amount of overlap between indene and tetrahydronaphthalene, the degree of which is controlled by the distillation range index of the top product. Therefore, the top product may vary slightly within the allowable range of this index.

[0047] The material at the bottom of the column is pumped out by the bottom discharge pump. The pumped material first enters the aforementioned bottom residual oil heat exchanger to exchange heat with the feed. After heat exchange, the temperature drops to 116℃, and then it enters the residual oil cooler. This cooler has a heat exchange area of ​​5.8 m², with circulating water flowing through the shell side. The material is cooled to 75℃ here and then sent to the product tank as wash oil. According to the total energy balance of the column, the heat load of the condenser is the difference between the heat load of the reboiler and the enthalpy difference between the inlet and outlet streams. Using 25℃ as the liquid phase enthalpy reference, a feed of 0.85 t·h⁻¹ carries in 60.6 kW at 160℃ (liquid phase heat is taken as 1.9 kJ·kg⁻¹·K⁻¹), a top product of 0.3775 t·h⁻¹ carries out 5.7 kW (taken as 1.8 kJ·kg⁻¹·K⁻¹) at 55℃, and a bottom product of 0.4725 t·h⁻¹ carries out 56.5 kW (taken as 2.1 kJ·kg⁻¹·K⁻¹) at 230℃. The enthalpy difference between the feed and output streams is 1.6 kW. Therefore, the heat load of the condenser is 349 kW minus 1.6 kW, approximately 347 kW. The condenser has both a condensation section and a subcooling section, and its overall heat transfer coefficient is lower than that of pure component condensation. The bottom product output is regulated by a level control valve on the pipeline according to the bottom liquid level, maintaining the bottom liquid level within the specified range. The wash oil production rate is 0.47 m³·h⁻¹, and the outlet pressure is 0.35 MPa. Based on its density of 1.0053 t·m⁻³, the annual output is 3402 t. The feed rate of 0.85 t·h⁻¹, the top rate of 0.3775 t·h⁻¹, and the bottom rate of 0.4725 t·h⁻¹ are consistent.

[0048] The operation control of this embodiment is carried out as follows. The temperature of the bottom material is measured by a thermocouple and maintained at 230°C by adjusting the steam flow rate of the reboiler. When the pressure rating of the steam used changes, it should be checked to ensure that its saturation temperature is higher than the bottom material temperature, and the difference does not exceed 25°C. If the difference exceeds 25°C, a steam pressure rating with a lower saturation temperature should be used, or the bottom material temperature should be increased to bring the difference back within the limit. The bottom liquid level is maintained by the liquid level regulating valve on the bottom product pipeline. When the liquid level drops to the lower limit, the product flow rate is reduced. If it is necessary to stop the steam supply to the reboiler for any reason, the steam should be cut off first, and the bottom circulation pump should be stopped only after the bottom material temperature drops below 150°C. The quality of crude benzene is determined by sampling and measuring its density and distillation range. The density should be 0.871~0.900 t·m⁻³, and the distillate volumetric percentage before 180℃ should not be less than 93%. If the distillate volume before 180℃ is less than 93%, the reflux ratio should be increased or the top product should be reduced. Because heavy benzene residue is prone to clogging, its liquid level is measured, for example, using a jacketed heated float level gauge or guided wave radar level gauge, and its flow rate is measured using a jacketed heated electric remote-transmission float flow meter.

[0049] The reboiler in this embodiment is a shell-and-tube heat exchanger. Any indirect heat exchange type that can separate the steam from the bottom material with a solid wall, such as a plate-and-shell heat exchanger, is also applicable to this invention; the selection is based on the structural compactness and cleanability while ensuring that the heat flux density does not exceed 6 kW·m⁻².

[0050] Example 2 The feed composition, feed rate, and annual operating time of this embodiment are the same as those of Embodiment 1. That is, the heavy benzene residue oil is composed as described in Embodiment 1, the processing capacity is 6120t, the annual operating time is 7200h, and the process sequence is also the same as that of Embodiment 1. The changes in this embodiment are the values ​​of the bottom material temperature, heating steam pressure, number of trays, and reflux ratio.

[0051] The temperature of the bottom material and the pressure of the heating steam cannot be selected independently; they must be determined in pairs. The saturation temperature of the heating steam at its pressure must be higher than the temperature of the bottom material; otherwise, heat cannot be transferred to the bottom of the column. At the same time, the difference between this saturation temperature and the bottom material temperature should not exceed 25°C; otherwise, the superheat on the wall will be too high. Table 1 lists five sets of feasible paired values.

[0052] Table 1. Paired values ​​of bottom material temperature and heating steam pressure The third group in Table 1, i.e., the configuration of Example 1, has its heat flux density calculated by multiplying the total heat transfer coefficient of 250.4 W / (m²·K) as stated in Example 1 by the wall temperature difference. The other groups are calculated in the same way. The steam saturation temperature and wall temperature difference in the table have been rounded to one decimal place, while the heat flux density is calculated using the unrounded values. Therefore, the result obtained by multiplying the listed temperature difference by the total heat transfer coefficient may differ from the listed heat flux density by a certain last digit. The first and fifth groups fall at the two ends of the bottom material temperature range of 220~240℃, respectively, and the heating steam gauge pressure also falls at the two ends of the range of 3.0~4.2 MPa, respectively. The wall temperature difference of all five groups does not exceed 25℃, and the heat flux density does not exceed 6 kW·m⁻². The wall temperature differences of the first and fourth groups are 15.8℃ and 19.8℃, respectively, which are close to the lower and upper ends of the range of 15~20℃. It should also be noted that the same steam pressure rating of 4.2 MPa corresponds to two different bottom material temperatures in the table, resulting in different wall temperature differences. Therefore, these two items must be determined in pairs.

[0053] The following two combinations are not acceptable. First, a bottom material temperature of 240℃ and a heating steam gauge pressure of 3.0 MPa: at this pressure, the steam saturation temperature is 235.8℃, which is 4.2℃ lower than the bottom material temperature, making heat transfer from the steam to the bottom material impossible. Second, a bottom material temperature of 220℃ and a heating steam gauge pressure of 4.2 MPa: at this pressure, the steam saturation temperature is 254.8℃, a difference of 34.8℃ from the bottom material temperature, exceeding the 25℃ limit, resulting in excessive superheat on the heat exchanger tube walls. Therefore, the correct order for selecting these values ​​is: first, determine the bottom material temperature according to the separation requirements, then select a steam pressure rating with a saturation temperature higher than that temperature, but not exceeding 25℃ above it.

[0054] The relationship between heat flux density and heat transfer area should be determined in the following order, and cannot be reversed. Heat flux density is equal to the product of the overall heat transfer coefficient of the reboiler and the wall temperature difference. Therefore, for a given overall heat transfer coefficient, it is uniquely determined by the wall temperature difference and is independent of the heat transfer area; the heat transfer area is obtained by dividing the heat load by the heat flux density. Taking Example 1 as an example, the overall heat transfer coefficient is 250 W / (m²·K), the wall temperature difference is 17°C, the heat flux density is 4.3 kW·m⁻², and the heat load of 349 kW divided by this heat flux density yields a heat transfer area of ​​approximately 82 m². The overall heat transfer coefficient is not a constant value; it depends on the reboiler structure, the flow velocity inside the tubes, and the boiling conditions. As the wall temperature difference increases, the heat transfer coefficient on the boiling side also increases, so the overall heat transfer coefficient also increases with the increase of the wall temperature difference. At the level of 250.4 W / (m²·K) in Example 1, the heat flux density reaches its upper limit of 6 kW·m⁻² when the wall temperature difference reaches approximately 24°C. Therefore, the actual wall temperature difference does not exceed approximately 24°C. However, when the overall heat transfer coefficient of the selected reboiler structure is not higher than 240 W / (m²·K), the heat flux density obtained by taking a wall temperature difference of 25°C is 6.0 kW·m⁻², still not exceeding the upper limit. At this point, both limits are simultaneously reached. For example, when the bottom material temperature is 230°C and the saturation temperature of the heating steam at its pressure is 255°C, the wall temperature difference is 25°C, the heat flux density is 6.0 kW·m⁻², and the heat transfer area obtained by dividing the heat load of 349 kW by this heat flux density is 58.2 m². It can be seen that the wall temperature difference of 25°C can be reached in reboiler structures with lower overall heat transfer coefficients, and the order in which it and the upper limit of heat flux density are reached is determined by which limit is reached first.

[0055] The operating pressure of the column can be selected between 0 and 0.05 MPa (gauge pressure). Increasing the operating pressure will raise the boiling point of each component in the column, and the temperature of the bottom material should also be increased. Therefore, the operating pressure and the bottom material temperature must be adjusted together. After adjustment, the wall temperature difference should still be checked according to Table 1.

[0056] The ratio of the number of trays to the reflux flow rate and the crude benzene yield is complementary and must be determined in pairs: when the number of trays is smaller, this ratio should be higher to achieve the same top product requirements as in Example 1; when the number of trays is larger, this ratio can be lower. Example 1 uses a configuration of 14 trays and a ratio of 6. Increasing this ratio has a saturation effect on separation improvement and will simultaneously increase the heat load on the reboiler and condenser, therefore it should not be too high; the selection between the number of trays and this ratio should be determined in conjunction with the investment in the column and the utilities required for operation.

[0057] In this embodiment, the gas phase temperature at the top of the column is 110~130℃, the temperature inside the reflux tank is 50~60℃, the feed is preheated to 150~170℃ before being sent into the column, and the wash oil collected from the bottom of the column is cooled to 70~80℃. All value groups are operated and controlled in accordance with the method described in Example 1.

[0058] Example 3 In this embodiment, the bottom material temperature is 230℃, the heating steam gauge pressure is 3.7MPa, and the reboiler heat exchange area is selected based on a heat flux density not exceeding 6kW·m⁻², all the same as in Embodiment 1; the process sequence is also the same as in Embodiment 1. The only changes in this embodiment are the composition of the feed and the location of the feed inlet.

[0059] This process is applicable to coking crude benzene heavy components with a total content of benzene, toluene, xylene, and trimethylbenzene of 30%~50% by mass and a total content of tetrahydronaphthalene and naphthalene of 20%~40% by mass. The heavy benzene residue oil processed in Example 1 has these two contents of 41.30% and 28.09%, respectively, falling within this range. This example processes two heavy benzene residue oils at opposite ends of this range: one is a light component-rich feedstock with a total content of 50% for benzene, toluene, xylene, and trimethylbenzene and a total content of 20% for tetrahydronaphthalene and naphthalene; the other is a light component-poor feedstock with these two contents of 30% and 40%, respectively. The sum of these two contents is limited to no more than 80%, with the remainder being components such as coumarone, indene, and methylnaphthalene; the sum of these two contents in the feedstock of Example 1 is 69.38%, while both of the above feedstocks have a content of 70%.

[0060] Regardless of fluctuations in feed composition, the cutting surface is always positioned between indene and tetrahydronaphthalene. This means that components with boiling points lower than tetrahydronaphthalene are collected from the top of the column, while tetrahydronaphthalene and components with boiling points higher than tetrahydronaphthalene are collected from the bottom. No naphthalene oil side stream is provided. The top collection rate increases with increasing light component content in the feed, while the bottom collection rate decreases accordingly.

[0061] The selection of feed inlets is based on the following criteria. Two to three feed inlets are installed along the height of the column. Feeds rich in light components have lower bubble points, which are similar to the liquid phase composition on the higher, cooler trays within the column; therefore, higher feed inlets are selected. Feeds leaning in light components have higher bubble points; therefore, lower feed inlets are selected. The selection criterion is to make the composition of the feed as close as possible to the liquid phase composition on the tray at the selected feed inlet, in order to minimize abrupt changes in concentration within the column caused by the feed.

[0062] Another operating parameter that needs to be adjusted according to the feed composition is the reflux ratio. When processing lean light component feedstocks, the overhead product yield is reduced, and a higher reflux ratio is used to maintain a distillation yield of at least 93% of the overhead product before 180°C; when processing light component-rich feedstocks, a lower reflux ratio can be used. The adjustment range for the reflux ratio is 5 to 12.

[0063] This embodiment uses the density and distillation range of the top product as the judgment indicators: samples are taken from the crude benzene pipeline after the reflux tank, and its density and the distillation rate before 180°C are measured. When the density is higher than 0.900 t·m⁻³ or the distillation rate before 180°C is lower than 93%, it indicates that the top fraction contains components with higher boiling points. In this case, the reflux ratio should be increased, or a lower feed inlet should be used, or both should be done simultaneously. After adjustment, samples should be taken and measured again until both indicators are within the requirements. When the density is lower than 0.871 t·m⁻³, it indicates that the top product yield is too low. In this case, the reflux ratio should be reduced or the top product yield should be increased.

[0064] If, after adjusting the feed inlet position and reflux ratio as described above, the two indicators of the top product still cannot fall within the requirements simultaneously, it indicates that the feed composition has exceeded the applicable range of this process. At this time, stop feeding into the tower and check the operating status of the upstream crude benzene hydrogenation refining unit. The material remaining in the tower should be handled according to the shutdown method described in Example 1, that is, first cut off the steam supply to the reboiler, and then stop the bottom circulation pump after the temperature of the bottom material drops below 150°C.

[0065] Example 4 This example is a comparative example, using a decompression operation, and is used to compare with Example 1.

[0066] The raw material source, raw material composition, processing scale, annual operating time, and product destination in this embodiment are the same as in Embodiment 1. The only changes are the tower operating pressure, and consequently, the bottom material temperature, heating steam grade, and reboiler heat exchange area. The parameters in this embodiment are taken from the equipment list in the pressure reduction scheme design document prepared during the feasibility study stage of the same device, and are at the same document level as the parameters in Embodiment 1, which are taken from the equipment list in the construction drawing stage design document of the same device.

[0067] The pressure reduction scheme operates at a gauge pressure of −0.055 MPa, with a bottom material temperature of 210℃. Steam with a gauge pressure of 5.0 MPa and a temperature of 265℃ is introduced into the shell side of the reboiler. The saturation temperature of the steam at this pressure is 265℃. The heat exchange area of ​​the reboiler is 15 m², and the steam flow rate is 800 kg·h⁻¹. The unit is also equipped with a liquid ring vacuum pump with a pumping capacity of 30 m³·h⁻¹ and a vacuum degree of −55 kPa, which is used to maintain the negative pressure inside the column.

[0068] The comparison between this embodiment and Embodiment 1 is listed in Table 2. Each item in the table represents the design parameters of the two schemes and the physical quantities calculated from them.

[0069] Table 2 Comparison of Example 1 and Example 4 project Example 4 (Decompression) Example 1 (at normal pressure) Tower operating pressure / MPa −0.055 0.008 Temperature of material at the bottom of the tower / °C 210 230 Heating steam gauge pressure / MPa 5.0 3.7 Steam saturation temperature / °C 265 247 Wall temperature difference / °C 55 17 Reboiler heat exchange area / m² 15 82 Reboiler heat load / kW 363 349 Heat flux density / kW·m⁻² 24.2 4.3 Steam flow rate / kg·h⁻¹ 800 700 Vacuum system Set up a liquid ring vacuum pump No As shown in Table 2, compared to the comparative example, the bulk temperature of the material at the bottom of the tower in Example 1 increased by 20°C, while the wall temperature difference of the reboiler decreased from 55°C to 17°C, a reduction of 38°C. The heat exchange area of ​​the reboiler increased from 15 m² to 82 m², which is 5.47 times the original size, while the heat load decreased from 363 kW to 349 kW. That is, while the heat exchange area increased, the heat transfer did not increase. The heat flux density decreased from 24.2 kW·m⁻² to 4.3 kW·m⁻², a reduction of approximately 82%. The total heat transfer coefficients of the two schemes are not the same: calculated from their respective heat loads, heat exchange areas, and wall temperature differences, the comparative example has a coefficient of approximately 440 W / (m²·K), while Example 1 has a coefficient of approximately 250 W / (m²·K). This difference is an inevitable result of reducing the wall temperature difference—the heat transfer coefficient on the boiling side increases with the increase of wall superheat. Therefore, this comparative example achieves a higher overall heat transfer coefficient at a wall temperature difference of 55°C, while in Example 1, the overall heat transfer coefficient decreases after the wall temperature difference is reduced to 17°C. This is precisely why the heat exchange area needs to be enlarged by 5.47 times: the area enlargement factor is not equal to the wall temperature difference compression factor, but rather the result of the combined effect of both and the change in the overall heat transfer coefficient. Simultaneously, the steam flow rate is reduced from 800 kg·h⁻¹ to 700 kg·h⁻¹, a reduction of 12.5%, and the entire vacuum system is eliminated. The reduction in steam flow rate must be understood in conjunction with the heat load: the reboiler heat loads of the two schemes are 363 kW and 349 kW, respectively, differing by only 3.8%; the reason for the 12.5% ​​reduction in steam flow rate is that the heat release per unit mass of the steam used in this invention increases from 1633 kJ·kg⁻¹ to 1795 kJ·kg⁻¹. The 162 kJ·kg⁻¹ increase consists of two parts: firstly, approximately 95 kJ·kg⁻¹ comes from the increased latent heat of vaporization due to the decrease in pressure level. This is a result of the present invention reducing the wall temperature difference, thereby allowing the use of steam with a lower pressure level. Secondly, approximately 67 kJ·kg⁻¹ comes from the fact that the steam used in this embodiment has a superheat of 23°C at 3.7 MPa (270°C to saturation temperature of 247°C), while the steam used in this comparative example is in a saturated state. This depends on the steam network conditions of the device and is irrelevant to the present invention. The pressure level can be reduced because the present invention limits the wall temperature difference to within 25°C—when the bottom material temperature is 230°C, only steam with a saturation temperature of 247°C is needed; this comparative example, however, uses steam with a saturation temperature of 265°C under the condition that the bottom material temperature is only 210°C, resulting in a wall temperature difference of 55°C.

[0070] The bulk temperature of the material and the wall superheat of the heat exchange surface are two different quantities; reducing the former does not equate to reducing the latter. In this comparative example, the bulk temperature of the material at the bottom of the column is reduced to 210°C by increasing the vacuum level, but the saturation temperature of the steam used is 265°C, and the heat exchange area of ​​the reboiler is only 15 m², resulting in a higher wall temperature difference and heat flux density than in Example 1. Example 1 accepts an increase in the bulk temperature of the material at the bottom of the column, instead using steam with a limited saturation temperature as the heat source and a larger heat exchange area, thereby reducing the wall superheat and helping to suppress coking on the heat exchange surface. The parameters, value ranges, and configuration methods in the above examples and comparative examples are merely illustrative and do not constitute a limitation on the scope of protection of this invention. Equivalent substitutions and improvements made by those skilled in the art based on the technical solutions of this invention all fall within the scope of protection of this invention.

[0071] Without conflict, the technical features in the above embodiments can be combined with each other.

Claims

1. A process for recovering and utilizing heavy benzene components from coking, wherein the heavy benzene components are heavy benzene residue discharged during the hydrorefining process of coking crude benzene, characterized in that, Includes the following steps: S1. The heavy benzene residue is fed into the distillation column from the middle section; S2. The distillation column is operated at a gauge pressure of 0~0.05MPa, and the temperature of the bottom material is maintained at 220~240℃; S3. The bottom material is drawn from the bottom of the column and fed into the tube side of the reboiler. High-pressure steam is introduced into the shell side of the reboiler. The high-pressure steam indirectly exchanges heat with the bottom material in the tube side through the heat exchange tube wall and is discharged after condensation. The bottom material after heat exchange is returned to the bottom of the column; the saturation temperature of the high-pressure steam at its pressure is higher than the temperature of the bottom material, and the difference is greater than the temperature of the bottom material. The temperature fluctuation range shall not exceed 25°C, and the heat flux density of the reboiler shall not exceed 6 kW / m²; S4. The gas phase of the component with a boiling point lower than tetrahydronaphthalene is collected from the top of the column, and after condensation, part of it is refluxed back to the top of the distillation column, and the other part is collected as crude benzene; S5. The tetrahydronaphthalene and the component with a boiling point higher than tetrahydronaphthalene are collected from the bottom of the column, and after cooling, they are collected as wash oil; wherein, the streams collected from the distillation column are only the top streams of step S4 and the bottom streams of step S5, and no side stream fractions are collected from the distillation column during the distillation process.

2. The coking crude benzene heavy component recovery and utilization process according to claim 1, characterized in that, In step S3, the saturation temperature of the high-pressure steam under its pressure is 15~20℃ higher than the temperature of the material at the bottom of the tower.

3. The coking crude benzene heavy component recovery and utilization process according to claim 1, characterized in that, The gauge pressure of the high-pressure steam in step S3 is 3.0~4.2MPa, and the temperature of the material at the bottom of the tower in step S2 is maintained at 225~235℃.

4. The coking crude benzene heavy component recovery and utilization process according to claim 1, characterized in that, In step S1, the heavy benzene residue oil is preheated in two stages before being fed into the tower: first, the heavy benzene residue oil is heated to 80~100°C with the steam condensate discharged from the reboiler in step S3, and then the heavy benzene residue oil is heated to 150~170°C with the component collected from the bottom of the tower in step S5; the component collected from the bottom of the tower in step S5 is cooled to 70~80°C after exchanging heat with the heavy benzene residue oil.

5. The coking crude benzene heavy component recovery and utilization process according to claim 1, characterized in that, In step S3, the bottom material is drawn from the bottom of the tower by a circulating pump and forcibly fed into the tube side of the reboiler. The circulation volume is 30 to 80 times the bottom material, and the circulating pump operates continuously throughout the process of introducing high-pressure steam into the reboiler.

6. The coking crude benzene heavy component recovery and utilization process according to claim 1, characterized in that, The distillation column has 10 to 20 trays, and 2 to 3 feed ports are opened in the middle of the distillation column. In step S1, the heavy benzene residual oil is fed into the column from one of the trays according to the composition of the heavy benzene residual oil.

7. The coking crude benzene heavy component recovery and utilization process according to claim 6, characterized in that, The distillation column has 14 trays.

8. The coking crude benzene heavy component recovery and utilization process according to claim 1, characterized in that, In step S4, the temperature of the gas phase extracted from the top of the tower is 110~130℃. After condensation, the gas phase enters the reflux tank, where the temperature is 50~60℃. The ratio of the reflux flow rate to the crude benzene output is 5~12.

9. The coking crude benzene heavy component recovery and utilization process according to claim 1, characterized in that, The total content of benzene, toluene, xylene and trimethylbenzene in the heavy benzene residue oil is 30% to 50%, and the total content of tetrahydronaphthalene and naphthalene is 20% to 40%, and the sum of these two contents does not exceed 80%, all by mass.

10. The coking crude benzene heavy component recovery and utilization process according to claim 1, characterized in that, The crude benzene extracted in step S4 is returned to the coking crude benzene hydrorefining unit as a feedstock for the unit.

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