WO2012055145A1 - 提高石油烃类馏分油收率的蒸馏塔和其进料方法 - Google Patents

提高石油烃类馏分油收率的蒸馏塔和其进料方法 Download PDF

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Publication number
WO2012055145A1
WO2012055145A1 PCT/CN2011/000665 CN2011000665W WO2012055145A1 WO 2012055145 A1 WO2012055145 A1 WO 2012055145A1 CN 2011000665 W CN2011000665 W CN 2011000665W WO 2012055145 A1 WO2012055145 A1 WO 2012055145A1
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Prior art keywords
distillation column
kpa
column
distillation
pressure
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PCT/CN2011/000665
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English (en)
French (fr)
Inventor
张占柱
毛俊义
候栓弟
秦娅
袁清
许克家
张同旺
王少兵
渠红亮
唐晓津
朱振兴
黄涛
Original Assignee
中国石油化工股份有限公司
中国石油化工股份有限公司石油化工科学研究院
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Application filed by 中国石油化工股份有限公司, 中国石油化工股份有限公司石油化工科学研究院 filed Critical 中国石油化工股份有限公司
Priority to GB1308855.4A priority Critical patent/GB2498500B/en
Priority to JP2013535241A priority patent/JP6000961B2/ja
Priority to US13/881,535 priority patent/US10544372B2/en
Publication of WO2012055145A1 publication Critical patent/WO2012055145A1/zh

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Classifications

    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G7/00—Distillation of hydrocarbon oils
    • C10G7/06—Vacuum distillation
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G7/00—Distillation of hydrocarbon oils
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G7/00—Distillation of hydrocarbon oils
    • C10G7/12—Controlling or regulating
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/10—Feedstock materials
    • C10G2300/1037—Hydrocarbon fractions
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/40—Characteristics of the process deviating from typical ways of processing
    • C10G2300/4012—Pressure
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/40—Characteristics of the process deviating from typical ways of processing
    • C10G2300/4025—Yield
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
    • C10G2400/02—Gasoline

Definitions

  • the present invention relates to a distillation column and method for increasing the distillate oil yield of a distillation apparatus, and more particularly to a distillation column and method for increasing the distillate oil yield of a heavy oil distillation process in the petroleum refining industry. Background technique
  • Distillation towers are widely used unit equipment in the petroleum refining industry.
  • some heavy oil products such as distillation of light distillate oil from oils such as crude oil and wax oil
  • the temperature of the distillation tower is higher, the heat source of the reboiler is high, not easy to obtain, and heavy oil The product is prone to thermal cracking at high temperatures. Therefore, the crude oil or heavy oil distillation column generally does not have a reboiler.
  • the heat source required for distillation is almost completely provided by the raw material.
  • the distillate oil is vaporized, and the vaporized distillate oil is from The top of the column and / or distilled from the side line, the unvaporized portion is distilled from the column.
  • Typical fractionation processes such as atmospheric distillation and vacuum distillation of crude oil.
  • the atmospheric and vacuum distillation of crude oil is the first process of crude oil processing. It supplies raw materials for the subsequent processing equipment of the refinery and directly supplies some products.
  • the basic process of crude oil distillation (for example, the fuel oil type) is that the crude oil is heated to about 220-260 ° C into the preliminary distillation column.
  • the first distillation column only takes one overhead product, that is, the reforming material or the light gasoline.
  • the normal process of the atmospheric tower is shown in Figure 1.
  • the bottom oil of the first distillation column is partially vaporized by heat exchange or normal pressure heating furnace 2, and then enters the atmospheric distillation column 8 through the oil transfer line 7, and the light components in the vaporization section of the distillation column.
  • Vaporization and ascending into the fractionation section the condensate of the reflux liquid is withdrawn from the top or side line to obtain the hydrazine oil, and the unvaporized part flows downward into the raking section, and the tray on the stripping section is in contact with the water vapor entering the bottom of the tower.
  • Get lighter components such as gasoline, kerosene, diesel, heavy diesel.
  • the unvaporized portion falls into the bottom of the tower and is taken up as an atmospheric residue.
  • the normal process of the vacuum distillation process is shown in Fig. 3.
  • the atmospheric residue is partially vaporized by heating in the vacuum heating furnace 2, and then enters the vacuum distillation column 6 through the oil transfer line 7, and the light component is vaporized in the vaporization section of the vacuum distillation column. It rises into the fractionation section, is condensed by the reflux liquid, and is withdrawn from the top or side line to obtain a distillate oil, and the unvaporized portion is taken out from the bottom of the column to obtain a vacuum residue.
  • the advantages and disadvantages of the design and operation of the crude oil distillation unit will have a great impact on the product quality, product yield and economic benefits of the refinery.
  • the extraction rate of the atmospheric distillation unit can be increased, so that the light components can be pulled out at the atmospheric pressure tower as much as possible, and no more into the vacuum distillation tower, on the one hand, more light fractions can be obtained, and on the other hand, more light fractions can be obtained.
  • the load of the vacuum furnace and the vacuum tower can be reduced; the extraction rate of the pressure reducing device can be increased, the yield of the distillate can be increased, and more raw materials can be provided for catalytic cracking and hydrocracking, thereby improving the economy of the refinery. benefit.
  • An important factor affecting the distillate yield of the atmospheric and vacuum distillation unit is the temperature of the vaporization section of the distillation column and the partial pressure of the oil vapor.
  • Another way to increase the distillation yield of the distillation column is to increase the temperature of the vaporization section.
  • the temperature of the vaporization section is affected by the outlet temperature of the furnace. The higher the outlet temperature of the furnace, the higher the temperature of the vaporization section.
  • the temperature of the furnace should not be too high, because the heavy oil may have a cracking reaction above 360 °C, and the coke generated by oil cracking will seriously affect the stability and long-term operation of the device. Therefore, in the industry, the furnace tube is gradually expanded in diameter and the large-diameter oil transfer line is generally used to reduce the outlet pressure of the heating furnace as much as possible, thereby lowering the temperature of the raw material of the heating furnace under the premise of ensuring the vaporization rate of the raw material.
  • the pressure of the top of the vacuum tower in the industrial plant has reached a minimum of 1 kPa (absolute pressure), and the feed section has reached 3 kPa (absolute pressure), and it has been very difficult to reduce the pressure.
  • the performance improvement of fillers and internal components is also becoming more and more difficult and the cost is greatly increased.
  • the density of the raw materials in the tube is continuously decreased, especially in the pressure reducing furnace tube, the density of the oil is decreased more, so that the heat transfer coefficient of the medium in the furnace tube is greatly reduced, thereby resulting in total furnace heat.
  • the heat transfer coefficient decreases.
  • the temperature difference must be increased, that is, the furnace and tube temperatures are increased. As a result, the wall temperature is locally too high, which may affect the service life of the tube.
  • the simulation results show that the large droplets are trapped in the furnace tube and the oil transfer line of the radiant section of the vacuum furnace.
  • the vapor phase flow rate is very fast, and the vapor-liquid two-phase interphase mass transfer area is small, so that the light fraction cannot be completely vaporized and is entrapped in the unvaporized heavy oil, resulting in low actual vaporization rate of the raw material entering the distillation section vaporization section.
  • Based on the theoretically calculated equilibrium vaporization rate a portion of the light components are present in the bottoms of the bottoms, thereby reducing the pull-out rate of the unit.
  • the domestic atmospheric and vacuum distillation unit generally cuts the vacuum residue design at 54CTC.
  • the fraction below 500 °C is more than 8w%, and the fraction below 538 °C is more than 10w%, and some even up to 30>% or more.
  • the equilibrium vaporization rate of atmospheric residue in the vaporization section of the vacuum tower is 59.0wi%, while the industrial extraction rate is only 51.9w%, indicating that the industry There is still a certain gap between the extraction rate and the equilibrium vaporization rate. It can be seen that the vacuum distillation still does not reach the equilibrium vaporization rate, and there is still much room for improvement in the extraction rate. Summary of the invention
  • the present invention provides a method for increasing the yield of a petroleum hydrocarbon distillate in a distillation column, the distillation column comprising a vaporization section and a fractionation section, the method comprising preheating the petroleum hydrocarbon feedstock oil to be fractionated, under pressure
  • the feed system enters the vaporization section of the fractionation column at a pressure higher than the distillation column vaporization section pressure of 100-100 kPa, preferably 200-800 kPa, more preferably 200-600 kPa, most preferably 200-400 kPa or 200-300 kPa, and the feedstock oil is in the vaporization section.
  • the atomization is simultaneously vaporized, and further distillation is carried out in the fractionation section of the distillation column, and the distillate product is taken out from the top and/or the side line, and the unvaporized heavy shield oil is taken out from the bottom of the column.
  • the petroleum hydrocarbon distillation column means that the heat source required for the distillation column mentioned in the foregoing is provided by the raw material, and the distillation column without the reboiler may be a flash column or a preliminary distillation column. , an atmospheric distillation column, a vacuum distillation column or a hydrogenation oil distillation column.
  • the distillation column generally comprises a vaporization section, a fractionation section, an optional overhead bottoms outlet, an optional mid-stage reflux, an optional extraction side line, an optional overhead vacuuming system, and optional stripping. Segment, optional wash section, etc.
  • the type of tower can be an empty tower, a tray tower or a packed tower.
  • the absolute pressure of the top of the distillation column is 0.5-240 kPa
  • the absolute pressure of the vaporization section is 1-280 kPa
  • the temperature of the vaporization section is 150-430 ° C
  • the top of the tower is absolutely The pressure is 1 10-180 kPa
  • the absolute pressure of the vaporization section is 130-200 kPa
  • the temperature of the vaporization section is 330-390 ° C
  • the absolute pressure of the top of the tower is 0.5-90 kPa, preferably 0.5-10 kPa or 0.5.
  • the absolute pressure of the vaporization section is l-98kPa
  • the vaporization section temperature is 300-430 ° C, preferably 370-410 ° C.
  • the vaporization section of the steaming tower described herein is between the upper part of the crucible section and the feed port, and the raw material introduced into the steaming tower through the feed port is fully vaporized in the vaporization section, and may be completely vaporized or partially vaporized, after vaporization.
  • the gas phase enters the upper partial fraction for heat exchange and further fractionation.
  • the temperature and pressure of the vaporization section are distributed in a gradient, the vaporization section temperature is a temperature range of the vaporization section, and the vaporization section absolute pressure is a pressure range of the vaporization section.
  • preheating is carried out by means of a heating furnace such as an atmospheric pressure furnace and a vacuum furnace.
  • the furnace outlet pressure is preferably 100-1000 kPa higher than the vaporization section pressure, preferably 200-800 kPa higher, more preferably 200-600 kPa higher, most preferably 200-400 kPa or 200-300 kPa higher, and the outlet temperature of the heating furnace is 360-460°. C, preferably 380-430 °C.
  • the furnace tube in the case of using a heating furnace, may be injected with or without steam, and it is preferable that steam is not injected.
  • the vacuum distillation column may or may not be injected with steam, and it is preferred that steam is not injected.
  • the pressure feed system includes a flow distribution system and one or more atomization devices.
  • the atomizing device may be in the distillation section of the distillation column, or outside the distillation column, or both.
  • the flow distribution system can ensure that each atomizing device can be ejected by liquid and vapor in any case, thereby ensuring the atomization effect of the raw material.
  • the flow distribution system may be a piping system consisting of in-line, misplaced, parallel, vertical, toroidal, tree-shaped, symmetrical and asymmetrical pipelines, the purpose of which is to preheat The subsequent material is distributed to each atomizing device, and the pipe arrangement selected for this purpose can be regarded as a flow distribution system.
  • the atomization device may be one or more nozzles extending into the vaporization section of the distillation column or other equipment capable of atomizing heavy oil, and/or extending into the distillation tower and One or more nozzles of the atomizing vessel in communication with the distillation column or other apparatus that can atomize the heavy oil.
  • the atomizing device such as a nozzle (including but not limited to the following nozzle forms, such as a swirling atomizing nozzle, a centrifugal atomizing nozzle, a variable area pressure atomizing nozzle, etc.) may be single hole or porous, open
  • the direction of the holes may be arbitrary, either with or without auxiliary steam, and the auxiliary atomizing steam may enter with the feedstock oil or may enter separately.
  • the size of the droplet after atomization can be sufficient to ensure a good vaporization effect, and the purpose of effectively fractionating the oil is achieved.
  • the flow distribution system can be placed outside the tower or placed in the tower.
  • the flow distribution system can be placed outside of the atomizing container or in the atomizing container.
  • the distribution system can be a flow distribution system with automatic control or a fully self-regulating flow distributor without automatic control.
  • the flow distribution system with automatic control consists mainly of piping and automatically controlled valves.
  • the flow distribution system without automatic control distributes the logistics to each atomizing device mainly by rationally designing the resistance of each branch pipe.
  • the atomization container is a container having sufficient space to atomize heavy oil.
  • the atomizing container is a shift line, a flash tank or a flash tower.
  • the use of the oil transfer line as the atomization container can achieve the advantage. If the flash tank is set as an atomization container, although the equipment investment will be increased, the flash tank not only provides more space and time for atomization and vaporization of the feedstock oil, but also facilitates vaporization and non-vaporization after vaporization. The mist drops are separated.
  • the atomizing device comprises one or more nozzles or other atomic containers that extend into the atomization vessel outside the distillation column and in communication with the distillation column.
  • the vapor phase stream formed in the atomization vessel enters the vaporization section of the distillation column, and the formed liquid phase stream directly enters the bottom of the fractionation column and is mixed with the residue of the bottom of the column, or is formed in the atomization container.
  • the vapor phase stream and the formed liquid phase stream enter the distillation column vaporization section from the same line.
  • the atomization device comprises one or more nozzles extending into the vaporization section of the distillation column or other equipment capable of atomizing heavy oil, and the petroleum hydrocarbon feedstock oil to be fractionated is preheated. After that, the pressure feed system is higher than the distillation section vaporization section pressure
  • a foaming element 9 may be disposed above the vaporization section, and/or a liquid collecting element 10 may be disposed below the vaporizing section.
  • the foaming element 9 is a foaming net or a vapor-liquid filter, which functions to reduce or eliminate entrainment of the mist and prevent the liquid from being carried into the fractionation section by the vapor phase.
  • the liquid collecting member 10 is a layer or a plurality of liquid collecting trays for collecting droplets which continuously gather during the collision of the droplets to fall into the bottom of the tower as the residual oil is taken out.
  • the foaming element 9 and the liquid collecting element 10 are both provided to increase the fractionation efficiency of the distillation column.
  • the present invention provides a steam for increasing the yield of petroleum hydrocarbon distillates
  • the distillation column comprises a vaporization section, characterized in that the distillation column comprises a pressure type for feeding the petroleum hydrocarbon feedstock oil to be fractionated at a pressure higher than a pressure of the distillation section vaporization section of 100-100 kPa. Material system.
  • the distillation column is a distillation column which is not provided with a reboiler, and preferably includes a flash column, a preliminary distillation column, an atmospheric distillation column, a vacuum distillation column or a hydrogenation oil distillation column.
  • a liquid collecting member is disposed under the inlet of the raw material oil, and/or a foaming member is disposed above the inlet of the raw material oil.
  • the pressure feed system comprises a flow distribution system and an atomization device, which may be in the distillation section of the distillation column, or outside the distillation column, or both.
  • the atomizing device is one or more nozzles or other means for atomizing the heavy oil which protrude into the vaporization section of the distillation column, and/or extends into the outside of the steaming tower and One or more nozzles of the atomizing vessel in communication with the distillation column or other apparatus that can atomize the heavy oil.
  • the flow distribution system is placed in the column and/or outside the atomizing vessel and/or in the atomizing vessel.
  • the atomization container is a shift line, a flash tank or a flash column.
  • Benefits that can be provided by the present invention include:
  • the feedstock oil to be fractionated is introduced into the distillation column through a pressure feed system under a certain pressure after preheating, and the atomization of the atomization device accelerates the vaporization of the feedstock oil in the vaporization section, so that the raw material oil is actually in the vaporization section.
  • the vaporization rate is closer to the equilibrium vaporization rate, so as to maximize the vaporization of the light distillate in the feedstock oil into the gas phase.
  • the vaporization rate is greatly increased due to the sharp increase in surface area, thereby facilitating the improvement of the fraction. Oil yield.
  • the pressure in the furnace tube is increased, the density of the oil in the furnace tube is increased, the heat transfer coefficient is increased, and the total heat transfer coefficient is increased accordingly.
  • the furnace temperature can be lowered.
  • the surface temperature of the tube and the degree of thermal cracking of the feedstock can be reduced.
  • the method or device provided by the invention is used for vacuum distillation, the extraction rate of the vacuum tower can be increased, and the diameter of the oil transfer line can be greatly reduced; for atmospheric distillation, the drawing of the atmospheric pressure tower can be improved.
  • Output rate reduce the load of the vacuum furnace and the vacuum tower; at the same time, it is used for the atmospheric tower and the vacuum tower, Increasing the total extraction rate of the atmospheric and vacuum distillation unit while reducing energy consumption and operating costs.
  • Figure 1 is a schematic flow chart of conventional atmospheric distillation
  • FIG. 2 is a schematic flow chart of a method provided by the present invention for atmospheric distillation
  • Figure 3 is a schematic flow chart of conventional vacuum distillation
  • FIG. 4 is a schematic flow chart of a method provided by the present invention for vacuum distillation
  • Figure 5 is a schematic flow chart of the atomization container being the oil transfer line
  • FIG. 6 is a schematic flow chart of the atomization container being a flash tank and the gas-liquid mixed phase feeding
  • FIG. 7 is a schematic flow chart of the atomization container being a flash tank and the gas-liquid two phases are respectively fed.
  • FIG. 4 is a process provided by the present invention for use in a vacuum distillation process.
  • the vacuum distillation column is divided into a vaporization section 1 1 , a washing section 12 and a fractionation section 13 , and the raw material oil (normal pressure residue) to be branched is driven into the heating furnace 2 through the feed pump 1 Heat, furnace 2 furnace outlet pressure is higher than the distillation column vaporization section 100- l OOOkPa, preferably high 200-800kPa, more preferably high 200-600kPa, most preferably high 200-400kPa or 200-300kPa, heating furnace tube outlet temperature is 360 -460 ° C, preferably 380-430 ° C.
  • the preheated feedstock oil is introduced into the bottom of the distillation column by the pressure feed system 3, and the pressure feed system includes a flow distribution system 4 and an atomization device 5, and the preheated feedstock oil is fixed by the flow distribution system 4 After the proportion is distributed, it is atomized into small droplets by the atomizing device 5, sprayed into the vaporization section of the vacuum distillation column, and rapidly vaporized. Since the droplets have a large specific surface area, during the droplet movement of the vaporization section, The vaporized fraction is fully vaporized in a very short time.
  • a foaming element 9 is arranged above the atomizing device 5, and a liquid collecting element 10 is arranged below the atomizing device 5.
  • the vaporized fraction in the vaporization section 11 is introduced upward into the washing section 12 and the fractionation section 13 of the vacuum distillation column, and after fractionation, is taken out from the top or side line to obtain a distillate product.
  • the structure of the washing section 12 and the fractionating section 13 is the same as that of the conventional vacuum tower.
  • the heavy fraction which is difficult to vaporize maintains a liquid phase state, and the droplets continuously aggregate to form large droplets during collision with each other, and are collected by the liquid collecting member 10 to fall to the bottom of the column, and are taken out as residual oil.
  • Fig. 5 an embodiment of the present invention will be described by taking a vacuum distillation as an example, in which the atomization container is a shift line.
  • the raw material oil to be fractionated (such as atmospheric residue) is preheated by the feed pump 1 into the heating furnace 2, and the pressure in the furnace tube of the heating furnace 2 is 100-1000 kPa higher than the vaporization section, preferably 200-800 kPa higher, more preferably 200 high. -600 kPa, most preferably 200-400 kPa or 200-300 kPa, and the outlet temperature of the heating furnace tube is 360-460 ° C, preferably 380-430 ° C.
  • the preheated feedstock is injected into the transfer line 7 from the pressure feed system 3, the pressure in the transfer line 7 is 2.0-60.0 kPa, and the temperature is 230-460° (:.
  • the mist drops at low oil vapor Fully vaporized under pressure, the vaporized vapor stream is introduced into the vaporization section 8 of the reduced pressure fractionation column 6. This embodiment allows the droplets to be fully vaporized, thereby increasing the extraction rate of the reduced pressure fractionation column.
  • a vacuum distillation in which the atomization container is a flash tank and is different from the atomization container of Fig. 5 in that the oil transfer line is preheated.
  • the subsequent feedstock is sprayed into the flash tank 9 by the pressure feed system 3, and the pressure in the flash tank 9 is 2.0-60.0 kPa and the temperature is 230-460 °C. Since the droplets have a large specific surface area, the fraction having a lower boiling point is flash vaporized under the condition of low oil vapor partial pressure in the flash tank.
  • the fully vaporized vapor stream is introduced into the vaporization section 8 of the reduced pressure fractionation column 6, which allows the droplets to be fully vaporized, thereby increasing the extraction rate of the reduced pressure fractionation column.
  • Comparative Example 1 illustrates the effect of fractional distillation of crude oil by the atmospheric distillation method of the prior art.
  • the properties of the mixed crude oil to be fractionated are shown in Table 1.
  • 1 is a schematic flow chart of an atmospheric pressure fractionation method in the prior art. As shown in FIG. 1, the mixed crude oil is first heated by an atmospheric pressure heating furnace 2, and the outlet temperature of the heating furnace is 368 ° C, and is subjected to atmospheric distillation through the oil transfer line 7.
  • the atmospheric distillation tower is a plate tower, having a diameter of 6.5 meters, having three side lines and two middle sections flowing back to obtain straight-run gasoline.
  • the kerosene, diesel and other fractions, atmospheric distillation tower operating conditions and product properties are shown in Table 2.
  • the extraction rate of the atmospheric distillation column was 30.2%.
  • Example 1 illustrates the effect of the method provided by the present invention on atmospheric distillation of crude oil.
  • the atmospheric distillation column 8 used is the same as that of the comparative example 1, and the raw material oil to be fractionated is the same as that of the comparative example 1, the raw material oil
  • the pressure feed system (including the flow distribution system 4 and the atomization device 5) is injected into the atmospheric distillation column 8 under the condition of a pressure of 500 kPa higher than the vaporization section of the distillation column, and the atmospheric distillation column is An atomizing device is installed, the atomizing device is a swirling atomizing nozzle, the swirling core is placed at the front of the nozzle, and a single-hole plate is installed at the top of the swirling core, and the swirling liquid is ejected through the hole to form a taper.
  • Example 1 Residual pressure at the top of the tower, kPa (absolute) 170.0 170.0
  • Atmospheric pressure furnace outlet temperature °c 368.0 372.0
  • Vaporization section temperature °C 365.5 364.8
  • Atmospheric pressure second line extraction temperature 253.4 255.9
  • Atmospheric pressure extraction rate 30.2 33.2 It can be seen from Table 2 that the method provided by the present invention is used for atmospheric distillation, and the atmospheric pressure heating furnace outlet pressure is increased by 166.4 kPa compared with the conventional pressure atmospheric pressure distillation method. The outlet temperature is increased by 4.0 °C. In the case where the temperature and pressure of the distillation section of the distillation column are substantially the same, the extraction rate of the distillation column reaches 33.2%, which is 3% higher than that of the conventional feed.
  • the method provided by the present invention is applied to an atmospheric pressure steaming tower, which can increase the extraction rate of the atmospheric pressure tower. Comparative example 2
  • Example 2 illustrates the effect of the prior art vacuum fractionation atmospheric residue.
  • FIG. 3 is a decompression in the prior art Schematic diagram of the distillation method, as shown in Fig. 3, the atmospheric pressure bottom oil is heated by the vacuum furnace 2, the outlet pressure of the vacuum furnace is 30.0 kPa (absolute), and the surface temperature of the vacuum furnace tube is 593 ° C, the vacuum furnace The outlet temperature was 410 ° C, and the preheated feedstock oil was introduced into the vacuum distillation column 6 via the oil transfer line 7.
  • the furnace tube of the vacuum furnace is continuously expanded from ⁇ 152mm to ⁇ D273mm, the diameter of the oil transfer line is 2.0m, and the length is 33.0m.
  • the feed is subjected to gas-liquid separation through a feed distributor in the distillation column.
  • the vacuum distillation column is a conventional fully packed column, 9.2 m in diameter, dry operation.
  • the vacuum distillation column is divided into a vaporization section, a washing section and a fractionation section, and the vaporization section temperature is 393.7 °C.
  • the washing section is filled with ZUPAC2 series packing (Tianjin University Beiyang Chemical Equipment Co., Ltd.) 1.5 meters
  • the separation section is filled with two layers of ZUPAC 1 packing (Tianjin University Beiyang Chemical Equipment Co., Ltd.).
  • the decompression tower includes four discharge ports from top to bottom for the top reduction, the minus one line, the minus two lines, the minus three lines, and the two middle sections.
  • the top vacuum system uses a three-stage vacuum.
  • the operating conditions and product properties of the vacuum distillation column are shown in Table 4.
  • the extraction rate of the vacuum distillation column was 57.6 %.
  • Example 2 illustrates the effect of the method provided by the present invention on a vacuum distillation column.
  • the raw material oil to be classified was an atmospheric residue, which was the same as Comparative Example 2.
  • the raw material oil is heated by the vacuum furnace 2, the diameter of the furnace tube is (D152mm, the heated feedstock oil enters the oil transfer line, and then passes through the pressure feed system (including the flow distribution system) at a pressure higher than the distillation section vapor pressure of 300 kPa. 4 and the atomization device 5) is sprayed into the vacuum distillation column 6, and the atomization device is installed in the vacuum distillation column.
  • the atomization device is as described in Example 1.
  • the operating conditions and product properties of the vacuum distillation column are shown in Table 4.
  • the method provided by the present invention is used in vacuum distillation, compared with the vacuum distillation method of the conventional feed of Comparative Example 2, at the same vaporization stage temperature and Under pressure, the extraction rate of the vacuum distillation column reached 60.2%, which was 2.6% higher than that of the conventional feed.
  • the outlet temperature of the vacuum furnace is increased by 18 °C, the surface temperature of the furnace tube is reduced by 33 °C, and the non-condensing amount of the top of the vacuum tower is reduced from 0.3% to 0.2%.
  • the pressure reducing furnace tube is stepwise The diameter expansion is more complicated, and the diameter and the oil transfer line diameter of the furnace tube are both (D 152 mm, the structure of the tube and the oil transfer line are compressed; in addition, compared with the comparative example 2 , the final boiling point of vacuum wax oil increased
  • Comparative Example 3 illustrates the effect of the fractional distillation of the atmospheric pressure fractionation column in the prior art by the reduced pressure fractionation process. '
  • the mixed crude oil to be fractionated is introduced into an atmospheric pressure fractionation column, and fractionated to obtain a straight-run gasoline, a kerosene, and a diesel fraction, and the atmospheric pressure tower extraction rate is 32 ⁇ ⁇ %.
  • the atmospheric pressure fractionator bottoms oil is sent to the vacuum distillation system heating furnace 2 through the oil pump 1, and after heating, it is introduced into the vacuum distillation column vaporization section 8 through the oil transfer line 7.
  • the outlet pressure of the furnace tube is 30.0 kPa, the wall temperature is 561 ° C, the furnace outlet temperature is 386 ° C, and the furnace tube is expanded step by step.
  • the vacuum distillation column is a high efficiency fully packed column, and the temperature of the vaporization section 8 of the vacuum distillation column is 374 °C.
  • the properties of the mixed crude oil are shown in Table 5.
  • the operating conditions and product properties of the vacuum distillation column are shown in Table 6.
  • the extraction rate of the vacuum distillation column was 29.8w%.
  • Example 3 illustrates the effect of the process provided by the present invention on the vacuum distillation of crude oil.
  • the atmospheric column system used and the mixed crude oil to be fractionated were the same as in Comparative Example 3, and the atmospheric pressure column extraction rate was 32 wi%.
  • the bottom oil obtained by fractionating the atmospheric distillation column is first sent to the vacuum distillation system heating furnace 2 through the oil pump 1, and the heated atmospheric base oil is sprayed into the oil transfer line 7 through the nozzle 5, often The bottom oil of the pressure tower is fully vaporized in the oil transfer line, and then introduced into the vaporization section 8 of the vacuum tower through the oil transfer line.
  • the pressure at the inlet of the transfer line is 14.0 kPa and the temperature is 386 °C.
  • the nozzle used was a centrifugal atomizing nozzle; the furnace tube of the heating furnace was not tapered.
  • the structure of the oil transfer line and the vacuum column used was the same as that of Comparative Example 3, and the temperature of the vaporization section of the vacuum column was 381 °C.
  • Example 4 In the case where the pressure of the vaporization section was the same as that of Comparative Example 3, the extraction rate of the raw material in Example 3 after passing through the vacuum distillation system reached 33.7 wt%, which was higher than that of Comparative Example 3 by 3.9 percentage points. Vacuum residue density and viscosity increase, vacuum residue small and medium The mass content of the fraction at 500 °C was also reduced from 10% of Comparative Example 3 to 5.8%.
  • Example 4
  • Example 4 illustrates the effect of the process provided by the present invention on the vacuum distillation of crude oil.
  • the atmospheric column system used and the mixed crude oil to be separated were the same as in Comparative Example 3, and the atmospheric pressure column extraction rate was 32 wi%.
  • the structure of the vacuum tower used was the same as that of Comparative Example 3, and the structure of the furnace used was the same as that of Example 3.
  • the difference is that a flash tank 9 is added after the vacuum furnace, and the atmospheric pressure bottom oil is distributed by the flow distribution system 4, and then sprayed into the flash tank through the nozzle 5, and after being fully vaporized, the vacuum distillation tower is introduced.
  • the flash tank pressure is 6.1 kPa and the temperature is 382 °C.
  • the other main operating conditions and product properties are shown in Table 6.
  • Example 4 As can be seen from the data in Table 6, Example 4, by providing an atomizing nozzle and a flash tank at the outlet of the heating furnace, the extraction rate of the vacuum distillate in the atmospheric bottom oil was 34.5 wi%, and Comparative Example 3 Compared with an increase of 4.7 percentage points.
  • Example 5 As can be seen from the data in Table 6, Example 4, by providing an atomizing nozzle and a flash tank at the outlet of the heating furnace, the extraction rate of the vacuum distillate in the atmospheric bottom oil was 34.5 wi%, and Comparative Example 3 Compared with an increase of 4.7 percentage points.
  • Example 5 As can be seen from the data in Table 6, Example 4, by providing an atomizing nozzle and a flash tank at the outlet of the heating furnace, the extraction rate of the vacuum distillate in the atmospheric bottom oil was 34.5 wi%, and Comparative Example 3 Compared with an increase of 4.7 percentage points.
  • Example 5 As can be seen from the data in Table 6, Example 4, by providing an atomizing nozzle and a flash tank at the outlet of the heating furnace, the extraction rate
  • Example 5 illustrates the effect of the method provided by the present invention on the vacuum distillation of crude oil.
  • the atmospheric column system used and the mixed crude oil to be fractionated were the same as in Comparative Example 3, and the atmospheric pressure column extraction rate was 32 wi%.
  • the structure of the vacuum tower used in Example 5 was the same as that in Example 4.
  • the structure of the furnace used was the same as in Example 4, and the flash tank used was the same as in Example 4.
  • the atmospheric pressure bottom oil is discharged into the flash tank 9 through the nozzle 5 after being distributed by the flow distribution system 3, and after being fully vaporized, the gas and the liquid are separately introduced into the vacuum distillation from different pipelines.
  • the main operating conditions and product properties of this example are shown in Table 6.

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Description

提高石油烃类馏分油收率的蒸馏塔和其进料方法 技术领域
本发明涉及一种提高蒸馏装置馏分油收率的蒸馏塔和方法，更具体 地说，涉及一种提高石油炼制工业中的重质油蒸馏过程馏分油收率的蒸 馏塔和方法。 背景技术
蒸馏塔是石油炼制工业中应用十分广泛的单元设备。对于一些重质 油品的分馏， 如从原油、 蜡油等油品分馏轻质馏分油时， 蒸馏塔塔釜温 度都较高， 再沸器的热源温位高， 不易获得， 而且重质油品在高温下易 发生热裂解， 所以原油或重油蒸馏塔一般都不设再沸器， 蒸馏所需的热 量来源几乎完全由原料提供， 原料经预热后馏分油汽化， 汽化后的馏分 油从塔顶和 /或从侧线馏出， 未汽化的部分从塔釜馏出。 典型的分馏过 程如原油的常压蒸馏和减压蒸馏。
原油常减压蒸馏是原油加工的第一道工序，它为炼厂后续加工装置 提供原料， 并直接提供部分产品。 原油蒸馏（以燃料油型为例）的基本 流程为， 原油被加热到 220-260°C左右进初馏塔， 通常初馏塔只取一个 塔顶产品， 即重整料或轻汽油镏分。 也有的初馏塔除塔顶产品外， 还有 一个侧线产品， 初镏塔塔底油送常压塔。
常压塔常规流程见图 1， 初馏塔的塔底油经换热或常压加热炉 2加 热部分汽化后， 经转油线 7进入常压蒸馏塔 8， 在蒸馏塔汽化段轻组分 汽化并上升进入分馏段，经过回流液体的冷凝从塔顶或侧线抽出得到熘 分油， 未汽化的部分向下流入提镏段， 在提馏段的塔板上与塔底进入的 水蒸汽接触，其中未汽化的轻馏分被汽提出来并随水蒸汽一起向上进入 分馏段。 得到汽油、 煤油、 柴油、 重柴油等较轻的组分。 未汽化的部分 落入塔底， 作为常压渣油引出。
减压蒸馏过程的常规流程见图 3 , 常压渣油经减压加热炉 2加热部 分汽化后， 经转油线 7进入减压蒸馏塔 6， 在减压蒸馏塔汽化段轻组分 汽化并上升进入分馏段，经过回流液体的冷凝后从塔顶或侧线抽出得到 馏分油， 未汽化的部分从塔底引出， 得到减压渣油。 原油蒸馏装置设计和操作的优劣， 会对炼油厂的产品质量、 产品收 率和经济效益产生很大影响。 在保证产品质量的前提下， 提高常压蒸馏 装置的拔出率， 可以使轻组分尽量在常压塔拔出， 不会再进入减压塔， 一方面可以得到更多的轻馏分，另一方面可以减少减压炉和减压塔的负 荷； 提高减压装置的拔出率， 可以增加馏分油的收率， 为催化裂化、 加 氢裂化提供更多的原料， 从而提高炼厂的经济效益。
影响常减压装置馏分油收率的重要因素是蒸馏塔汽化段的温度和 油汽分压。 汽化段温度越高， 油汽分压越低则原料的汽化率越高， 馏分 油的拔出率也就越高。
目前， 工业上降低汽化段压力的方法主要有两种： 一是降低蒸馏塔 塔顶压力， 对于常压蒸馏塔， 一般是减少塔顶油汽管线和冷凝冷却器的 压降。 对于减压蒸馏塔而言， 高性能的抽真空设备可以有效降低塔顶压 力。 二是采用高性能的填料、 塔板和塔内件， 有效降低塔内阻力， 从而 使汽化段压力显著降低。
提高蒸馏塔镏分油收率的另一途径是提高汽化段温度。汽化段的温 度受加热炉出口温度的影响，加热炉出口温度越高，汽化段温度就越高。 但是加热炉温度又不能太高， 因为重油在 360 °C以上有发生裂解反应的 可能， 油品裂解生成的焦炭会严重影响装置的稳定性和长周期运行。 因 此， 工业上一般采用加热炉炉管逐级扩径和大直径转油线， 尽可能降低 加热炉出口压力， 从而在保证原料汽化率的前提下降低加热炉原料温 度。
目前， 工业装置中减压塔顶压力最低已经达到 1 kPa (绝压） ， 进 料段已经达到 3 kPa (绝压） ， 再降低压力已经非常困难。 填料和内构 件的性能提高也越来越困难， 成本大幅增加。 采用炉管逐级扩径和大直 径转油线也存在一定的限制，一是炉管的扩径必须根据原料油的性质和 加热炉的特性进行合理设计， 而原料种类繁多， 使得精确的炉管设计非 常困难。 二是随着原料在炉管内大量汽化， 管内原料的密度不断下降， 特别是在减压炉管内， 油品密度下降更大， 使得炉管内介质的给热系数 大大减小， 从而导致炉内总传热系数下降， 为达到相同的传热强度必须 提高温差， 亦即提高炉膛和炉管温度， 其结果会导致出现管壁温度局部 过高， 易影响炉管使用寿命。
模拟计算结果表明， 在减压炉辐射段炉管和转油线内， 夹杂大液滴 的汽相流速很快， 而且汽液两相相际传质面积较小， 使得轻馏分不能完 全汽化而被包裹在未汽化的重质油中，导致进入蒸馏塔汽化段原料的实 际汽化率低于理论计算的平衡汽化率，一部分轻质组分存在于塔底渣油 中， 从而降低装置的拔出率。 目前国内常减压装置一般将减压渣油设计 切割点在 54CTC，许多减压渣油中低于 500 °C馏分含量大于 8w % , 低于 538 °C馏分含量大于 10w %, 有的甚至高达 30> %以上。 以中石化海南 炼化公司炼厂常减压装置为例，常压渣油在减压塔汽化段温度和压力下 平衡汽化率为 59.0wi%， 而工业拔出率仅为 51.9w %, 说明工业拔出率 与平衡汽化率仍有一定的差距。 由此可见， 减压蒸馏仍未达到平衡汽化 率， 拔出率还有很大的提升空间。 发明内容
本发明目的是提供一种提高石油烃蒸馏塔馏分油收率、尤其是提高 常减压蒸馏塔馏分油收率的蒸馏塔和方法。
本发明提供了一种在蒸馏塔中提高石油烃类馏分油收率的方法，所 述的蒸馏塔包括汽化段和分馏段，所述方法包括将待分馏的石油烃原料 油预热， 经过压力式进料***在高于蒸馏塔汽化段压力 100- lOOOkPa, 优选 200-800kPa，更优选 200-600kPa，最优选 200-400kPa或 200-300kPa 的条件下进入分馏塔汽化段， 原料油在汽化段雾化同时汽化， 进而在蒸 馏塔的分馏段进行蒸馏分离， 塔顶和 /或侧线引出馏分油产品， 塔底引 出未汽化的重盾油。
本发明提供的方法中，所述的石油烃蒸馏塔是指前文中提到的蒸馏 塔所需的热量来源由原料提供， 不设再沸器的蒸馏塔， 可以是闪蒸塔、 初馏塔、 常压蒸馏塔、 减压蒸馏塔或加氢生成油蒸馏塔。 所述的蒸馏塔 一般包括汽化段、 分馏段、 任选的塔顶塔底出料口、 任选的中段回流、 任选的抽出侧线、 任选的塔顶抽真空***、 任选的提馏段、 任选的洗涤 段等。 塔的类型可以是空塔、 板式塔或者填料塔。
本发明提供的方法中， 蒸馏塔塔顶绝对压力为 0.5-240kPa 、 汽化 段绝对压力为 l-280kPa、 汽化段温度为 150-430°C ; 具体地， 在常压蒸 馏情况下，塔顶绝对压力为 1 10-180kPa、汽化段绝对压力为 130-200kPa， 汽化段温度为 330-390°C ; 和在减压蒸馏情况下， 塔顶绝对压力为 0.5-90kPa, 优选地 0.5-10kPa或 0.5-3kPa， 汽化段绝对压力为 l-98kPa, 优选地 l -5kPa， 汽化段温度为 300-430°C， 优选地 370-410°C。
本文中所述的蒸镏塔汽化段介于上部分熘段和进料口之间，经进料 口引入蒸镏塔的原料在汽化段充分汽化， 可以是全部汽化或者部分汽 化， 汽化后的气相向上进入上部分馏段进行热交换并进一步分馏。 所述 的汽化段的温度和压力成梯度分布，所述的汽化段温度为汽化段的温度 范围， 所述的汽化段绝对压力为汽化段的压力范围。
本发明提供的方法中， 预热是通过加热炉 （例如常压炉和减压炉） 进行的。 所述的加热炉出口压力优选比汽化段压力高 100- 1000kPa， 优 选高 200-800kPa， 更优选高 200-600kPa， 最优选高 200-400kPa 或 200-300kPa, 加热炉出口温度为 360-460°C、 优选 380-430°C。
本发明提供的方法中，在使用加热炉的情况下所述的加热炉炉管可 以注入蒸汽， 也可以不注入蒸汽， 优选的方案是不注入蒸汽。
本发明提供的方法中， 在减压蒸馏的情况下， 所述的减压蒸馏塔可 以注入蒸汽， 也可以不注入蒸汽， 优选的方案是不注入蒸汽。
本发明提供的方法中，所述的压力式进料***包括流量分配***和 一个或多个雾化 i殳备。 所述的雾化设备可以在蒸馏塔汽化段内， 或者在 蒸馏塔外， 或者两者兼之。
本发明提供的方法中，所述的流量分配***能保证每个雾化设备在 任何情况下都能有液体和汽体喷出， 从而保证原料的雾化效果。 所述的 流量分配***可以是由直列的、 错置的、 平行的、 竖直的、 环型的、 树 型的、对称的和不对称的管路组成的管系， 其目的就是把预热后的原料 分配到每一个雾化设备，为此目的选用的管道排列方式均可视为流量分 配***。
本发明提供的方法中，所述的雾化设备可以是伸入蒸馏塔汽化段内 的一个或多个喷嘴或其它可以使重油雾化的设备， 和 /或伸入位于蒸馏 塔塔外的且与蒸馏塔连通的雾化容器的一个或多个喷嘴或其它可以使 重油雾化的设备。所述雾化设备如喷嘴（包括但不限于以下的喷嘴形式， 例如旋流式雾化喷嘴、 离心式雾化喷嘴、 变面积压力式雾化喷嘴等） ， 可以是单孔或多孔的， 开孔方向可以是任意的， 可以是带有辅助雾化蒸 汽或不带辅助蒸汽的，辅助雾化蒸汽可以与原料油一起进入也可分别进 入。 雾化后的雾滴尺寸可足以保证良好的汽化效果， 达到将油品有效分 馏的目的。 流量分配***可以放置在塔外， 也可以放置在塔内。 流量分配*** 可以放置在雾化容器外， 也可以放置在雾化容器内。 分配***可以是带 自动控制的流量分配***，也可以是不带自动控制的完全自我调节流量 的分布器。带自动控制的流量分配***主要有管路和自动控制的阀门组 成。 不带自动控制的流量分配***主要通过合理设计各分支管路的阻 力， 将物流分配到各雾化设备中。
本发明提供的方法中，所述的雾化容器为有足够空间可以使重油雾 化的容器。 例如雾化容器为转油线、 闪蒸罐或闪蒸塔。 对于现有装置的 改造而言， 采用转油线作为雾化容器可以实现利旧。 若设置闪蒸罐为雾 化容器， 虽然会增加设备投资，但闪蒸罐不仅可以为原料油的雾化和汽 化提供更足够的空间和时间，而且更有利于汽化后的油汽与未汽化的雾 滴分离。
本发明提供的方法中，一种方案是所述的雾化设备包括伸入位于蒸 馏塔塔外的且与蒸镏塔连通的雾化容器的一个或多个喷嘴或其它可以 使重油雾化的设备；在所述的雾化容器中形成的汽相物流进入蒸馏塔汽 化段， 形成的液相物流直接进入分馏塔的底部与塔底的渣油混合， 或者 在所述的雾化容器中形成的汽相物流和形成的液相物流从同一个管线 进入蒸馏塔汽化段。
本发明提供的方法中，一种方案是所述的雾化设备包括伸入蒸馏塔 汽化段内的一个或多个喷嘴或其它可以使重油雾化的设备，待分馏的石 油烃原料油预热后， 经过压力式进料***在高于蒸馏塔汽化段压力
100- lOOOkPa,优选 200-800kPa,更优选 200-600kPa，最优选 200-400kPa 或 200-300kPa的条件下雾化， 同时全部或部分汽化， 进入蒸馏塔汽化 段， 塔顶和 /或侧线引出馏分油产品， 塔底引出重质油。
本发明提供的方法中， 所述的蒸镏塔内， 在汽化段上方可以设置破 沫元件 9， 和 /或汽化段的下方可以设置液体收集元件 10。 所述的破沫 元件 9为破沫网或汽液过滤网， 其作用是减少或消除雾沫夹带， 避免液 体被汽相带入分馏段。 所述的液体收集元件 10为一层或多层集液盘， 用来收集雾滴相互碰撞过程中不断聚集形成大的液滴，使之落入塔底作 为渣油被引出。 设置破沫元件 9和液体收集元件 10 , 均可以提高蒸馏 塔的分馏效率。
在另一方面，本发明提供了一种用于提高石油烃类馏分油收率的蒸 熘塔， 所述的蒸馏塔含有汽化段， 其特征在于所述的蒸馏塔包括用于使 待分馏的石油烃原料油以高于蒸馏塔汽化段压力 100-lOOOkPa 的压力 进料的压力式进料***。
根据本发明的蒸馏塔， 所述的蒸馏塔为不设置再沸器的蒸馏塔， 优 选包括闪蒸塔、初馏塔、常压蒸馏塔、减压蒸馏塔或加氢生成油蒸馏塔。
根据 发明的蒸馏塔， 所述的蒸馏塔内， 喷入原料油的入口下方设 置液体收集元件， 和 /或喷入原料油的入口上方设置破沫元件。
根据本发明的蒸馏塔，所述的压力式进料***包括流量分配***和 雾化设备， 所述的雾化设备可以在蒸馏塔汽化段内， 或者在蒸馏塔外， 或者两者兼之。
根据本发明的蒸馏塔，所述的雾化设备是伸入蒸馏塔汽化段内的一 个或多个喷嘴或其它可以使重油雾化的设备， 和 /或伸入位于蒸熘塔塔 外的且与蒸馏塔连通的雾化容器的一个或多个喷嘴或其它可以使重油 雾化的设备。
根据本发明的蒸熘塔， 所述的流量分配***被放置在塔内和 /或在 雾化容器外和 /或在雾化容器内。
根据本发明的蒸馏塔，所述的雾化容器为转油线、闪蒸罐或闪蒸塔。 通过本发明能够提供的益处包括：
首先，待分馏的原料油经预热后在一定压力下经压力式进料***引 入蒸馏塔， 经雾化设备的雾化作用加速原料油在汽化段的汽化，使原料 油在汽化段的实际汽化率更接***衡汽化率，从而最大程度地使原料油 中的轻质馏分油汽化到气相中， 同时雾化成雾滴后， 由于表面积急剧增 加， 汽化速率也会大幅提高， 从而有利于提高馏分油收率。
其次， 加热炉炉管内压力提高， 炉管内油品的密度增加， 给热系数 增大， 总传热系数相应增大， 在相同传热强度或相同的炉出口温度下， 加热炉炉膛温度可以降低，从而可以降低炉管表面温度和原料油的热裂 化程度。
第三， 由于炉管内压力高， 油品基本不汽化， 因此炉管不需要多次 扩径， 从而使得加热炉结构简化同时转油线直径也可大幅减小。
第四， 本发明提供的方法或装置用于减压蒸馏， 可以提高减压塔的 拔出率， 转油线直径也可大幅减小； 用于常压蒸馏， 可以提高常压塔的 拔.出率， 降低减压炉和减压塔的负荷； 同时用于常压塔和减压塔， 可以 提高常减压蒸馏装置的总拔出率同时又降低了能耗和操作费用。 附图说明
图 1为常规常压蒸馏的流程示意图；
图 2为本发明提供的方法用于常压蒸馏的流程示意图；
图 3为常规减压蒸馏的流程示意图；
图 4为本发明提供的方法用于减压蒸馏的流程示意图；
图 5为雾化容器为转油线的流程示意图；
图 6为雾化容器为闪蒸罐且气液混相进料的流程示意图； 图 7为雾化容器为闪蒸罐且气液两相分别进料的流程示意图。
具体实施方式
下面结合附图具体说明本发明提供的提高石油烃馏分油收率的方 法和相关装置， 但本发明并不因此而受到限制。
以下参考图 4， 以减压蒸熘为例， 说明本发明的一种实施方案。 图 4为本发明提供的方法用于减压蒸馏过程。如图 4所示， 减压蒸 熘塔分为汽化段 1 1、 洗涤段 12和分馏段 13， 待分镏的原料油（常压渣 油）经进料泵 1打入加热炉 2中预热， 加热炉 2炉出口压力比蒸馏塔汽 化段高 100- l OOOkPa, 优选高 200-800kPa, 更优选高 200-600kPa， 最优 选高 200-400kPa或 200-300kPa,加热炉管出口温度为 360-460°C、优选 380-430°C。 预热后的原料油由压力式进料*** 3 引入蒸馏塔底部， 所 述的压力式进料***包括流量分配*** 4和雾化设备 5， 预热后的原料 油经流量分配*** 4按一定比例进行分配后由雾化设备 5 雾化为小液 滴， 喷入减压蒸馏塔汽化段 1 1， 并迅速汽化， 由于雾滴具有极大的比 表面积，在汽化段雾滴运动过程中可汽化的馏分在极短的时间内充分汽 化。 在所述的雾化设备 5上方设置破沫元件 9， 雾化设备 5的下方设置 液体收集元件 10。 在汽化段 1 1汽化后的馏分向上进入减压蒸馏塔的洗 涤段 12和分馏段 13， 分馏后从塔顶或侧线引出得到馏分油产品。 洗涤 段 12和分馏段 13结构与常规减压塔相同。较难汽化的重馏分则保持液 相状态， 雾滴相互碰撞过程中不断聚集形成大的液滴， 在液体收集元件 10的作用下收集落到塔底， 作为渣油被引出。 以下参考图 5， 以减压蒸馏为例， 说明本发明的一种实施方案， 其 中雾化容器为转油线。 待分馏的原料油（如常压渣油）经进料泵 1打入 加热炉 2中预热， 加热炉 2炉管内压力比汽化段高 100-1000kPa， 优选 高 200-800kPa， 更优选高 200-600kPa， 最优选高 200-400kPa 或 200-300kPa, 加热炉管出口温度为 360_460°C、 优选 380-430°C。 预热后 的原料油由压力式进料*** 3喷入转油线 7 中， 转油线 7 中的压力为 2.0-60.0 kPa,温度为 230-460° (：。雾滴在低油汽分压的条件下充分汽化， 汽化后的蒸汽物流引入减压分馏塔 6的汽化段 8 , 该实施方式可以使雾 滴充分汽化， 从而提高减压分馏塔的拔出率。
以下参考图 6， 以减压蒸馏为例， 说明本发明的一种实施方案， 其 中雾化容器为闪蒸罐并且与图 5 中的雾化容器为转油线的方案不同的 是， 预热后的原料油由压力式进料*** 3 喷入闪蒸罐 9 中， 闪蒸罐 9 中的压力为 2.0-60.0 kPa, 温度为 230-460°C。 由于雾滴具有极大的比表 面积， 在闪蒸罐中低油汽分压的条件下， 其中沸点较低的馏分被闪蒸汽 化。 经充分汽化后的蒸汽物流引入减压分馏塔 6的汽化段 8 , 该实施方 式可以使雾滴充分汽化， 从而提高减压分馏塔的拔出率。
以下参考图 7， 以减压蒸馏为例， 说明本发明的一种实施方案。 该 实施方案类似于图 6中的雾化容器为闪蒸罐的实施方案， 区别在于在闪 蒸罐 9中， 沸点较低的馏分被闪蒸汽化， 雾滴中未被汽化的馏分通过相 互碰撞重新聚集成较大的液滴落入闪蒸罐底部， 如图 7所示， 将闪蒸罐 内的气相物流从罐顶或贴近罐顶的壁面处通过管线 10引入减压分馏塔 6 的汽化段 8， 而罐底液相物流通过管线 1 1直接输送到减压分馏塔的塔釜 与其中的减压渣油汇合。该实施方式可以使闪蒸罐内未汽化的重组分雾 滴与蒸汽物流得到更好地分离，从而使减压分馏塔内雾沫夹带进一步减 少。 对比例 1 ,
对比例 1说明现有技术中的常压蒸馏方法分馏混合原油的效果。 待分馏的混合原油的性质见表 1。 图 1 为现有技术中常压分馏方法 的流程示意图， 如图 1所示， 混合原油首先由常压加热炉 2加热， 加热 炉出口温度为 368 °C， 经转油线 7进入常压蒸馏塔 8。 所述的常压蒸馏 塔为板式塔， 直径 6.5米，有三个侧线和两个中段回流， 获得直馏汽油、 煤油、 柴油等馏分， 常压蒸馏塔操作条件及产品性质见表 2。 常压蒸馏 塔的拔出率为 30.2 % 。 实施例 1
实施例 1说明本发明提供的方法用于原油常压蒸馏的效果。
图 2为本发明提供的方法用于常压蒸馏过程的流程示意图， 如图 2 所示， 采用的常压蒸馏塔 8与对比例 1相同， 待分馏的原料油与对比例 1 相同， 原料油由常压炉 2加热后在高于蒸馏塔汽化段压力 500kPa的条件 下经压力式进料***（包括流量分配*** 4和雾化设备 5 )喷入常压蒸馏 塔 8， 常压蒸馏塔内安装雾化设备， 所述雾化设备为旋流式雾化喷嘴， 旋流芯置于喷嘴前部， 旋流芯顶端安装有单孔板，被旋流的液体经孔喷 出后形成锥形液膜， 由于具有较大的径向速率和角向速率， 液膜与周围 气体速度差导致的摩擦将液膜撕扯成细小雾滴， 实现液相的良好雾化。 常压蒸馏塔操作条件及产品性质见表 2。 表 1 : 混合原油性质
Figure imgf000011_0001
蒸馏塔操作条件及产 项目 对比例 1 实施例 1 塔顶残压， kPa (绝） 170.0 170.0
全塔压降， kPa 27.0 27.0
汽化段压力， kPa (绝） 197.0 197.0
常压炉出 口压力， kPa
246.1 412.5
(绝）
常压炉出口温度， °c 368.0 372.0
常压炉炉管表面温度， °c 568.0 550.2
汽化段温度， °C 365.5 364.8
塔顶温度， V 1 18.1 1 19.5
常压一线抽出温度， °c 193.1 193.8
常压二线抽出温度， 253.4 255.9
常压三线抽出温度， °c 304.0 308.5
塔底温度， V 352.1 353.9
产品
产品收率， w /0
常顶油 5.0 5.2
常一线 7.0 7.2
常二线 9.9 1 1.0
常三线 8.3 9.8
常底油 69.8 66.8
常压拔出率 30.2 33.2 由表 2可见， 本发明提供的方法用于常压蒸馏时， 与常规进料的常 压蒸镏方法相比， 常压加热炉出口压力提高 166.4kPa， 常压炉出口温度 提高 4.0°C。 在蒸馏塔汽化段温度和压力基本相同的情况下， 蒸馏塔的 拔出率达到 33.2 % , 比常规进料的常压蒸馏拔出率提高 3 %。 本发明提 供的方法用于常压蒸镏塔， 可以提高常压塔的拔出率。 对比例 2
对 匕例 2说明现有技术中减压分馏常压渣油的效果。
待分馏的原料油为常压渣油， 性质见表 3。 图 3为现有技术中减压 蒸馏方法的流程示意图， 如图 3所示， 常压塔底油通过减压炉 2加热， 减压炉出口压力为 30.0kPa (绝） ， 减压炉管表面温度为 593 °C , 减压 炉出口温度为 410°C , 预热后的原料油经转油线 7引入减压蒸馏塔 6。 减压炉炉管从 φ 152mm不断扩径到 <D273mm， 转油线直径为 2.0m， 长 度 33.0m。 进料经蒸馏塔内进料分布器进行气液分离。 减压蒸馏塔为常 规全填料塔， 直径 9.2米， 干式操作。 所述的减压蒸馏塔分为汽化段、 洗涤段和分馏段， 所述的汽化段温度为 393.7°C。 洗涤段装填 ZUPAC2 系列填料 （天津大学北洋化工设备有限公司 ） 1.5米， 分馏段装填两层 ZUPAC 1填料（天津大学北洋化工设备有限公司 ）。 减压塔包括四个出 料口从上至下为减顶、 减一线、 减二线、 减三线， 以及两个中段回流。 塔顶抽真空***采用三级抽真空。减压蒸镏塔的操作条件及产品性质见 表 4。 减压蒸馏塔的拔出率为 57.6 % 。 实施例 2
实施例 2说明本发明提供的方法用于减压蒸馏塔的效果。
图 4为本发明提供的方法用于减压蒸馏过程的流程示意图。 待分熘 的原料油为常压渣油， 与对比例 2相同。 原料油经减压炉 2加热， 炉管直 径为 (D152mm， 加热后的原料油进入转油线， 然后在高于蒸馏塔汽化段 压力 300kPa的条件下经压力式进料*** （包括流量分配*** 4和雾化设 备 5 )喷入减压蒸馏塔 6， 减压蒸镏塔内安装雾化设备， 雾化设备如实施 例 1所述。 减压蒸馏塔的操作条件及产品性质见表 4。
表 3 : 常压渣油性质
Figure imgf000014_0001
表 4 减压蒸馏塔操作条件及产品性质 项目 对比例 2 实施例 2 塔顶残压， kPa (绝 ) 2.6 2.6 全塔压降， kPa (绝） 1.1 1.1 汽化段压力， kPa (绝） 3.7 3.7 减压炉出口压力， kPa (绝） 30.0 279.0 减压炉入口压力， kPa (绝） 470.0 470.0 常底泵出口压力， MPa (绝） 1.05 1.05 减压炉出口温度， °C 410.0 428.0 减压炉炉管表面温度， 。c 593.0 560.0 汽化段温度， °C 393.7 392.0 塔顶温度， °C 55.0 49.1 减压一线抽出温度， °c 1 16.1 120.5 减压二线抽出温度， °c 232.6 237.1 减压三线抽出温度， °c 312.7 320.8 塔底温度， V 374.5 376.8 产品
产品收率， w %
不凝气 0.3 0.2 减一线 5.2 5.8 项目 对比例 2 实施例 2 减二线 34.1 35.1
减三线 18.0 19.1
减； ΐ查油 42.4 39.8
拔出率， w % 57.6 60.2
馏份油性质
密度 （20°C ) ， kg/m3 905.3 912.4
混合蜡油残炭， ％ ( W ) 0.2 0.5
混合蜡油 的 C7 不 溶物
60.0 120.0
( mg/kg )
混合蜡油重金属含量（ mg/kg ) 0.2 0.5
混合蜡油馏程 ASTM D6352
初愤点 282 282
50% 439 447
终馏点 540 565
；查油性质
渣油密度 （20°C ) ， kg/m3 977.3 985.8
渣油 100°C运动粘度， mm2/s 857.0 1 189.0
；'查油残炭， （mg/kg ) 18 22
渣油 <500°C 馏分含量， ％ 4.3 1.3
渣油 500-550°C 熘分含量， ％ 12.6 8.7
渣油 550-600°C 熘分含量， ％ 18.0 14.1
渣油 >600°C 馏分含量， ％ 65.1 75.9 由表 4可见， 本发明提供的方法用于减压蒸馏时， 与对比例 2常规进料 的减压蒸馏方法相比， 在相同汽化段温度和压力下， 减压蒸馏塔的拔出 率达到 60.2 % , 与常规进料相比提高了 2.6 %。 减压加热炉出口温度提 高 18 °C , 炉管表面温度降低了 33 °C， 减压塔塔顶产品不凝气量从 0.3% 减至 0.2%; 常规进料时， 减压炉炉管逐级扩径， 较为复杂， 而采用本 发明提供的方法， 其炉管管径和转油线直径均为（D 152mm, 筒化了炉管 和转油线的结构； 此外， 与对比例 2 相比， 减压蜡油的终馏点提高了
25 °C , 其密度、 粘度、 重金属含量、 残炭均有提高， 但仍然满足下游装 置原料要求。 减压渣油中 500°C以下馏分含量从 4.3%降低到 1.3% , 600°C以上馏分含量从 65.1%增加到 75.9% , 渣油的密度、 粘度、 残炭 均有较大提高。 对比例 3
对比例 3说明现有技术中的减压分馏方法分馏常压分馏塔塔底油的 效果。 '
将待分馏的混合原油引入常压分馏塔， 分馏得到直馏汽油、 煤油、 柴油馏分， 常压塔拔出率为 32νν %。 类似于图 5但包括压力式进料系 统 3， 将常压分馏塔塔底油通过油泵 1输送到减压蒸馏***加热炉 2， 加热后通过转油线 7引入减压蒸馏塔汽化段 8。 加热炉炉管出口压力为 30.0kPa, 炉壁温度为 561°C， 炉出口温度为 386°C, 加热炉炉管逐级扩 径。减压蒸馏塔为高效全填料塔，减压蒸馏塔汽化段 8的温度为 374°C。 混合原油的性质见表 5， 减压蒸馏塔操作条件及产品性质见表 6。 减压 蒸馏塔的拔出率为 29.8w %。 实施例 3
实施例 3说明本发明提供的方法用于原油减压蒸馏的效果。
所用的常压塔***和待分馏的混合原油与对比例 3相同， 常压塔拔 出率为 32 wi%。 如图 5所示， 先将常压蒸馏塔分馏得到的塔底油通过油 泵 1输送到减压蒸馏***加热炉 2， 加热后的常压塔底油通过喷嘴 5喷入 转油线 7, 常压塔底油在转油线内充分汽化， 再通过转油线引入减压塔 汽化段 8。 转油线入口处压力为 14.0kPa， 温度为 386°C。 所用的喷嘴为 离心式雾化喷嘴； 加热炉的炉管没有变径。 所用的转油线及减压塔结构 与对比例 3相同， 减压塔汽化段的温度为 381°C。
从表 6可见， 通过在转油线上设置喷嘴， 在减压塔汽化段压力与对 比例 3相同的情况下， 加热炉管的出口压力达到 280.0 kPa， 炉壁温度为 556°C, 较对比例 34氏 5 °C。 而加热炉的出口温度达到 418 °C， 高于对比例 3达 22°C, 喷入转油线的雾滴通过在转油线内闪蒸汽化， 进入减压塔汽 化段仍能维持与对比例 3基本相同的温度。在汽化段压力与对比例 3相同 的情况下， 实施例 3中原料通过减压蒸馏***后拔出率达到 33.7 wt% , 高于对比例 3达 3.9个百分点。 减压渣油密度和粘度提高， 减压渣油中小 于 500 °C馏分的质量含量， 也从对比例 3的 10 %， 降低到 5.8 % 。 实施例 4
实施例 4说明本发明提供的方法用于原油减压蒸馏的效果。
所用的常压塔***和待分镏的混合原油与对比例 3相同， 常压塔拔 出率为 32 wi %。 如图 6所示， 所用减压塔结构与对比例 3相同， 所用加 热炉结构与实施例 3相同。 所不同的是在减压炉后增设一个闪蒸罐 9， 常 压塔底油由流量分配*** 4进行流量分配后经喷嘴 5喷入闪蒸罐内，经充 分汽化后， 引入减压蒸馏塔汽化段， 其中闪蒸罐压力为 6.1 kPa， 温度为 382 °C , 其他主要操作条件和产品性质见表 6。
由表 6中数据可以看出，实施例 4通过在加热炉出口设置雾化喷嘴和 闪蒸罐， 使常压塔底油中减压馏分油的拔出率为 34.5wi%， 和对比例 3 相比提高了 4.7个百分点。 实施例 5
实施例 5说明本发明提供的方法用于原油减压蒸馏的效果。
所用的常压塔***和待分馏的混合原油与对比例 3相同， 常压塔拔 出率为 32 wi %。 实施例 5所用减压塔结构同实施例 4, 所用加热炉结构 同实施例 4 , 所用闪蒸罐同实施例 4。 如图 7所示， 常压塔底油由流量分 配*** 3进衧流量分配后经喷嘴 5喷入闪蒸罐 9内， 经充分汽化后， 将气 体和液体从不同的管线分别引入减压蒸馏塔， 其中闪蒸罐压力为 6.1kPa, 温度为 382 °C。 该实施例的主要操作条件和产品性质见表 6。
由表 6中数据可以看出，实施例 5通过在加热炉出口设置喷雾式压力 进料***和闪蒸罐， 常压塔底油中减压馏分油的拔出率为 35.1 νί%， 和 对比例 3相比提高了 5.3个百分点。 原油的性质
Figure imgf000018_0001
表 6: 减压蒸馏***操作条件及产品性质
Figure imgf000019_0001

Claims

权 利 要 求
1. 一种在蒸馏塔中提高石油烃类馏分油收率的方法， 所述的蒸镏 塔包括汽化段和分馏段， 所述方法包括将待分馏的石油烃原料油预热， 经过压力式进料***在高于蒸馏塔汽化段压力 100-1000kPa， 优选 200-800kPa, 更优选 200-600kPa， 最优选 200-400kPa或 200-300kPa的 条件下进入蒸馏塔的汽化段，雾化同时汽化进而在蒸镏塔的分镏段进行 蒸馏分离， 塔顶和 /或侧线引出馏分油产品， 塔底引出未汽化的重质油。
2. 按照权利要求 1的方法， 其特征在于所述的蒸馏塔为不设置再 沸器的蒸馏塔， 优选包括闪蒸塔、 初馏塔、 常压蒸馏塔、 减压蒸馏塔或 加氢生成油蒸馏塔。
3. 按照权利要求 1的方法，其特征在于所述的蒸馏塔操作条件为： 蒸馏塔塔顶绝对压力为 0.5-240kPa、 汽化段绝对压力为 l -280kPa、 汽化 段温度为 150-430°C ; 具体地， 在常压蒸馏情况下， 塔顶绝对压力为 1 10- 180kPa、 汽化段绝对压力为 130-200kPa， 汽化段温度为 330-390°C ; 和在减压蒸馏情况下， 塔顶绝对压力为 0.5-90kPa, 优选地 0.5- 10kPa， 汽化段绝对压力为 l -98kPa， 优选地 l -5kPa， 汽化段温度为 300-430°C， 优选地 370-410 °C。
4. 按照权利要求 1的方法， 其特征在于所述的预热是通过加热炉 进行的， 所述的加热炉出 口压力比汽化段高 100-1000kPa， 优选
200-800kPa, 更优选 200-600kPa， 最优选 200-400kPa或 200-300kPa， 加热炉出口温度为 360-460°C、 优选 380-430°C。
5. 按照权利要求 1的方法， 其特征在于所述的蒸馏塔内， 喷入原 料油的入口下方设置液体收集元件， 和 /或喷入原料油的入口上方设置 破沫元件。
6. 按照权利要求 1的方法， 其特征在于所述的压力式进料***包 括流量分配***和雾化设备， 所述的雾化设备可以在蒸馏塔汽化段内， 或者在蒸馏塔外， 或者两者兼之。
7. 按照权利要求 6的方法， 其特征在于所述的雾化设备是伸入蒸 镏塔汽化段内的一个或多个喷嘴或其它可以使重油雾化的设备， 和 /或 伸入位于蒸馏塔塔外的且与蒸馏塔连通的雾化容器的一个或多个喷嘴 或其它可以使重油雾化的设备。
8. 按照权利要求 7的方法， 其特征在于所述的流量分配***被放 置在塔内和 /或在雾化容器外和 /或在雾化容器内。
9. 按照权利要求 7的方法，其特征在于所述的雾化容器为转油线、 闪蒸罐或闪蒸塔。
10. 按照权利要求 6 的方法， 其特征在于所述的雾化设备包括伸 入位于蒸馏塔塔外的且与蒸馏塔连通的雾化容器的一个或多个喷嘴或 其它可以使重油雾化的设备；在所述的雾化容器中形成的汽相物流进入 蒸馏塔汽化段，形成的液相物流直接进入分馏塔的底部与塔底的渣油混 合，或者在所述的雾化容器中形成的汽相物流和形成的液相物流从同一 个管线进入蒸馏塔汽化段。
1 1. 按照权利要求 6 的方法， 其特征在于所述的雾化设备包括伸 入蒸馏塔汽化段内的一个或多个喷嘴或其它可以使重油雾化的设备，待 分馏的石油烃原料油预热后，经过压力式进料***在高于蒸馏塔汽化段 压力 100- l OOOkPa的条件下雾化， 同时全部或部分汽化而进入蒸馏塔汽 化段， 塔顶和 /或侧线引出馏分油产品， 塔底引出未汽化的重质油。
12. 一种用于提高石油烃类馏分油收率的蒸馏塔， 所述的蒸馏塔 包括汽化段，其特征在于所述的蒸馏塔包括用于使待分馏的石油烃原料 油以高于蒸馏塔汽化段压力 100- l OOOkPa , 优选 200-800kPa , 更优选 200-600kPa, 最优选 200-400kPa或 200-300kPa的压力进料的压力式进 料***。
13. 按照权利要求 12的蒸馏塔， 其特征在于所述的蒸馏塔为不设 置再沸器的蒸馏塔， 优选包括闪蒸塔、 初馏塔、 常压蒸馏塔、 减压蒸馏 塔或加氢生成油蒸馏塔。
14. 按照权利要求 12的蒸馏塔， 其特征在于所述的蒸馏塔内， 喷 入原料油的入口下方设置液体收集元件， 和 /或喷入原料油的入口上方 设置破沫元件。
15. 按照权利要求 12的蒸馏塔， 其特征在于所述的压力式进料系 统包括流量分配***和雾化设备，所述的雾化设备可以在蒸馏塔汽化段 内， 或者在蒸馏塔外， 或者两者兼之。
16. 按照权利要求 12的蒸馏塔， 其特征在于所述的雾化设备是伸 入蒸馏塔汽化段内的一个或多个喷嘴或其它可以使重油雾化的设备，和 /或伸入位于蒸馏塔塔外的且与蒸馏塔连通的雾化容器的一个或多个喷 嘴或其它可以使重油雾化的设备。
17. 按照权利要求 16 ό 蒸馏塔， 其特征在于所述的流量分配*** 被放置在塔内和 /或在雾化容器外和 /或在雾化容器内。
18. 按照权利要求 16的蒸馏塔， 其特征在于所述的雾化容器为转 油线、 闪蒸罐或闪蒸塔。
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Citations (5)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
GB324376A (en) * 1928-08-21 1930-01-21 Standard Oil Co Indiana Improvements in or relating to the distillation of hydrocarbon oils
CN101376068A (zh) * 2008-09-17 2009-03-04 天津大学 带有减压闪蒸塔的常减压蒸馏方法及设备
EP2174697A1 (en) * 2008-10-10 2010-04-14 Heurtey Petrochem S.A. Petroleum residue recycling process and unit.
CN101717658A (zh) * 2009-12-03 2010-06-02 中国石油天然气集团公司 一种油品分馏塔的多次汽化进料方法
CN102079984A (zh) * 2009-11-26 2011-06-01 中国石油化工股份有限公司 一种提高馏分油收率的分馏塔进料方法

Family Cites Families (14)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US1394488A (en) * 1919-07-10 1921-10-18 Edward H French Process of fractionally distilling mixtures containing hydrocarbons
US2740753A (en) * 1949-10-07 1956-04-03 Teerverwertung Mit Beschrankte Continuous fractionation of tar
JPS515001B1 (zh) * 1970-04-21 1976-02-17
JPS515001A (ja) 1974-06-29 1976-01-16 Casio Computer Co Ltd Johosentakuyomidashihoshiki
US4053392A (en) * 1976-08-12 1977-10-11 Uop Inc. Fractionation process
US4234391A (en) * 1978-10-13 1980-11-18 University Of Utah Continuous distillation apparatus and method
JPS55123684A (en) * 1979-03-15 1980-09-24 Osaka Gas Co Ltd Separator of heavy oil
JPS5731987A (en) * 1980-08-04 1982-02-20 Idemitsu Kosan Co Ltd Method and apparatus for distilling crude oil
JPS636086A (ja) * 1986-06-26 1988-01-12 Mitsui Eng & Shipbuild Co Ltd 原油の蒸留方法
US5306418A (en) * 1991-12-13 1994-04-26 Mobil Oil Corporation Heavy hydrocarbon feed atomization
JPH111691A (ja) * 1997-06-12 1999-01-06 Mitsubishi Heavy Ind Ltd ガスタービン燃料処理装置
WO2001042393A1 (fr) * 1998-08-03 2001-06-14 Jgc Corporation Procede et dispositif de traitement du petrole
JP2001129301A (ja) * 1999-11-09 2001-05-15 Minoru Yoshimoto 減圧蒸発分離装置
CN1600831A (zh) * 2003-09-25 2005-03-30 田原宇 重油固相热载体循环裂解和气化工艺

Patent Citations (5)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
GB324376A (en) * 1928-08-21 1930-01-21 Standard Oil Co Indiana Improvements in or relating to the distillation of hydrocarbon oils
CN101376068A (zh) * 2008-09-17 2009-03-04 天津大学 带有减压闪蒸塔的常减压蒸馏方法及设备
EP2174697A1 (en) * 2008-10-10 2010-04-14 Heurtey Petrochem S.A. Petroleum residue recycling process and unit.
CN102079984A (zh) * 2009-11-26 2011-06-01 中国石油化工股份有限公司 一种提高馏分油收率的分馏塔进料方法
CN101717658A (zh) * 2009-12-03 2010-06-02 中国石油天然气集团公司 一种油品分馏塔的多次汽化进料方法

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