Method and system for separating alpha-olefin from Fischer-Tropsch light oil

Abstract

The invention relates to a method and system for separating alpha-olefin from Fischer-Tropsch light oil, and the method comprises the following steps: S1, carrying out first fraction cutting on the Fischer-Tropsch light oil to obtain C7-C9 fractions and C10-C12 fractions; s2, carrying out oxygen-containing compound removal treatment on the C7-C9 fraction and the C10-C12 fraction by adopting hollow fiber membrane extraction, so as to obtain a C7-C9 deoxidized fraction and a C10-C12 deoxidized fraction; s3, respectively carrying out second fraction cutting on the C7-C9 deoxidized fraction and the C10-C12 deoxidized fraction, so as to obtain a plurality of deoxidized alkene material flows after cutting; and S4, respectively carrying out alkane-olefin adsorption separation treatment on the cut deoxidized alkane-olefin material flow by adopting a sequential simulated moving bed. The method disclosed by the invention is low in energy consumption, wide in application range, high in product yield and purity and low in oxygen-containing compound content.

Description

Technical Field

[0001] This disclosure relates to the field of Fischer-Tropsch light oil separation, and more specifically, to a method and system for separating α-olefins from Fischer-Tropsch light oil. Background Technology

[0002] The naphtha fraction of Fischer-Tropsch synthesis products contains a large amount of α-olefins. These α-olefins have advantages such as continuous carbon number distribution, high content of straight-chain alkenes, and absence of sulfur and nitrogen heteroatoms and dienes, making them high-quality raw materials for developing downstream high-end polyolefin materials, polyα-olefin lubricants, α-alkenyl sulfonates, and higher alcohols. Due to the difficulty in separating and purifying α-olefins, industrially, Fischer-Tropsch light oil rich in α-olefins is usually cracked to produce ethylene and propylene, but the α-olefins are not effectively utilized. Compared with the ethylene oligomerization process for producing α-olefins, separating α-olefin products from Fischer-Tropsch oil has advantages such as a shorter process flow, lower raw material costs, and a complete range of odd and even carbon atoms. This can compensate for the limitation of ethylene oligomerization, which can only produce even-numbered carbon α-olefins, and leverage the unique properties of odd-numbered carbon α-olefins in downstream product production. Therefore, developing purification and separation technologies for α-olefins in Fischer-Tropsch light oil is of great significance.

[0003] The current process technology combinations can be summarized into two main categories: (1) Deoxygenation process with extraction as the core combined with alkane-alkene separation process with distillation as the core. However, the extractive distillation method used in the alkene-alkane separation unit of this type of method has a relatively complex process route and high operating energy consumption, resulting in high investment and operating costs; for products with different carbon numbers, the extractant used in extractive distillation is specific and can only separate C6-C8 components, and cannot completely separate components with other carbon numbers. (2) Deoxygenation process combined with simulated moving bed adsorption alkane-alkene separation process. In this type of method, oxygen-containing compounds are removed using a simulated moving bed adsorption method. Oxygen-containing compounds have a strong binding effect with the adsorbent, which can easily lead to adsorbent saturation and reduce its stability. The oxygen-containing compound content in Fischer-Tropsch light oil is about 2-6 wt%. If the content of one of the components to be separated is significantly low, it will lead to an imbalance between the two streams of extract and raffinate, increasing operating costs and weakening the separation effect. In the alkene separation stage, the material entering is an oil product with a mixed number of carbon atoms. The effect of simulated moving bed adsorption in separating alkenes and alkenes with mixed carbon atoms is worse than that in separating alkenes with a single carbon atom, which affects the purity of the α-olefin product. Summary of the Invention

[0004] The purpose of this disclosure is to provide a method and system for separating α-olefins from Fischer-Tropsch light oil.

[0005] To achieve the above objectives, the first aspect of this disclosure provides a method for separating α-olefins from Fischer-Tropsch light oil, the method comprising the following steps: S1. The Fischer-Tropsch light oil is fractionated into C7-C9 and C10-C12 fractions. S2. The C7-C9 fraction and the C10-C12 fraction are subjected to oxygen-containing compound removal treatment by hollow fiber membrane extraction to obtain C7-C9 deoxygenated fraction and C10-C12 deoxygenated fraction; S3. The C7-C9 deoxygenated fraction and the C10-C12 deoxygenated fraction are respectively subjected to a second fractionation to obtain multiple deoxyalkene streams after fractionation. S4. The cut deoxyalkane stream is subjected to alkane adsorption and separation treatment using a sequential simulated moving bed.

[0006] Optionally, in step S2, the oxygen-containing compound removal treatment of the C7-C9 fraction using hollow fiber membrane extraction includes: filling the tubular side of the membrane module of the first hollow fiber membrane extractor with a first extractant, so that the membrane pores of the membrane module are filled with the first extractant; then injecting a first preparation solution containing the first extractant and a first stripping agent into the tubular side of the membrane module; then injecting the C7-C9 fraction into the shell side of the membrane module; and washing the outflow from the shell side outlet with water to obtain the deoxygenated C7-C9 fraction. The first extractant comprises an aqueous methanol solution, wherein the concentration of methanol is 50-90% by weight; the first stripping agent comprises water; and in the first preparation solution, the weight ratio of the first extractant to the first stripping agent is 2-10. The first hollow fiber membrane extractor has a packing factor of 0.1-0.6 and a mass flow rate ratio of 0.5-4 between the tube side and the shell side.

[0007] Optionally, in step S2, the oxygen-containing compound removal treatment of the C10-C12 fraction using hollow fiber membrane extraction includes: filling the tube side of the membrane module of the second hollow fiber membrane extractor with a second extractant, so that the membrane pores of the membrane module are filled with the second extractant; then injecting a second preparation solution containing the second extractant and a second stripping agent into the tube side of the membrane module; then injecting the C10-C12 fraction into the shell side of the membrane module; and washing the outflow from the shell side outlet with water to obtain the deoxygenated C10-C12 fraction. The second extractant comprises an aqueous ethanol solution, wherein the concentration of the ethanol is 50-85% by weight; the second stripping agent comprises water; and in the second preparation solution, the weight ratio of the second extractant to the second stripping agent is 1-3. The second hollow fiber membrane extractor has a packing factor of 0.1-0.6 and a mass flow rate ratio of 0.5-4 between the tube side and the shell side.

[0008] Optionally, the hollow fiber membrane is made of at least one of polysulfone, polyethersulfone, and polyvinylidene fluoride, and has an inner diameter of 0.2-3.0 mm, an outer diameter of 0.3-4.0 mm, and a porosity of 30-80%.

[0009] Optionally, step S4 includes: performing a first adsorption separation treatment on the first deoxyalkane stream obtained by cutting the C7-C9 deoxygenated fraction using a first sequential simulated moving bed; The second deoxyalkene stream obtained by cutting the C10-C12 deoxygenated fraction is subjected to a second adsorption separation treatment using a second sequential simulated moving bed.

[0010] Optionally, the number of adsorption columns in the first sequential simulated moving bed is 6-12, and the first adsorbent used includes one or more of CaNaA type molecular sieves, NaX type molecular sieves and NaY type molecular sieves. The first desorbent used includes a first desorbing component and a first solvent; the first desorbing component includes one or more C6 hydrocarbons, and the first solvent includes n-hexane and / or cyclohexane. The content of the first desorbing component is 5-60% by weight relative to the weight of the first desorbent.

[0011] Optionally, the operating temperature of the first sequential simulated moving bed is 20-40 ℃, the operating pressure is 0.1-0.4 MPa, and the mass flow ratio of the first deoxyalkane stream to the first desorbent is 0.1-5.

[0012] Optionally, the number of adsorption columns in the second sequential simulated moving bed is 8-16, and the second adsorbent used includes one or more of CaNaA type molecular sieves, NaX type molecular sieves and NaY type molecular sieves. The second desorbent used includes a second desorbing component and a second solvent; the second desorbing component includes one or more of C7 hydrocarbons and C8 hydrocarbons, and the second solvent includes n-heptane and / or n-octane; The content of the second desorbing component is 10-70% by weight relative to the weight of the second desorbent.

[0013] Optionally, the operating temperature of the second sequential simulated moving bed is 20-40 ℃, the operating pressure is 0.1-0.4 MPa, and the mass flow ratio of the second deoxyalkene stream to the second desorbent is 0.1-5.

[0014] Optionally, the Fischer-Tropsch light oil has an initial boiling point of 20-30°C, a final boiling point of 200-230°C, and an oxygen content of 0.1-10% by weight.

[0015] The second aspect of this disclosure provides a system for the method for separating α-olefins from Fischer-Tropsch light oil as described in the first aspect of this disclosure, the system comprising a first fractionation unit, an oxygen-containing compound removal unit, a second fractionation unit, and an adsorption separation unit; The first fractionation unit is used to fractionate the Fischer-Tropsch light oil to obtain C7-C9 fractions and C10-C12 fractions. The oxygen-containing compound removal unit is used to remove oxygen-containing compounds from the C7-C9 fraction and the C10-C12 fraction to obtain the C7-C9 deoxygenated fraction and the C10-C12 deoxygenated fraction. The oxygen-containing compound removal unit includes a hollow fiber membrane extraction unit; The second fraction cutting unit is used to cut the C7-C9 deoxygenated fraction and the C10-C12 deoxygenated fraction to obtain multiple cut deoxyalkane streams; The adsorption separation unit includes a sequential simulated moving bed for adsorption and separation of the deoxyalkane stream.

[0016] Through the above technical solution, this disclosure employs hollow fiber membrane extraction to remove oxygen-containing compounds from Fischer-Tropsch light oil. The removal effect is stable, with the feed liquid and solvent phases flowing on both sides of the membrane, avoiding liquid dispersion and aggregation, and reducing entrainment loss of the extractant. Simultaneously, the hollow fiber membrane has a larger mass transfer surface area, improving mass transfer efficiency and deoxygenation efficiency. After oxygen-containing compound removal, the resulting deoxygenated alkane-alkene stream is first cut, and then a sequential simulated moving bed is used for alkane-alkene separation, achieving better separation results. Furthermore, the sequential simulated moving bed has lower energy consumption and is widely used for separating alkanes and alkenes with different carbon groups (C7-C12). The intermittent feed and intermittent discharge mode requires fewer packed columns, allows for more flexible adjustment of operating parameters, effectively solves the backmixing problem within the system, increases product purity and yield, and reduces energy and solvent consumption.

[0017] Other features and advantages of this disclosure will be described in detail in the following detailed description section. Attached Figure Description

[0018] The accompanying drawings are provided to further illustrate the present disclosure and form part of the specification. They are used together with the following detailed description to explain the present disclosure, but do not constitute a limitation thereof. In the drawings: Figure 1 This is a process flow diagram of one embodiment of the present disclosure for separating α-olefins from Fischer-Tropsch light oil.

[0019] Explanation of reference numerals in the attached figures A: Fischer-Tropsch light oil coarse cut distillation column; B: C5-C9 coarse cut distillation column; C: First hollow fiber membrane extractor; C': Second hollow fiber membrane extractor; D: First water washing column; D': Second water washing column; E: First extractant recovery column I; F: First extractant recovery column II; E': Second extractant recovery column I; F': Second extractant recovery column II; G: C7-C9 single carbon cut distillation column I; H: C7-C9 single carbon cut distillation column II; G': C10-C12 single carbon cut distillation column I; H': C10-C12 single carbon cut distillation column I; Distillation column II; I: C7 deoxyalkene stream storage tank; J: C8 deoxyalkene stream storage tank; K: C9 deoxyalkene stream storage tank; I': C10 deoxyalkene stream storage tank; J': C11 deoxyalkene stream storage tank; K': C12 deoxyalkene stream storage tank; L: First multi-port switching valve; L': Second multi-port switching valve; M: First sequential simulated moving bed; M': Second sequential simulated moving bed; N: First hydrocarbon by-product distillation column; O: First olefin product distillation column; N': Second hydrocarbon by-product distillation column; O': Second olefin product distillation column 1: Fischer-Tropsch light oil; 2: C5-C9 fraction; 3: C10-C12 fraction; 4: C7 fraction; 5: C7-C9 fraction; 6: First extractant; 7: C7-C9 deoxygenated fraction to be treated; 8: First processing liquid; 9: Regenerated first extractant; 10: Mixture of water and first oxygenated compound; 11: Regenerated water; 12: First oxygenated compound stream; 13: Water; 14: C7-C9 deoxygenated fraction; 15: The first... 16: Washing solution; 17: C7 deoxyalkene stream; 18: C8 deoxyalkene stream; 19: C9 deoxyalkene stream; 20: C7, C8, or C9 deoxyalkene stream; 21: First desorbent; 22: Mixture of first olefin product and first desorbent; 23: Mixture of first hydrocarbon by-product and first desorbent; 24: Regenerated first desorbent; 25: First olefin product; 26: Regenerated first desorbent. 1. Desorbent; 27. First hydrocarbon byproduct; 28. Second extractant; 29. ​​C10-C12 deoxygenated fraction to be treated; 30. Second processing liquid; 31. Regenerated second extractant; 32. Mixture of water and second oxygenated compound; 33. Regenerated water; 34. Second oxygenated compound stream; 35. Water; 36. C10-C12 deoxygenated fraction; 37. Second washing liquid; 38. C10 deoxyalkane stream; 39. C11- 40: C12 deoxyalkene stream; 41: C12 deoxyalkene stream; 42: C10, C11, or C12 deoxyalkene stream; 43: Second desorbent; 44: Second olefin product and second desorbent mixture; 45: Second hydrocarbon by-product and second desorbent mixture; 46: Regenerated second desorbent; 47: Second olefin product stream; 48: Regenerated second desorbent; 49: Second hydrocarbon by-product Detailed Implementation The specific embodiments of this disclosure will be described in detail below with reference to the accompanying drawings. It should be understood that the specific embodiments described herein are for illustration and explanation only and are not intended to limit this disclosure.

[0020] The first aspect of this disclosure provides a method for separating α-olefins from Fischer-Tropsch light oil, the method comprising the following steps: S1. The Fischer-Tropsch light oil is fractionated into C7-C9 and C10-C12 fractions. S2. The C7-C9 fraction and the C10-C12 fraction are subjected to oxygen-containing compound removal treatment using hollow fiber membranes to obtain C7-C9 deoxygenated fraction and C10-C12 deoxygenated fraction; S3. The C7-C9 deoxygenated fraction and the C10-C12 deoxygenated fraction are respectively subjected to a second fractionation to obtain a deoxyalkene stream; S4. The deoxyalkane stream is subjected to alkane adsorption and separation treatment using a sequential simulated moving bed.

[0021] This disclosure employs hollow fiber membrane extraction to remove oxygen-containing compounds from Fischer-Tropsch light oil. Compared to adsorption deoxygenation methods, the hollow fiber membrane extraction method of this disclosure exhibits more stable deoxygenation performance over long-term operation, overcoming the drawback of adsorbent saturation and reduced deoxygenation stability in adsorption methods. Compared to traditional extraction methods, the feed liquid phase and solvent phase flow on both sides of the membrane, avoiding liquid dispersion and aggregation, and reducing extractant entrainment losses. It also boasts a larger mass transfer specific surface area, 10-50 times larger than that in general extraction equipment, improving mass transfer efficiency and enabling same-stage extraction and back-extraction membrane processes. After removing oxygen-containing compounds, the resulting deoxyalkane-alkene stream is first cleaved, and then a sequential simulated moving bed is used for alkane-alkene separation. Unlike existing technologies that first separate mixed-carbon-number alkane-alkene streams and then perform single-carbon cleavage, the steps of this disclosure achieve better separation results. Compared with extractive distillation for alkane and alkene separation, sequential simulated moving bed adsorption separation has lower energy consumption and is suitable for separating alkanes and alkenes with different carbon groups (C7-C12). Compared with traditional simulated moving bed adsorption separation, sequential simulated moving bed adopts an intermittent feeding and discharging mode, requiring fewer packing columns and allowing for more flexible adjustment of operating parameters. It effectively solves the problem of backmixing of materials within the system, resulting in higher product purity and lower energy and solvent consumption.

[0022] According to one embodiment of this disclosure, the Fischer-Tropsch light oil has an initial boiling point of 20-30°C and a final boiling point of 200-230°C, and the Fischer-Tropsch light oil contains oxygenated compounds, C5-C12 alkanes and C5-C12 olefins.

[0023] According to one embodiment of this disclosure, the content of oxygenated compounds in the Fischer-Tropsch light oil is 0.1-10% by weight, and the oxygenated compounds include alcohols, ketones, esters, etc.

[0024] According to one embodiment of this disclosure, in step S1, the cutting temperature of the C7-C9 fraction is 85-160°C, and the cutting temperature of the C10-C12 fraction is 165-225°C.

[0025] According to one embodiment of this disclosure, step S1 can also obtain C7. - The fraction is cut at a temperature of 25-80℃.

[0026] According to one embodiment of this disclosure, step S1 is carried out in a distillation column, and the specific steps and conditions are conventional in the art.

[0027] According to one embodiment of this disclosure, in step S2, the removal of oxygenated compounds from the C7-C9 fraction using hollow fiber membrane extraction includes: filling the tubular side of the membrane module of the first hollow fiber membrane extractor with a first extractant, so that the membrane pores of the membrane module are filled with the first extractant; then injecting a first preparation solution containing the first extractant and a first back-extraction agent into the tubular side of the membrane module; then injecting the C7-C9 fraction into the shell side of the membrane module; the outflow from the shell side outlet is washed with water to obtain the C7-C9 deoxygenated fraction; and the first treatment liquid flows out from the tubular side outlet. The specific steps of water washing include: passing the outflow from the shell side outlet into a water washing tower for water washing; obtaining the first washing liquid at the top of the water washing tower and obtaining the C7-C9 deoxygenated fraction at the bottom of the tower. The first extractant comprises an aqueous methanol solution with a methanol concentration of 50-90% by weight. The first stripping agent comprises water. In the first preparation solution, the weight ratio of the first extractant to the first stripping agent is 2-10. The first hollow fiber membrane extractor has a packing factor of 0.1-0.6 and a mass flow rate ratio of tube side to shell side of 0.5-4.

[0028] According to one embodiment of the present disclosure, the method further includes: recovering the first extractant from the first treatment liquid and the first washing liquid, wherein the recovered first extractant can be recycled, and the recovery can be carried out in a recovery tower to improve the utilization efficiency of the first extractant.

[0029] According to one embodiment of this disclosure, step S2, the removal of oxygenated compounds from the C10-C12 fraction using hollow fiber membrane extraction, includes: filling the tubular side of the membrane module of the second hollow fiber membrane extractor with a second extractant, so that the membrane pores of the membrane module are filled with the second extractant; then injecting a second preparation solution containing the second extractant and a second back-extraction agent into the tubular side of the membrane module; then injecting the C10-C12 fraction into the shell side of the membrane module; the outflow from the shell side outlet is washed with water to obtain the deoxygenated C10-C12 fraction; and the second treatment solution flows out from the tubular side outlet. The specific steps of the water washing process include: feeding the shell-side effluent into a water washing tower for washing; obtaining a second washing liquid at the top of the water washing tower and a C10-C12 deoxygenated fraction at the bottom; the second extractant comprising an aqueous ethanol solution with a concentration of 50-85% by weight; the second stripping agent comprising water; and the weight ratio of the second extractant to the second stripping agent in the second preparation solution being 1-3; the packing factor of the second hollow fiber membrane extractor being 0.1-0.6, and the mass flow rate ratio of the tube side to the shell side being 0.5-4.

[0030] According to one embodiment of this disclosure, the method further includes: recovering a second extractant from the second treatment liquid and the second washing liquid, wherein the recovered second extractant can be recycled, and the recovery can be carried out in a recovery tower to improve the utilization efficiency of the first extractant.

[0031] According to one embodiment of this disclosure, the hollow fiber membrane is made of at least one of polysulfone (PSF), polyethersulfone (PES), and polyvinylidene fluoride (PVDF). The hollow fiber membrane has an inner diameter of 0.2-3.0 mm, an outer diameter of 0.3-4.0 mm, and a porosity of 30-80%. Using a hollow fiber membrane with the above structure can further improve the removal efficiency of oxygen-containing compounds.

[0032] According to one embodiment of this disclosure, step S3, cutting the C7-C9 deoxygenated fraction, may include: cutting the C7-C9 deoxygenated fraction to obtain C7 deoxyalkane stream, C8 deoxyalkane stream, and C9 deoxyalkane stream; or, cutting the C7-C9 deoxygenated fraction to obtain C7-C8 deoxyalkane stream and C9 deoxyalkane stream; or, cutting the C7-C9 deoxygenated fraction to obtain C7 deoxyalkane stream and C8-C9 deoxyalkane stream; preferably, cutting the C7-C9 deoxygenated fraction to obtain C7 deoxyalkane stream, C8 deoxyalkane stream, and C9 deoxyalkane stream; the cutting can be carried out in a distillation column.

[0033] According to one embodiment of this disclosure, step S3, cutting the C10-C12 deoxygenated fraction, may include: cutting the C10-C12 deoxygenated fraction to obtain C10 deoxyalkane stream, C11 deoxyalkane stream, and C12 deoxyalkane stream; or, cutting the C10-C12 deoxygenated fraction to obtain C10-C11 deoxyalkane stream and C12 deoxyalkane stream; or, cutting the C10-C12 deoxygenated fraction to obtain C10 deoxyalkane stream and C11-C12 deoxyalkane stream; preferably, cutting the C10-C12 deoxygenated fraction to obtain C10 deoxyalkane stream, C11 deoxyalkane stream, and C12 deoxyalkane stream; the cutting can be carried out in a distillation column.

[0034] According to one embodiment of this disclosure, step S4 includes: performing a first adsorption separation treatment on the first deoxygenated alkene stream obtained by cutting the C7-C9 deoxygenated fraction using a first sequential simulated moving bed. The first deoxygenated alkene stream can be any one of C7 alkene stream, deoxygenated C8 alkene stream, deoxygenated C9 alkene stream, deoxygenated C7-C8 alkene stream, and deoxygenated C8-C9 alkene stream, preferably any one of C7 alkene stream, deoxygenated C8 alkene stream, and deoxygenated C9 alkene stream, i.e., deoxygenated single-carbon alkene stream.

[0035] According to a specific embodiment of this disclosure, the number of adsorption columns in the first sequential simulated moving bed is 6-12, preferably 6 or 8; the first adsorbent used includes one or more of CaNaA type molecular sieves, NaX type molecular sieves, and NaY type molecular sieves, and when the first adsorbent includes two or more, their proportion is not specifically limited; the first desorbent used includes a first desorption component and a first solvent; the first desorption component includes one or more of C6 hydrocarbons, and when the first desorption component includes two or more, their proportion is not specifically limited, the first desorption component is preferably cyclohexene, and the first solvent includes n-hexane and / or cyclohexane; relative to the weight of the first desorbent, the content of the first desorption component is 5-60% by weight, and the specific combination of adsorbent and desorbent can improve the alkane-alkene separation effect, thereby improving the purity and yield of olefin products.

[0036] In order to obtain better alkane-alkene separation efficiency and improve the yield and purity of olefin products, according to a specific embodiment of the present disclosure, the operating temperature of the first sequential simulated moving bed is 20-40 °C, the operating pressure is 0.1-0.4 MPa, and the mass flow ratio of the first deoxyalkane stream to the first desorbent is 0.1-5, preferably 0.5-2.

[0037] According to one embodiment of this disclosure, step S4 includes: performing a second adsorption separation treatment on the second deoxygenated alkene stream obtained by cutting the C10-C12 deoxygenated fraction using a second sequential simulated moving bed. The second deoxygenated alkene stream can be any one of C10 alkene stream, deoxygenated C11 alkene stream, deoxygenated C12 alkene stream, deoxygenated C10-C11 alkene stream, and deoxygenated C11-C12 alkene stream, preferably any one of C10 alkene stream, deoxygenated C11 alkene stream, and deoxygenated C12 alkene stream, i.e., deoxygenated single-carbon alkene stream.

[0038] According to a specific embodiment of this disclosure, the number of adsorption columns in the second sequential simulated moving bed is 8-16, preferably 10 or 12; the second adsorbent used includes one or more of CaNaA type molecular sieves, NaX type molecular sieves, and NaY type molecular sieves, and when the first adsorbent includes two or more, their proportions are not specifically limited; the second desorbent used includes a second desorption component and a second solvent; the second desorption component includes one or more of C7 hydrocarbons and C8 hydrocarbons, and when the second desorption component includes two or more, their proportions are not specifically limited, the second desorption component is preferably toluene or ethylbenzene, and the second solvent includes n-heptane and / or n-octane; relative to the weight of the second desorbent, the content of the second desorption component is 10-70% by weight, and the specific combination of adsorbent and desorbent can improve the alkane-alkene separation effect, thereby improving the purity and yield of olefin products.

[0039] In order to obtain better alkane-alkene separation efficiency and improve the yield and purity of olefin products, according to a specific embodiment of the present disclosure, the operating temperature of the second sequential simulated moving bed is 20-40 °C, the operating pressure is 0.1-0.4 MPa, and the mass flow ratio of the second deoxyalkane stream to the second desorbent is 0.1-5, preferably 0.5-2.

[0040] According to one embodiment of this disclosure, the sequential simulated moving bed has four regions: Region I (adsorption zone), Region II (distillation zone), Region III (desorption zone), and Region IV (buffer zone), where the sum of the number of adsorption columns in the four regions equals the total number of adsorption columns. The complete switching time for each adsorption column includes the following three sub-steps: (a) All-in-all-out stage: The four zones are separated. The front end of Zone III receives the material to be separated, and the rear end of Zone III discharges the residual liquid. The front end of Zone I receives the desorbent, and the rear end of Zone I discharges the extract. (b) Small circulation stage: Zones I, II and III are connected in series, and Zone IV is separated. The desorbent enters at the front end of Zone I and the residual liquid is discharged at the rear end of Zone III; (c) Large circulation stage: 4 regions are connected in series and closed, with no inflow or outflow, and the mobile phase circulates throughout the system; After the three sub-steps are completed, all inlets and outlets are moved back by one adsorption column (moving to the first adsorption column when reaching the last adsorption column), forming a new area division. The above three stages are then repeated in sequence. The time ratio of the third stage (large circulation stage) to the first stage (all-in, all-out stage) is 1-15, preferably 2-5; the time ratio of the third stage (large circulation stage) to the second stage (small circulation stage) is 1-20, preferably 5-10; and the time ratio of the first stage (all-in, all-out stage) to the second stage (small circulation stage) is 0.1-5, preferably 3-5.

[0041] According to one embodiment of this disclosure, the method further includes: passing the raffinate from the first sequential simulated moving bed into a distillation column for distillation, obtaining a regenerated first desorbent at the top of the column, which is returned to the first sequential simulated moving bed for recycling, and obtaining a by-product at the bottom of the column; and the evaporate from the first sequential simulated moving bed entering a distillation column for distillation, obtaining a regenerated first desorbent at the top of the column, which is returned to the first sequential simulated moving bed for recycling, and obtaining an α-olefin product at the bottom of the column.

[0042] According to one embodiment of this disclosure, the method further includes: passing the raffinate from the second sequential simulated moving bed into a distillation column for distillation, obtaining a regenerated second desorbent at the top of the column, which is returned to the second sequential simulated moving bed for recycling, and obtaining a by-product at the bottom of the column; and the evaporate from the second sequential simulated moving bed entering a distillation column for distillation, obtaining a regenerated second desorbent at the top of the column, which is returned to the second sequential simulated moving bed for recycling, and obtaining an α-olefin product at the bottom of the column.

[0043] The second aspect of this disclosure provides a system for the method for separating α-olefins from Fischer-Tropsch light oil as described in the first aspect of this disclosure, the system comprising a first fractionation unit, an oxygen-containing compound removal unit, a second fractionation unit, and an adsorption separation unit; The first fractionation unit is used to fractionate the Fischer-Tropsch light oil to obtain C7-C9 fractions and C10-C12 fractions. The oxygen-containing compound removal unit is used to remove oxygen-containing compounds from the C7-C9 fraction and the C10-C12 fraction to obtain the C7-C9 deoxygenated fraction and the C10-C12 deoxygenated fraction. The oxygen-containing compound removal unit includes a hollow fiber membrane extraction unit; The second fraction cutting unit is used to cut the C7-C9 deoxygenated fraction and the C10-C12 deoxygenated fraction to obtain multiple cut deoxyalkane streams; The adsorption separation unit includes a sequential simulated moving bed for adsorption and separation of the cut deoxyalkene stream.

[0044] According to one embodiment of this disclosure, the first fractionation unit includes multiple distillation columns to fractionate Fischer-Tropsch light oil to obtain C7. - The product comprises C5-C9 and C10-C12 fractions. Specifically, it may include two distillation columns. Fischer-Tropsch light oil is fed into a Fischer-Tropsch light oil coarse-cut distillation column to coarsely cut the oil into C5-C9 and C10-C12 fractions. The outlet of the Fischer-Tropsch light oil coarse-cut distillation column is connected to the inlet of the C5-C9 coarse-cut distillation column, allowing the C5-C9 fraction to be fed into the C5-C9 coarse-cut distillation column to further cut the oil into C7-C9 fractions. - Fractions and C7-C9 fractions.

[0045] According to one embodiment of this disclosure, the hollow fiber membrane extraction unit includes a first hollow fiber membrane extractor and a second hollow fiber membrane extractor; the inlet of the first hollow fiber extractor is connected to the C7-C9 fraction outlet of the C5-C9 coarse cutting distillation column, for removing oxygen-containing compounds from the C7-C9 fraction; the first hollow fiber extractor is also provided with a first extractant inlet and a first preparation liquid inlet for introducing the first extractant and the first preparation liquid; the inlet of the second hollow fiber extractor is connected to the C10-C12 fraction outlet of the first fraction cutting unit, for removing oxygen-containing compounds from the C10-C12 fraction; the second hollow fiber extractor is also provided with a second extractant inlet and a second preparation liquid inlet for introducing the second extractant and the second preparation liquid.

[0046] According to one embodiment of the present disclosure, the oxygen-containing compound removal unit further includes a first water washing tower, a second water washing tower, a first extractant recovery tower I, a first extractant recovery tower II, a second extractant recovery tower I, and a second extractant recovery tower II; The inlet of the first water washing tower is connected to the outlet of the C7-C9 deoxygenated fraction to be treated in the first hollow fiber membrane extractor, and the outlet of the first water washing liquid is connected to the inlet of the first extractant recovery tower I. The C7-C9 fraction after deoxygenation is washed with water to recover the first extractant. The outlet of the C7-C9 deoxygenated fraction of the first water washing tower is connected to the inlet of the second fraction cutting unit. The first water washing tower is also equipped with a water inlet for water to be introduced for washing. The inlet of the second water washing tower is connected to the outlet of the C10-C12 deoxygenated fraction to be treated in the second hollow fiber membrane extractor, and the outlet of the second water washing liquid is connected to the inlet of the second extractant recovery tower I. The C10-C12 fraction after deoxygenation is washed with water to recover the second extractant. The outlet of the C10-C12 deoxygenated fraction of the second water washing tower is connected to the inlet of the second fraction cutting unit. The second water washing tower is also equipped with a water inlet for water to be introduced for washing. The inlet of the first extractant recovery tower I is connected to the outlet of the first washing liquid and the outlet of the first processed liquid of the first hollow fiber membrane extractor to recover the first extractant in the first washing liquid and the first processed liquid. The regenerated first extractant outlet of the first extractant recovery tower I is connected to the first extractant inlet on the first hollow fiber extractor. The inlet of the first extractant recovery tower II is connected to the outlet of the mixture of water and the first oxygen-containing compound in the first extractant recovery tower I, so as to separate the oxygen-containing compound in the mixture. At the same time, the reclaimed water obtained can be recycled. The inlet of the second extractant recovery tower I is connected to the outlet of the second washing liquid and the outlet of the second treatment liquid of the second hollow fiber membrane extractor to recover the second extractant in the second washing liquid and the second treatment liquid. The regenerated second extractant outlet of the second extractant recovery tower I is connected to the second extractant inlet on the second hollow fiber extractor. The inlet of the second extractant recovery tower II is connected to the outlet of the mixture of water and oxygenated compounds in the second extractant recovery tower I, which separates the oxygenated compounds in the mixture. At the same time, the reclaimed water obtained can be recycled.

[0047] According to one embodiment of this disclosure, the second fraction cutting unit includes a C7-C9 deoxygenated fraction cutting unit and a C10-C12 deoxygenated fraction cutting unit. The C7-C9 deoxygenated fraction cutting unit is used to cut the C7-C9 deoxygenated fraction, and the deoxygenated alkane stream obtained from the cutting is subjected to a further alkane-alkene separation. The C10-C12 deoxygenated fraction cutting unit is used to cut the C10-C12 deoxygenated fraction, and the deoxygenated alkane stream obtained from the cutting is subjected to a further alkane-alkene separation. Depending on actual usage requirements, the C7-C9 deoxygenated fraction cutting unit and the C10-C12 deoxygenated fraction cutting unit may include one or more distillation columns.

[0048] According to one specific embodiment of this disclosure, the C7-C9 deoxygenated fraction cutting unit includes two distillation columns to cut the C7-C9 deoxygenated fraction into C7 deoxyalkane streams, C8 deoxyalkane streams, and C9 deoxyalkane streams; or, the C7-C9 deoxygenated fraction cutting unit includes one distillation column to cut the C7-C9 deoxygenated fraction into C7-C8 deoxyalkane streams and C9 deoxyalkane streams; or, the C7-C9 deoxygenated fraction to cut into C7 deoxyalkane streams and C8-C9 deoxyalkane streams.

[0049] According to one specific embodiment of this disclosure, the C10-C12 deoxygenated fraction cutting unit includes two distillation columns to cut the C10-C12 deoxygenated fraction into C10 deoxyalkane streams, C11 deoxyalkane streams, and C12 deoxyalkane streams; or, the C10-C12 deoxygenated fraction cutting unit includes one distillation column to cut the C10-C12 deoxygenated fraction into C10-C11 deoxyalkane streams and C12 deoxyalkane streams; or, the C10-C12 deoxygenated fraction is cut into C10 deoxyalkane streams and C11-C12 deoxyalkane streams.

[0050] According to one embodiment of this disclosure, the adsorption separation unit includes a first sequential simulated moving bed and a second sequential simulated moving bed. The inlet of the first sequential simulated moving bed is connected to the outlet of the deoxygenated alkene stream of the C7-C9 deoxygenated fraction cutting unit, and is used for alkene adsorption separation of the deoxygenated alkene stream. The inlet of the second sequential simulated moving bed is connected to the outlet of the deoxygenated alkene stream of the C10-C12 deoxygenated fraction cutting unit, and is used for alkene adsorption separation of the deoxygenated alkene stream.

[0051] According to one embodiment of this disclosure, the system further includes an olefin purification unit, which comprises a first hydrocarbon by-product distillation column, a first olefin product distillation column, a second hydrocarbon by-product distillation column, and a second olefin product distillation column. The inlet of the first hydrocarbon by-product distillation column is connected to the raffinate outlet of the first sequential simulated moving bed, separating and purifying to obtain a first desorbent and a first hydrocarbon by-product, with the first desorbent returned to the system for recycling. The inlet of the first olefin product distillation column is connected to the raffinate outlet of the first sequential simulated moving bed, separating and purifying to obtain a first desorbent and a first olefin product, with the first desorbent returned to the system for recycling. The inlet of the second hydrocarbon by-product distillation column is connected to the raffinate outlet of the second sequential simulated moving bed, separating and purifying to obtain a second desorbent and a second hydrocarbon by-product, with the second desorbent returned to the system for recycling. The inlet of the second olefin product distillation column is connected to the raffinate outlet of the second sequential simulated moving bed, separating and purifying to obtain a second desorbent and a second olefin product, with the second desorbent returned to the system for recycling.

[0052] The present disclosure is further illustrated by the following examples, but the present disclosure is not limited thereto. All raw materials used in the examples are commercially available.

[0053] In the embodiments, the initial boiling point of Fischer-Tropsch light oil is 25°C, the final boiling point is 220°C, and the total content of oxygenated compounds in Fischer-Tropsch light oil is 2.83% by weight. The specific distribution is shown in Table 1. In Table 1, the content of oxygenated compounds refers to the content of oxygenated compounds in the corresponding carbon fraction of Fischer-Tropsch light oil.

[0054] Table 1

[0055] Method for testing the purity of α-olefins: gas chromatography.

[0056] Test method for oxygen-containing compound content: full two-dimensional gas chromatography.

[0057] The method for testing the yield of α-olefins is: (extract flow rate × α-olefin concentration) / (feed flow rate × α-olefin concentration) × 100%.

[0058] Test method for the content of α-olefin products in the mixture: gas chromatography.

[0059] Example 1 Adopting such Figure 1 The process shown involves the following steps for separating α-olefins from Fischer-Tropsch light oil: (1) The Fischer-Tropsch light oil shown in Table 1 was subjected to coarse fractionation to obtain C7. - Fractions, C7-C9 fractions, C10-C12 fractions: (a) First, Fischer-Tropsch light oil 1 enters Fischer-Tropsch light oil coarse-cut distillation column A, and is coarsely cut into C5-C9 fraction 2 and C10-C12 fraction 3: The Fischer-Tropsch light oil coarse-cut distillation column A operates at atmospheric pressure and has 50 theoretical plates. Fischer-Tropsch light oil is fed from the 29th theoretical plate at a feed rate of 10 kg / h, with a reflux ratio of 2 and a reboiler temperature of 160°C. The overhead effluent is C5-C9 fraction 2 at a flow rate of 9.1 kg / h; the bottom effluent is C10-C12 fraction 3 at a flow rate of 0.9 kg / h. (b) The outlet of the Fischer-Tropsch light oil coarse cut distillation column is connected to the inlet of the C5-C9 coarse cut distillation column. The C5-C9 fraction 2, which is the overhead effluent from the Fischer-Tropsch light oil coarse cut distillation column A, enters the C5-C9 coarse cut distillation column B for further coarse cutting into C7 fractions. - Fraction 4 and C7-C9 fraction 5: The C5-C9 coarse-cut distillation column B operates at atmospheric pressure and has 40 theoretical plates. The C5-C9 fraction is fed from the 20th theoretical plate at a feed rate of 9.1 kg / h. The reflux ratio is 1.5, and the reboiler temperature is 110°C. The overhead effluent is C7. - Fraction 4 has a flow rate of 1.7 kg / g; the bottom effluent is C7-C9 fraction 5, with a flow rate of 7.4 kg / h.

[0060] (2) The C7-C9 fraction and the C10-C12 fraction were subjected to oxygen-containing compound removal using hollow fiber membranes, and the extractant was regenerated and recycled to obtain the C7-C9 deoxygenated fraction and the C10-C12 deoxygenated fraction, respectively: (a) The inlet of the first hollow fiber extractor is connected to the outlet of the C7-C9 fraction of the C5-C9 coarse cutting distillation column. The C7-C9 fraction 5 enters the first hollow fiber membrane extractor C. The length of the extractor is 500 mm, the inner diameter is 20 mm, the hollow fiber membrane material is polysulfone (PSF), the inner diameter is 0.3 mm, the outer diameter is 0.4 mm, the porosity is 80%, and the packing factor is 0.5. A 65% by weight methanol-water solution was used as the first extractant 6 and injected into the tube side of the membrane module to fill the membrane pores with the first extractant. Then, a first preparation solution containing the first extractant and the first stripping agent water (the weight ratio of methanol-water solution to water was 9:1) was injected into the tube side of the membrane module at an inlet flow rate of 7 kg / h. The C7-C9 fraction 5 was injected into the membrane module from the shell side at a flow rate of 7.4 kg / h. The two streams were operated in countercurrent. The first processed liquid 8 was obtained from the tube side outlet, and the untreated C7-C9 deoxygenated fraction 7 was obtained from the shell side outlet. The mass flow rate ratio of the tube side to the shell side was 0.94. (b) The inlet of the first water washing tower D is connected to the outlet of the C7-C9 deoxygenated fraction to be treated in the first hollow fiber membrane extractor. The C7-C9 deoxygenated fraction 7 to be treated is injected into the first water washing tower D, and water 13 is injected to remove the first extractant by water washing countercurrent contact. The operating temperature is 40°C, and the mass flow ratio of water to the C7-C9 deoxygenated fraction to be treated is 1:1. C7-C9 deoxygenated fraction 14 is obtained at the bottom of the first water washing tower D. (c) The outlet of the first washing liquid in the first water washing tower D and the outlet of the first hollow fiber membrane extractor are connected to the inlet of the first extractant recovery tower. The first washing liquid 15 is obtained at the top of the first water washing tower D, and enters the first extractant recovery tower IE together with the first processed liquid 8. The first extractant recovery tower IE has 40 theoretical plates, operates at atmospheric pressure, has a bottom temperature of 80°C, and a reflux ratio of 1.5. The top of the tower yields regenerated methanol, i.e., the regenerated first extractant 9 is recycled. The bottom of the first extractant recovery tower IE yields a mixture 10 of water and the first oxygen-containing compound, which enters the first extractant recovery tower II F. The first extractant recovery tower II F has 60 theoretical plates, operates at atmospheric pressure, has a bottom temperature of 105°C, and a reflux ratio of 5. The top of the tower yields regenerated water 11, which is recycled, and the bottom of the tower yields the first oxygen-containing compound stream 12. (d) The inlet of the second hollow fiber extractor is connected to the outlet of the C10-C12 fraction of the first fraction cutting unit. The C10-C12 fraction 3 enters the second hollow fiber membrane extractor C'. The length of the extractor is 400 mm, the inner diameter is 10 mm, the hollow fiber membrane material is polyethersulfone (PES), the inner diameter is 0.25 mm, the outer diameter is 0.35 mm, and the filling factor is 0.6. A 75% wt% ethanol-water solution was used as the second extractant 28 and injected into the tube side of the membrane module to fill the membrane pores. Then, a second preparation solution (ethanol-water solution to water weight ratio of 1.5:1) containing the second extractant and the second stripping agent water was injected into the tube side of the membrane module at an inlet flow rate of 1.5 kg / h. The C10-C12 fraction 3 was injected into the membrane module from the shell side at a flow rate of 0.9 kg / h. The two streams were operated in countercurrent. The tube side outlet yielded the second processed liquid 30, and the shell side outlet yielded the C10-C12 deoxygenated fraction 29 to be treated. The mass flow rate ratio of the tube side to the shell side was 1.67. (e) The inlet of the second water washing tower D' is connected to the outlet of the C10-C12 deoxygenated fraction to be treated in the second hollow fiber membrane extractor. The C10-C12 deoxygenated fraction 29 to be treated is injected into the second water washing tower D', and water 35 is injected to remove the second extractant by water washing countercurrent contact. The operating temperature is 50°C, and the mass flow ratio of water to the C10-C12 deoxygenated fraction to be treated is 1.5. C10-C12 deoxygenated fraction 36 is obtained at the bottom of the second water washing tower D'. (f) The outlet of the second washing liquid in the second washing tower D' and the outlet of the second hollow fiber membrane extractor are connected to the inlet of the second extractant recovery tower. The second washing liquid 37 is obtained at the top of the second washing tower D', and enters the second extractant recovery tower IE' together with the second processed liquid 30. IE' has 40 theoretical plates, operates at atmospheric pressure, has a bottom temperature of 90°C, and a reflux ratio of 2. The ethanol azeotrope is obtained at the top of the tower, which is used to regenerate the second extractant 31 for recycling. The mixture 32 of water and the second oxygenated compound is obtained at the bottom of the second extractant recovery tower IE' and enters the second extractant recovery tower II F', which has 50 theoretical plates, operates at atmospheric pressure, has a bottom temperature of 115°C, and a reflux ratio of 3. The regenerated water 33 is obtained at the top of the tower for recycling, and the second oxygenated compound stream 34 is obtained at the bottom of the tower.

[0061] (3) The C7-C9 deoxygenated fraction and the C10-C12 deoxygenated fraction were subjected to single-carbon distillation to obtain single-carbon C7 deoxyalkane streams, C8 deoxyalkane streams, C9 deoxyalkane streams, and single-carbon C10 deoxyalkane streams, C11 deoxyalkane streams, and C12 deoxyalkane streams: (a) The C7-C9 deoxygenated fraction 14 enters the C7-C9 single-carbon cutting distillation column IG, where it is cut into C7 deoxyalkane stream 16 and C8-C9 deoxyalkane stream 17. The single-carbon cutting distillation column IG operates at atmospheric pressure and has 30 theoretical plates. The C7-C9 deoxygenated fraction is fed from the 15th theoretical plate at a feed rate of 7.2 kg / h, with a reflux ratio of 1 and a bottom temperature of 110°C. The overhead effluent is C7 deoxyalkane stream 16 at a flow rate of 3.0 kg / h; the bottom effluent is C8-C9 deoxyalkane stream 17 at a flow rate of 4.2 kg / h. (b) C8-C9 deoxyalkane stream 17 enters the C7-C9 single-carbon cutting distillation column II H, where it is cut into C8 deoxyalkane stream 18 and C9 deoxyalkane stream 19. The C7-C9 single-carbon cutting distillation column II H operates at atmospheric pressure and has 25 theoretical plates. The C8-C9 deoxyalkane stream is fed from the 12th theoretical plate at a feed rate of 4.2 kg / h, with a reflux ratio of 1.5 and a bottom temperature of 135°C. The overhead effluent is C8 deoxyalkane stream 18 at a flow rate of 2.5 kg / h; the bottom effluent is C9 deoxyalkane stream 19 at a flow rate of 1.7 kg / h. (c) The C10-C12 deoxygenated fraction 36 enters the C10-C12 single-carbon cutting distillation column I G', which is cut into C10 deoxyalkane stream 38 and C11-C12 deoxyalkane stream 39. The C10-C12 single-carbon cutting distillation column I G' operates at atmospheric pressure and has 30 theoretical plates. The C10-C12 deoxygenated fraction is fed from the 15th theoretical plate at a feed rate of 0.9 kg / h, the reflux ratio is 2.5, and the reboiler temperature is 160°C. The overhead distillate is C10 deoxyalkane stream 38 at a flow rate of 0.6 kg / h. The bottom effluent is C11-C12 deoxyalkane stream 39 at a flow rate of 0.3 kg / h. (d) C11-C12 deoxyalkane stream 39 enters the C10-C12 single-carbon cutting distillation column II H', which is cut into C11 deoxyalkane stream 40 and C12 deoxyalkane stream 41. The C10-C12 single-carbon cutting distillation column II H' operates at atmospheric pressure and has 30 theoretical plates. The C11-C12 deoxyalkane stream is fed from the 18th theoretical plate at a feed rate of 0.3 kg / h, with a reflux ratio of 2 and a bottom temperature of 205°C. The overhead effluent is C11 deoxyalkane stream 40 at a flow rate of 0.2 kg / h; the bottom effluent is C12 deoxyalkane stream 41 at a flow rate of 0.1 kg / h.

[0062] (4) Single-carbon C7 deoxyalkene streams, C8 deoxyalkene streams, and C9 deoxyalkene streams are stored in C7 deoxyalkene stream storage tank I, C8 deoxyalkene stream storage tank J, and C9 deoxyalkene stream storage tank K, respectively. These streams are connected to a first sequential simulated moving bed M via a first multi-port switching valve L. C7, C8, or C9 deoxyalkene stream 20 is introduced into the first sequential simulated moving bed M, and alkane-alkene separation is performed using a first desorbent 21 in the first sequential simulated moving bed L. The raffinate from the first sequential simulated moving bed M yields a first hydrocarbon by-product and a first desorbent mixture 23, while the extract yields a first olefin product and a first desorbent mixture 22. Distillation separation yields C7, C8, or C9 α-olefin products. (a) The first sequential simulated moving bed M has 8 adsorption columns, and the total volume of the adsorption columns is 75 m³. 3 A 1:1 weight ratio of NaX molecular sieve and NaY molecular sieve was used as the adsorbent, and 30% wt% cyclohexene and 70% wt% cyclohexane were used as the first desorbent. The operating temperature was 30℃, and the operating pressure was atmospheric pressure. The feed flow rate of the first desorbent was 2.8 kg / h. The time for the first stage (all in, all out) was t1 = 250 s, the time for the second stage (small circulation) was t2 = 85 s, and the time for the third stage (large circulation) was t3 = 850 s. Based on the time conversion of the three stages, the actual feed rate of the first desorbent was 0.79 kg / h, specifically calculated as: actual feed rate of the first desorbent = (t1 + t2) / (t1 + t2 + t3) × feed flow rate of the first desorbent. (b) When the C7 deoxyalkene stream enters the first sequential simulated moving bed M, the flow rate is 3.0 kg / h. Based on the three-stage time, the feed rate is calculated to be 0.63 kg / h. Specifically, the actual feed rate of the C7 deoxyalkene stream = t1 / (t1+t2+t3) × the feed flow rate of the C7 deoxyalkene stream. The mass flow rate ratio of the C7 deoxyalkene stream to the first desorbent is 0.80. The raffinate yields a mixture of C7 by-product and the first desorbent at a flow rate of 0.83 kg / h; the extract yields a mixture of C7 α-olefin product and the first desorbent at a flow rate of 0.59 kg / h. The product mass fraction in the mixture is 42.4 wt%, the purity of the C7 α-olefin is 97.5 wt%, the content of oxygen-containing compounds is less than 1 ppm, and the yield is 90.9%. When the C8 deoxyalkene stream enters the first sequential simulated moving bed M, the flow rate is 2.5 kg / h. Based on the three-stage time conversion, the feed rate is 0.53 kg / h. Specifically, the conversion method is: actual feed rate of C8 deoxyalkene stream = t1 / (t1+t2+t3) × feed flow rate of C8 deoxyalkene stream. The mass flow rate ratio of C8 deoxyalkene stream to the first desorbent is 0.67. The raffinate yields a mixture of C8 by-product and the first desorbent at a flow rate of 0.73 kg / h; the extract yields a mixture of C8 α-olefin product and the first desorbent at a flow rate of 0.59 kg / h. The product mass fraction in the mixture is 39.4 wt%, the purity of C8 α-olefin is 96.2 wt%, the oxygen content is less than 1 ppm, and the yield is 88.5%. When the C9 deoxyalkene stream enters the first sequential simulated moving bed M, the flow rate is 1.7 kg / h. Based on the three-stage time conversion, the feed rate is 0.36 kg / h. Specifically, the conversion method is: actual feed rate of C9 deoxyalkene stream = t1 / (t1+t2+t3) × feed flow rate of C9 deoxyalkene stream. The mass flow rate ratio of C9 deoxyalkene stream to the first desorbent is 0.46. The raffinate yields a mixture of C9 hydrocarbon byproducts and the first desorbent at a flow rate of 0.56 kg / h; the extract yields a mixture of C9 α-olefin product and the first desorbent at a flow rate of 0.59 kg / h. The product mass fraction in the mixture is 36.8 wt%, the purity of C9 α-olefin is 95.4 wt%, the oxygen content is less than 1 ppm, and the yield is 86.2%. (c) The mixture 22 of the first olefin product and the first desorbent enters the first olefin product distillation column O, and the product is obtained under the same process parameters. It has 20 theoretical plates, the 10th theoretical plate is fed, the operating pressure is atmospheric pressure, the bottom temperature is 85°C, the regenerated first desorbent 24 is obtained at the top of the column and returned to the sequential simulated moving bed M for recycling; the first olefin product 25 is obtained at the bottom of the column.

[0063] The mixture 23 of the first hydrocarbon by-product and the first desorbent enters the first hydrocarbon by-product distillation column N, where the product is obtained under the same process parameters. The column has 25 theoretical plates, with the 17th theoretical plate as feed. The operating pressure is atmospheric pressure, and the bottom temperature is 85°C. The regenerated first desorbent 26 obtained at the top of the column is returned to the sequential simulated moving bed M for recycling. The first hydrocarbon by-product 27 is obtained at the bottom of the column.

[0064] (5) Single-carbon C10 deoxyalkene streams, C11 deoxyalkene streams, and C12 deoxyalkene streams are stored in C10 deoxyalkene stream storage tank I', C11 deoxyalkene stream storage tank J', and C12 deoxyalkene stream storage tank K', respectively. These streams are connected to a second sequential simulated moving bed M' via a second multi-port switching valve L'. C10, C11, or C12 deoxyalkene stream 42 is introduced into the second sequential simulated moving bed M', and alkane-alkene separation is performed using a second desorbent 43 in the second sequential simulated moving bed L'. The raffinate from the second sequential simulated moving bed M' yields a second hydrocarbon byproduct and a second desorbent mixture 45, while the extract yields a second olefin product and a second desorbent mixture 44. Distillation separation yields C10, C11, or C12 α-olefin products. (a) The second sequential simulated moving bed M' has 12 adsorption columns and a total adsorption column volume of 15 m³. 3 A mixture of NaX molecular sieve and CaNaA molecular sieve at a mass ratio of 3:1 was used as the adsorbent, and 20% by weight toluene and 80% by weight n-heptane were used as the second desorbent. The operating temperature was 40℃ and the operating pressure was 0.3 MPa. The feed flow rate of the second desorbent was 0.4 kg / h. The time for the first stage (all in, all out) was t1 = 150 s, the time for the second stage (small circulation) was t2 = 55 s, and the time for the third stage (large circulation) was t3 = 700 s. Based on the time conversion of the three stages, the actual feed rate of the second desorbent was 0.09 kg / h, specifically calculated as: actual feed rate of the second desorbent = (t1 + t2) / (t1 + t2 + t3) × feed flow rate of the second desorbent. (b) When the C10 deoxyalkene stream enters the second sequential simulated moving bed M', the flow rate is 0.6 kg / h. Based on the three-stage time conversion, the feed rate is 0.10 kg / h. The specific conversion method is: actual feed rate of C10 deoxyalkene stream = t1 / (t1+t2+t3) × feed flow rate of C10 deoxyalkene stream. The mass flow rate ratio of C10 deoxyalkene stream to the second desorbent is 1.11. The raffinate yields a mixture of C10 byproduct and the second desorbent at a flow rate of 0.12 kg / h; the extract yields a mixture of C10 α-olefin product and the second desorbent at a flow rate of 0.07 kg / h. The product mass fraction in the mixture is 45.6 wt%, the purity of C10 α-olefin is 96.9 wt%, the content of oxygen-containing compounds is less than 1 ppm, and the yield is 89.8%. When the C11 deoxyalkene stream enters the second sequential simulated moving bed M', the flow rate is 0.2 kg / h. Based on the three-stage time conversion, the feed rate is 0.03 kg / h. Specifically, the conversion method is: actual feed rate of C11 deoxyalkene stream = t1 / (t1+t2+t3) × feed flow rate of C11 deoxyalkene stream. The mass flow rate ratio of C11 deoxyalkene stream to the second desorbent is 0.33. The raffinate yields a mixture of C11 byproduct and the second desorbent at a flow rate of 0.05 kg / h; the extract yields a mixture of C11 α-olefin product and the second desorbent at a flow rate of 0.07 kg / h. The product mass fraction in the mixture is 40.7 wt%, the purity of C11 α-olefin is 95.8 wt%, the content of oxygen-containing compounds is less than 1 ppm, and the yield is 87.3%. When the C12 deoxyalkene stream enters the second sequential simulated moving bed M', the flow rate is 0.1 kg / h. Based on the three-stage time conversion, the feed rate is 0.02 kg / h. Specifically, the conversion method is: actual feed rate of C12 deoxyalkene stream = t1 / (t1+t2+t3) × feed flow rate of C12 deoxyalkene stream. The mass flow rate ratio of C12 deoxyalkene stream to the second desorbent is 0.22. The raffinate yields a mixture of C12 byproduct and the second desorbent at a flow rate of 0.04 kg / h; the extract yields a mixture of C12 α-olefin product and the second desorbent at a flow rate of 0.07 kg / h. The product mass fraction in the mixture is 41.2 wt%, the purity of C12 α-olefin is 95.1 wt%, the oxygen content is less than 1 ppm, and the yield is 85.6%. (c) The mixture 44 of the second olefin product and the second desorbent enters the second olefin product distillation column O', and the product is obtained under the same process parameters. It has 20 theoretical plates, with the 10th theoretical plate as feed. The operating pressure is atmospheric pressure, and the bottom temperature is 110 °C. The regenerated second desorbent 46 obtained at the top of the column is returned to the sequential simulated moving bed M' for recycling; the second olefin product 47 is obtained at the bottom of the column. The mixture 45 of the second hydrocarbon by-product and the second desorbent enters the second hydrocarbon by-product distillation column N', where the product is obtained under the same process parameters. The column has 25 theoretical plates, with the 17th theoretical plate as feed. The operating pressure is atmospheric pressure, and the bottom temperature is 110 °C. The regenerated second desorbent 48 obtained at the top of the column is returned to the sequential simulated moving bed M' for recycling. The second hydrocarbon by-product 49 is obtained at the bottom of the column.

[0065] Example 2 The method of Example 1 is adopted, except that the membrane material of the first hollow fiber membrane extractor C is polyvinylidene fluoride (PVDF), with an inner diameter of 0.4 mm, an outer diameter of 0.5 mm, a porosity of 75%, and a packing factor of 0.3. An 80% wt% methanol aqueous solution is used as the first extractant and is injected into the tube side of the membrane module to fill the membrane pores with the first extractant. Then, a first preparation solution of the first extractant and the first stripping agent water (the weight ratio of methanol aqueous solution to water is 8:1) is injected into the tube side of the membrane module with an inlet flow rate of 6.5 kg / h. The C7-C9 fraction 5 is injected into the membrane module from the shell side with a flow rate of 7.4 kg / h. The mass flow rate ratio of the tube side to the shell side is 0.88. The membrane material of the second hollow fiber membrane extractor C' is polyvinylidene fluoride (PVDF), with an inner diameter of 0.4 mm, an outer diameter of 0.5 mm, a porosity of 75%, and a packing factor of 0.4. A 60% by weight aqueous ethanol solution is used as the second extractant. Then, a second preparation of the second extractant and the second stripping agent water (the weight ratio of aqueous ethanol solution to water is 1:1) is injected into the tube side of the membrane module at an inlet flow rate of 2.0 kg / h. The C10-C12 fraction 3 is injected into the membrane module from the shell side at a flow rate of 0.9 kg / h. The mass flow rate ratio of the tube side to the shell side is 2.22.

[0066] The oxygen-containing compound content of the final C7-C12 α-olefins was less than 1 ppm.

[0067] Example 3 The method of Example 1 was adopted, except that 45% wt% cyclohexene and 55% wt% cyclohexane were used as the first desorbent, the operating pressure was atmospheric pressure, and the operating temperature was 35°C. The feed flow rate of the first desorbent was 2.0 kg / h, the first stage (all in, all out) time t1 = 200 s, the second stage (small circulation) time t2 = 100 s, and the third stage (large circulation) time t3 = 700 s. Based on the time conversion of the three stages, the actual feed rate of the first desorbent was 0.60 kg / h, specifically calculated as: actual feed rate of the first desorbent = (t1 + t2) / (t1 + t2 + t3) × feed flow rate of the first desorbent; (a) When the C7 deoxyalkene stream enters the first sequential simulated moving bed, the flow rate is 3.0 kg / h. Based on the three-stage time, the feed rate is calculated to be 0.60 kg / h. Specifically, the actual feed rate of the C7 deoxyalkene stream = t1 / (t1+t2+t3) × the feed flow rate of the C7 deoxyalkene stream. The mass flow rate ratio of the C7 deoxyalkene stream to the first desorbent is 1.0. The raffinate yields a mixture of C7 by-product and the first desorbent at a flow rate of 0.80 kg / h; the extract yields a mixture of C7 α-olefin product and the first desorbent at a flow rate of 0.40 kg / h. The product mass fraction in the mixture is 58.9 wt%, the purity of the C7 α-olefin is 98.6 wt%, the content of oxygen-containing compounds is less than 1 ppm, and the yield is 91.3%. When the C8 deoxyalkene stream enters the first sequential simulated moving bed, the flow rate is 2.5 kg / h. Based on the three-stage time conversion, the feed rate is 0.5 kg / h. Specifically, the conversion method is: actual feed rate of C8 deoxyalkene stream = t1 / (t1+t2+t3) × feed flow rate of C8 deoxyalkene stream. The mass flow rate ratio of C8 deoxyalkene stream to the first desorbent is 0.83. The raffinate yields a mixture of C8 by-product and the first desorbent at a flow rate of 0.70 kg / h; the extract yields a mixture of C8 α-olefin product and the first desorbent at a flow rate of 0.40 kg / h. The product mass fraction in the mixture is 51.4 wt%, the purity of C8 α-olefin is 97.7 wt%, the oxygen content is less than 1 ppm, and the yield is 90.1%. When the C9 deoxyalkene stream enters the first sequential simulated moving bed, the flow rate is 1.7 kg / h. Based on the three-stage time conversion, the feed rate is 0.34 kg / h. Specifically, the conversion method is: actual feed rate of C9 deoxyalkene stream = t1 / (t1+t2+t3) × feed flow rate of C9 deoxyalkene stream. The mass flow rate ratio of C7 deoxyalkene stream to the first desorbent is 0.57. The raffinate yields a mixture of C9 hydrocarbon byproducts and the first desorbent at a flow rate of 0.54 kg / h; the extract yields a mixture of C9 α-olefin product and the first desorbent at a flow rate of 0.40 kg / h. The product mass fraction in the mixture is 48.6 wt%, the purity of C9 α-olefin is 96.3 wt%, the content of oxygen-containing compounds is less than 1 ppm, and the yield is 88.3%. (b) Using 40% by weight toluene and 60% by weight n-heptane as the second desorbent, the operating temperature was 40°C and the operating pressure was 0.4 MPa. The feed flow rate of the second desorbent was 0.3 kg / h. The time for the first stage (all in, all out) was t1 = 170 s, the time for the second stage (small circulation) was t2 = 65 s, and the time for the third stage (large circulation) was t3 = 750 s. Based on the time conversion of the three stages, the actual feed rate of the second desorbent was 0.07 kg / h. The specific conversion method was: actual feed rate of the second desorbent = (t1 + t2) / (t1 + t2 + t3) × feed flow rate of the second desorbent; When the C10 deoxyalkene stream enters the second sequential simulated moving bed, the flow rate is 0.6 kg / h. Based on the three-stage time conversion, the feed rate is 0.10 kg / h. Specifically, the conversion method is: actual feed rate of C10 deoxyalkene stream = t1 / (t1+t2+t3) × feed flow rate of C10 deoxyalkene stream. The mass flow rate ratio of C10 deoxyalkene stream to the second desorbent is 1.43. The raffinate yields a mixture of C10 byproduct and the second desorbent at a flow rate of 0.12 kg / h; the extract yields a mixture of C10 α-olefin product and the second desorbent at a flow rate of 0.05 kg / h. The product mass fraction in the mixture is 52.7 wt%, the purity of C10 α-olefin is 97.8 wt%, the oxygen content is less than 1 ppm, and the yield is 90.5%. When the C11 deoxyalkene stream enters the second sequential simulated moving bed, the flow rate is 0.2 kg / h. Based on the three-stage time conversion, the feed rate is 0.03 kg / h. Specifically, the conversion method is: actual feed rate of C11 deoxyalkene stream = t1 / (t1+t2+t3) × feed flow rate of C11 deoxyalkene stream. The mass flow rate ratio of C11 deoxyalkene stream to the second desorbent is 0.43. The raffinate yields a mixture of C11 byproduct and the second desorbent at a flow rate of 0.05 kg / h; the extract yields a mixture of C11 α-olefin product and the second desorbent at a flow rate of 0.05 kg / h. The product mass fraction in the mixture is 43.2 wt%, the purity of C11 α-olefin is 96.4 wt%, the content of oxygen-containing compounds is less than 1 ppm, and the yield is 88.5%. When the C12 deoxyalkene stream enters the second sequential simulated moving bed, the flow rate is 0.1 kg / h. Based on the three-stage time conversion, the feed rate is 0.02 kg / h. Specifically, the conversion method is: actual feed rate of C12 deoxyalkene stream = t1 / (t1+t2+t3) × feed flow rate of C12 deoxyalkene stream. The mass flow rate ratio of C12 deoxyalkene stream to the second desorbent is 0.28. The raffinate yields a mixture of C12 byproduct and the second desorbent at a flow rate of 0.04 kg / h; the extract yields a mixture of C12 α-olefin product and the second desorbent at a flow rate of 0.05 kg / h. The product mass fraction in the mixture is 40.5 wt%, the purity of C12 α-olefin is 95.8 wt%, the content of oxygen-containing compounds is less than 1 ppm, and the yield is 86.4%.

[0068] Comparative Example Using Example 1 of patent CN 116836035 A, 1-hexene, 1-heptene, and 1-octene were separated from Fischer-Tropsch naphtha. The composition of the Fischer-Tropsch naphtha and the specific separation conditions are described in patent CN 116836035 A. The purity of the separated 1-hexene was 99.1% by weight, the content of oxygenated compounds was 4 ppm, and the yield of 1-hexene was 87.1%; the purity of 1-heptene was 98.9% by weight, the content of oxygenated compounds was 4 ppm, and the yield of 1-heptene was 88.2%; the purity of 1-octene was 98.8% by weight, the content of oxygenated compounds was 5 ppm, and the yield of 1-octene was 88.3%.

[0069] As can be seen from the above data, the method disclosed herein uses hollow fiber membrane extraction for deoxygenation. The feed liquid phase and solvent phase flow on both sides of the membrane, eliminating liquid dispersion and aggregation processes, thus reducing entrainment losses. It also features a larger mass transfer surface area, improving mass transfer efficiency, and achieving a deoxygenation effect of less than 1 ppm. The method disclosed herein employs a sequential simulated moving bed adsorption method for alkane-alkene separation, resulting in lower energy consumption, effectively improving product yield, and is widely applicable to the separation of alkanes and alkenes with a wider range of carbon atoms (C7-C12). This method is used to separate single-carbon olefins from C7-C12, improving deoxygenation efficiency and reducing energy consumption without significantly reducing the purity of the α-olefin product. The prepared α-olefins meet industrial requirements.

[0070] The preferred embodiments of this disclosure have been described in detail above with reference to the accompanying drawings. However, this disclosure is not limited to the specific details of the above embodiments. Within the scope of the technical concept of this disclosure, various simple modifications can be made to the technical solutions of this disclosure, and these simple modifications all fall within the protection scope of this disclosure.

[0071] It should also be noted that the various specific technical features described in the above embodiments can be combined in any suitable manner without contradiction. To avoid unnecessary repetition, this disclosure will not describe the various possible combinations separately.

[0072] Furthermore, various different embodiments of this disclosure can be combined in any way, as long as they do not violate the spirit of this disclosure, they should also be regarded as the content disclosed in this disclosure.

Claims

1. A method for separating α-olefins from Fischer-Tropsch light oil, characterized in that, The method includes the following steps: S1. The Fischer-Tropsch light oil is fractionated into C7-C9 and C10-C12 fractions. S2. The C7-C9 fraction and the C10-C12 fraction are subjected to oxygen-containing compound removal treatment by hollow fiber membrane extraction to obtain C7-C9 deoxygenated fraction and C10-C12 deoxygenated fraction; S3. The C7-C9 deoxygenated fraction and the C10-C12 deoxygenated fraction are respectively subjected to a second fractionation to obtain multiple deoxyalkene streams after fractionation. S4. The cut deoxyalkane stream is subjected to alkane adsorption and separation treatment using a sequential simulated moving bed.

2. The method according to claim 1, wherein, In step S2, the oxygen-containing compound removal treatment of the C7-C9 fraction using hollow fiber membrane extraction includes: filling the tubular side of the membrane module of the first hollow fiber membrane extractor with a first extractant, so that the membrane pores of the membrane module are filled with the first extractant; then injecting a first preparation solution containing the first extractant and a first back-extraction agent into the tubular side of the membrane module; then injecting the C7-C9 fraction into the shell side of the membrane module; and washing the outflow from the shell side outlet with water to obtain the deoxygenated C7-C9 fraction. The first extractant comprises an aqueous methanol solution, wherein the concentration of methanol is 50-90% by weight; the first stripping agent comprises water; and in the first preparation solution, the weight ratio of the first extractant to the first stripping agent is 2-10. The first hollow fiber membrane extractor has a packing factor of 0.1-0.6 and a mass flow rate ratio of 0.5-4 between the tube side and the shell side.

3. The method according to claim 1, wherein, In step S2, the oxygen-containing compound removal treatment of the C10-C12 fraction using hollow fiber membrane extraction includes: filling the tubular side of the membrane module of the second hollow fiber membrane extractor with a second extractant, so that the membrane pores of the membrane module are filled with the second extractant; then injecting a second preparation solution containing the second extractant and a second stripping agent into the tubular side of the membrane module; then injecting the C10-C12 fraction into the shell side of the membrane module; and washing the outflow from the shell side outlet with water to obtain the deoxygenated C10-C12 fraction. The second extractant comprises an aqueous ethanol solution, wherein the concentration of the ethanol is 50-85% by weight; the second stripping agent comprises water; and in the second preparation solution, the weight ratio of the second extractant to the second stripping agent is 1-3. The second hollow fiber membrane extractor has a packing factor of 0.1-0.6 and a mass flow rate ratio of 0.5-4 between the tube side and the shell side.

4. The method according to claim 1, wherein, The hollow fiber membrane is made of at least one of polysulfone, polyethersulfone, and polyvinylidene fluoride. The hollow fiber membrane has an inner diameter of 0.2-3.0 mm, an outer diameter of 0.3-4.0 mm, and a porosity of 30-80%.

5. The method according to claim 1, wherein, Step S4 includes: performing a first adsorption separation treatment on the first deoxyalkane stream obtained by cutting the C7-C9 deoxygenated fraction using a first sequential simulated moving bed; The second deoxyalkene stream obtained by cutting the C10-C12 deoxygenated fraction is subjected to a second adsorption separation treatment using a second sequential simulated moving bed.

6. The method according to claim 5, wherein, The number of adsorption columns in the first sequential simulated moving bed is 6-12, and the first adsorbent used includes one or more of CaNaA type molecular sieves, NaX type molecular sieves and NaY type molecular sieves. The first desorbent used includes a first desorbing component and a first solvent; the first desorbing component includes one or more C6 hydrocarbons, and the first solvent includes n-hexane and / or cyclohexane. The content of the first desorbing component is 5-60% by weight relative to the weight of the first desorbent.

7. The method according to claim 6, wherein, The operating temperature of the first sequential simulated moving bed is 20-40℃, the operating pressure is 0.1-0.4 MPa, and the mass flow ratio of the first deoxyalkane stream to the first desorbent is 0.1-5.

8. The method according to claim 5, wherein, The number of adsorption columns in the second sequential simulated moving bed is 8-16, and the second adsorbent used includes one or more of CaNaA type molecular sieves, NaX type molecular sieves and NaY type molecular sieves. The second desorbent used includes a second desorbing component and a second solvent; the second desorbing component includes one or more of C7 hydrocarbons and C8 hydrocarbons, and the second solvent includes n-heptane and / or n-octane; The content of the second desorbing component is 10-70% by weight relative to the weight of the second desorbent.

9. The method according to claim 8, wherein, The operating temperature of the second sequential simulated moving bed is 20-40℃, the operating pressure is 0.1-0.4 MPa, and the mass flow ratio of the second deoxyalkene stream to the second desorbent is 0.1-5.

10. The method according to claim 1, wherein, The Fischer-Tropsch light oil has an initial boiling point of 20-30°C, a final boiling point of 200-230°C, and an oxygen content of 0.1-10% by weight.

11. A system for the method of separating α-olefins from Fischer-Tropsch light oil according to any one of claims 1-10, characterized in that, The system includes a first fraction cutting unit, an oxygen-containing compound removal unit, a second fraction cutting unit, and an adsorption separation unit; The first fractionation unit is used to fractionate the Fischer-Tropsch light oil to obtain C7-C9 fractions and C10-C12 fractions. The oxygen-containing compound removal unit is used to remove oxygen-containing compounds from the C7-C9 fraction and the C10-C12 fraction to obtain the C7-C9 deoxygenated fraction and the C10-C12 deoxygenated fraction. The oxygen-containing compound removal unit includes a hollow fiber membrane extraction unit; The second fraction cutting unit is used to cut the C7-C9 deoxygenated fraction and the C10-C12 deoxygenated fraction to obtain multiple cut deoxyalkane streams; The adsorption separation unit includes a sequential simulated moving bed for adsorption and separation of the cut deoxyalkene stream.

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