Compositions for jet fuel and methods for blending

Abstract

The fuel composition comprises at least 85% by weight of isoparaffins and at least 3% by weight of C17+ hydrocarbons, the C17+ hydrocarbons having a pour point not exceeding -47°C, a final boiling point not exceeding 300°C, and a density of 730 kg / m³ to 772 kg / m³ at 15°C. Blended fuel compositions are also disclosed, wherein a synthetic fuel composition is blended with a natural fuel composition to provide Jet A-1 with a lower pour point. Methods for blending are also disclosed.

Description

Priority Statement

[0001] This application claims the benefit of U.S. Patent Application Serial No. 63 / 623,160, filed January 19, 2024, the entire disclosure of which is incorporated herein by reference. Technical Field

[0002] This field is aviation fuel. This field can specifically involve jet fuel produced from renewable feedstock materials. Background Technology

[0003] The ethanol-to-jet fuel process is one of the promising routes to minimize or eliminate net carbon combustion. The end products of this process are jet fuel and diesel fuel produced from bioethanol or methanol. Jet fuel is a sustainable aviation fuel designed to replace jet fuel produced from conventional sources such as crude oil.

[0004] Bioethanol can be produced through the fermentation of biological feedstocks. The bioethanol can then be dehydrated to produce ethylene. Methanol can be produced from carbon oxides and hydrogen. Hydrogen can be obtained through the electrolysis of water. Methanol can be converted into ethylene and propylene in the methanol-to-olefins (MTO) process.

[0005] Ethylene can be oligomerized into olefins, such as C4, C6, and C8 olefins. Olefin oligomerization is the process by which smaller olefins are oligomerized into larger olefins. More specifically, olefins can be converted into distillates, which include products ranging from jet fuels to diesel fuels. Olefin oligomer distillates can be hydrogenated for use as transportation fuels.

[0006] Jet fuel is one of the few petroleum fuels that cannot be easily replaced by electric motor systems because it requires high energy output to fuel aircraft, which electric motors cannot provide. Currently, some regions offer substantial incentives for green jet fuel to reduce the environmental impact of fossil-derived jet fuels.

[0007] Jet fuels contain n-alkanes, isoalkanes, cycloalkanes, and aromatic compounds. Typically, jet fuel molecules range from containing 8 to 16 carbon atoms.

[0008] Jet fuels made from synthetic hydrocarbons, also known as synthetic aviation fuels (SAF), must meet the basic requirements specified in ASTM D7566-20a. Typically, the T10 is limited to 205°C, and the final boiling point should not exceed 300°C. The flash point should not be lower than 38°C, and the density at 15°C should be 730 kg / m³. 3 With 772kg / m 3Between these limits, and for Jet A standard, the freezing point should not exceed -40°C, and for the more stringent Jet A-1 standard, the freezing point should not exceed -47°C. Synthetic aviation fuel (SAF) that meets ASTM D7566-20a can be blended with conventional fuel that meets ASTM D1655-18a (a standard commonly used for aviation fuels).

[0009] There is a great need for aviation fuel produced from synthetic hydrocarbons that meets standards applicable to the fuel industry. Summary of the Invention

[0010] We have formulated a composition of highly isomerized aviation fuel from renewable sources. The degree of isomerization allows alkanes with more than 16 carbon atoms to be included in the jet fuel range, resulting in higher jet fuel yields. Furthermore, the high degree of isomerization provides a composition with a very low pour point well below -47°C. Additionally, the aviation fuel can be blended with conventional jet fuels. Conventional jet fuels can have pour points below -47°C, therefore the blended jet fuel can provide a pour point below -47°C. Attached Figure Description

[0011] Figure 1 The graph shows the bulk modulus as a function of temperature.

[0012] Figure 2 The graph shows the bulk modulus as a function of pressure.

[0013] Figure 3 This is a schematic diagram of the oligomerization section of the method and apparatus of this disclosure.

[0014] Figure 4 This is a schematic diagram of the hydrogenation section of the method and apparatus of this disclosure.

[0015] definition

[0016] The term "connectivity" refers to the operative permission for fluid flow between enumerated components, which can be characterized as "fluid connectivity".

[0017] The term "downstream connectivity" means that in downstream connectivity, at least a portion of the fluid flowing toward the body can be operatively flowed from the object with which it is fluidly connected.

[0018] The term "upstream connectivity" means that at least a portion of the fluid flowing out of the main body can be operatively directed to an object in fluid communication with it.

[0019] The term "direct connection" means that fluid flow from an upstream component enters a downstream component without passing through any other intermediary container.

[0020] The term "indirect connection" refers to fluid flow from an upstream component entering a downstream component after passing through an intermediary container.

[0021] The term "bypass" means that an object is disconnected from the downstream entity at least within the scope of the bypass.

[0022] As used herein, the terms “major” or “most” mean greater than 50%, appropriately greater than 75%, and preferably greater than 90%.

[0023] The term "tower" refers to one or more distillation columns used to separate one or more components with different volatility. Unless otherwise specified, each column includes a condenser at the top of the column for condensing a portion of the overhead feed and returning it to the top of the column, and a reboiler at the bottom of the column for vaporizing a portion of the bottom feed and returning it to the bottom of the column. The feed to the column can be preheated. The top pressure is the pressure of the vapor at the top of the column at the vapor outlet. The bottom temperature is the liquid temperature at the bottom outlet. Top and bottom lines refer to the net lines from any downstream reflux or reboiler to the column. A stripping column may omit the reboiler at the bottom of the column, instead providing the heating requirements and separation power for a liquefied inert medium such as steam. Stripping columns typically feed from the top tray and remove the main product from the bottom.

[0024] As used herein, the term "separator" refers to a vessel having an inlet and at least one top vapor outlet and a bottom liquid outlet, and may also have an outlet for an aqueous feed stream from a boot. A flash tank is a type of separator that can be connected downstream of a separator that can operate at higher pressures. As used herein, the term "boiling point temperature" refers to the atmospheric equivalent boiling point (AEBP) calculated from the observed boiling temperature and distillation pressure, as provided in Appendix A7 of ASTM D1160, entitled "Practice for Converting Observed Vapor Temperatures to Atmospheric Equivalent Temperatures".

[0025] As used herein, the term “true boiling point” (TBP) refers to a test method conforming to ASTM D-2892 for determining the boiling point of a substance. ASTM D-2892 is used to produce standardized masses of liquefied gases, distillate fractions, and residues for which analytical data are available, and to determine the yield of the aforementioned fractions by both mass and volume, based on which a graph of distillation temperature versus mass% is obtained using fifteen theoretical plates in a column with a reflux ratio of 5:1.

[0026] As used herein, the terms “T10,” “T90,” or “T95” refer to the temperatures at which 10%, 90%, or 95% of a sample (as the case may be) boil using ASTM D-86 or TBP, respectively.

[0027] As used in this article, the term “initial boiling point” (IBP) refers to the temperature at which a sample begins to boil using ASTM D-7169, ASTM D-86, or TBP (as applicable).

[0028] As used in this article, the term “final boiling point” (FBP) refers to the temperature at which a sample, determined using ASTM D-7169, ASTM D-86, or TBP (as the case may be), reaches complete boiling.

[0029] As used herein, the term "diesel" means hydrocarbons that boil in the following ranges: IBP between 125°C (257°F) and 175°C (347°F), or T5 between 150°C (302°F) and 200°C (392°F), and "diesel fractionation point," including T95 between 343°C (650°F) and 399°C (750°F) using the TBP distillation method, or T90 between 280°C (536°F) and 340°C (644°F) using ASTM D-86. The term "green diesel" means diesel containing hydrocarbons not derived from fossil fuels.

[0030] As used herein, the term "jet fuel" refers to a hydrocarbon that boils within a T10 range of 190°C (374°F) to 215°C (419°F) and whose endpoint is between 290°C (554°F) and 310°C (590°F). The term "green jet fuel" refers to a jet fuel containing hydrocarbons that are not derived from fossil fuels.

[0031] As used herein, the term "Cx" should be understood to refer to a molecule having a number of carbon atoms represented by the variable "x". Similarly, the term "Cx-" refers to a molecule containing less than or equal to x, and preferably x and fewer carbon atoms. The term "Cx+" refers to a molecule having more than or equal to x, and preferably x and more carbon atoms. Detailed Implementation

[0032] Jet fuels contain n-alkanes, isoalkanes, cycloalkanes, and aromatic compounds. Typically, jet fuel molecules range from 8 to 16 carbon atoms. However, if the alkanes are highly isomerized, the jet fraction can include molecules with more than 16 carbon atoms while still meeting all fuel specifications. Oligomerization of ethylene molecules produces highly isomerized alkanes within the jet fuel range, allowing the inclusion of C17+ hydrocarbons in the jet fraction, thus maximizing jet fuel yield. After hydrogenation and fractionation, the jet products contain at least 85% by weight of isoalkanes, with lower concentrations of cycloalkanes and even lower concentrations of n-alkanes. Oligomerized fuels contain very small amounts (if any) of aromatics.

[0033] The olefins used in this method can be produced by a methanol-to-olefins process known as MTO. In an MTO reactor, methanol is converted into light olefin products. The methanol feedstock can be obtained by converting carbon dioxide into syngas using hydrogen from the electrolysis of water. Molecular sieves (such as microporous crystalline zeolite and non-zeolite catalysts, particularly aluminosilicate phosphate (SAPO)) facilitate the conversion of oxygen-containing compounds (such as methanol) into hydrocarbon mixtures, particularly those consisting primarily of light olefins. SAPO catalysts and their formulations are generally taught in US 4,499,327A, US 10,358,394, and US 10,384,986. The light olefins produced by the MTO process are concentrated in ethylene and propylene, but include C4 to C6 olefins.

[0034] A methanol feedstream is introduced into an MTO reactor and contacted with an MTO catalyst under MTO reaction conditions to convert methanol into olefins and water. The methanol feedstream may include methanol, dimethyl ether, ethanol, or a combination thereof. The MTO reaction conditions involve contact with a SAPO catalyst at pressures between 2 MPa and 3.8 MPa. The MTO reaction temperature should be between 325°C and 450°C. The weight hourly space velocity (“WHSV”) in the MTO reactor is maintained at 2 hours. -1 Up to 15 hours -1 Within the range. After the MTO reaction, the MTO catalyst is separated from the product olefin stream.

[0035] Ethanol can be produced by the fermentation of biomass. Ethanol produced by fermentation can be concentrated and dehydrated in a dehydration reactor to produce ethylene. Dehydration requires heating the ethanol feed stream to 400°C to 550°C and pressurizing it to 455 kPa (gauge pressure) 65 psig to 630 kPa (gauge pressure) (90 psig), and contacting it with a dehydration catalyst, such as an alumina-based catalyst. Dehydration produces ethylene and water, which must be separated to provide a dry ethylene feed stream.

[0036] Ethylene from any process can be oligomerized into fuel-range olefins and hydrogenated to produce synthetic aviation fuels conforming to ASTM D7566-20a. Synthetic materials boiled in the jet fuel range according to the standard have a T10 not exceeding 205°C and an FBP not exceeding 300°C. The T95 of the synthetic jet fuel composition may be at least 265°C. The jet fuel has a flash point of at least 38°C to meet the standard. The density of the jet fuel boiled material at 15°C is also 730 kg / m³. 3 Up to 772kg / m 3 This also conforms to SAF standards.

[0037] The resulting SAF composition contains at least 85% by weight of isoparaffins. Table 1 shows the range of isoparaffin distribution in SAFs produced by ethylene oligomerization in our pilot plant.

[0038] Table 1

[0039]

[0040] Total isoparaffins may be up to 90% by weight of the jet fuel, and preferably up to 91.5% by weight of the jet fuel. Cyclic alkanes may constitute at least 7% by weight of the SAF composition, while n-alkanes may constitute at least 1% by weight, suitably at least 2% by weight, and sometimes up to 7% by weight of the SAF composition. Aromatic compounds are substantially absent in the jet fuel composition, typically less than 0.1% by weight. The ratio of isoparaffins to n-alkanes is at least 80, suitably at least 35, typically at least 11, and readily at least 10. Less than 5% by weight of aromatic compounds are observed in the jet fuel boiling range, suitably not more than 1% by weight, and typically not more than 0.01% by weight of aromatic compounds are observed in the jet fuel boiling range of the compositions found.

[0041] The synthetic jet fuel material uniquely comprises monomethyl isomerized hydrocarbons and polymethyl isomerized hydrocarbons, meaning that the only alkyl group on the main hydrocarbon chain is methyl. In the SAF composition, at least 85% by weight of the alkanes are monomethyl and polymethyl isomerized hydrocarbons, preferably at least 90% by weight, and more preferably at least 92% by weight of the alkanes. Highly isomerized alkanes with a higher carbon number boil within the jet fuel range. Isoalkanes with 17 to 20 carbon atoms boil within the jet fuel range, resulting in a higher yield of jet fuel in the jet fraction.

[0042] Table 2 shows typical amounts of C17+ hydrocarbons and cyclic alkanes from the lower carbon number range of conventional jet fuels in the SAF of this disclosure produced in our test facility.

[0043] Table 2

[0044]

[0045] The synthetic jet fuel composition disclosed herein comprises at least 3% by weight of C17+ hydrocarbons, suitably at least 4% by weight of C17+ hydrocarbons, and preferably at least 5% by weight of C17+ hydrocarbons. The synthetic jet fuel composition comprises at least 3% by weight of C17 isoparaffins. The synthetic jet fuel comprises at least 3% by weight of C17+ alkanes, suitably at least 4% by weight of C17+ alkanes, and preferably at least 5% by weight of C17+ alkanes. The synthetic jet fuel composition comprises at least 3% by weight of C17 isoparaffins, suitably at least 4% by weight of C17+ isoparaffins, and preferably at least 5% by weight of C17+ isoparaffins. The synthetic jet fuel composition comprises at least 3% by weight of C17 isoparaffins. The jet fuel composition comprises at least 1% by weight of C18 isoparaffins. The jet fuel composition comprises at least 0.5% by weight of C19 isoparaffins. The jet fuel composition contains at least 0.05% by weight and typically at least 0.1% by weight of C20 isoparaffins.

[0046] The SAF disclosed herein has a low packing modulus. Figure 1 The graph shows the packing modulus as a function of temperature at different temperatures. Figure 2 The graph shows the bulk modulus as a function of pressure at different pressures. The bulk modulus of SAF is at least 2% lower than that of conventional Jet A fuel, and typically 2% to 10% lower.

[0047] ASTM 7566 requires synthetic jet fuels to have a pour point not exceeding -40°C to meet Jet A standards and not exceeding -47°C to meet the more stringent Jet A1 standards. Synthetic jet fuel compositions have been found to exhibit pour points below -70°C and even below -80°C. Conventional jet fuels from mineral sources must meet the specifications specified in ASTM D1655-18a, which requires a pour point of -40°C to qualify as Jet A fuels. If a conventional jet fuel has a pour point above -40°C or even below -47°C, we have found that conventional Jet A fuels can be blended with up to 50% by volume of SAF as disclosed in this disclosure to achieve a blended jet fuel with a pour point below -47°C, thus qualifying it as a premium Jet A-1 fuel and meeting ASTM D1655-18a standards.

[0048] A method for blending jet fuel components includes blending a first jet fuel component containing SAF and a second jet fuel component containing conventional jet fuel conforming to Jet A standards under ASTM D1655-18a. The SAF composition may contain at least 85% by weight of isoparaffins, suitably at least 90% by weight of isoparaffins, and preferably at least 92% by weight of isoparaffins. The SAF composition may exhibit a pour point not exceeding -47°C, typically not exceeding -70°C, and preferably not exceeding -80°C. The SAF may also exhibit a final boiling point not exceeding 300°C and a density of 730 kg / m³ to 772 kg / m³ at 15°C. The SAF in the first fuel component contains at least 3% by weight of C17+ hydrocarbons, suitably at least 4% by weight of C17+ hydrocarbons, and preferably at least 5% by weight of C17+ hydrocarbons. The SAF in the first fuel component contains at least 3% by weight of C17+ alkanes, suitably at least 4% by weight of C17+ alkanes, and preferably at least 5% by weight of C17+ alkanes. The SAF in the first fuel component contains at least 3% by weight of C17+ isoparaffins, suitably at least 4% by weight of C17+ isoparaffins, and preferably at least 5% by weight of C17+ isoparaffins. The second jet fuel component, which may contain Jet A fuel, may have a pour point above -47°C. The second jet fuel component may have at least 5% by volume of aromatics, suitably at least 10% by volume of aromatics, more suitably at least 15% by volume of aromatics, and typically at least 20% by volume of aromatics. The blended jet fuel has a pour point not exceeding -47°C, and preferably not exceeding -55°C, and is qualified as Jet A-1 fuel. The blend may contain at least 50% by volume of the second jet fuel component. The blend may have less than 5% by volume of aromatics.

[0049] Tables 3 and 4 show the properties of the jet fuel compositions of this disclosure prepared in our test facility, and how the compositions meet the applicable ASTM specifications in ASTM D1655 for conventional jet fuels or ASTM D7566 for synthetic jets. In Tables 3 and 4, “BD” indicates a blend of 50 / 50 SAF and conventional Jet A, and “NT” indicates 100% SAF. Table 3 shows a comparison of the fuels suitable for Jet A with ASTM D1655 and the blend compositions suitable for SAF with D7566.

[0050] Table 3

[0051]

[0052] Table 4 shows a comparison of the fuel with ASTM D4054.

[0053] Table 4

[0054]

[0055]

[0056]

[0057] Tables 3 and 4 show SAFs prepared according to this disclosure in our test facility that conform to ASTM D7566, and SAFs blended with up to 50% conventional Jet A fuel that conform to the Jet A-1 fuel standard in ASTM D1655. The blended jet fuel will have at least 50% by volume of conventional Jet A fuel, but will be considered as 100% by volume of Jet A1 fuel because it has a pour point not exceeding -47°C and meets other requirements for Jet A1 fuel in ASTM D1655.

[0058] The blended jet fuel composition comprises at least 5% by weight, suitably at least 10% by weight, and preferably at least 15% by weight, aromatic compounds derived from conventional jet fuel components. The blended jet fuel composition comprises at least 25% by weight, suitably at least 30% by weight, more suitably at least 35% by weight, preferably at least 40% by weight, and more preferably at least 45% by weight, isoparaffins derived from SAF components. Of the isoparaffins in the blended jet fuel, at least 25% by weight, suitably at least 30% by weight, more suitably at least 35% by weight, preferably at least 40% by weight, and more preferably at least 45% by weight, are mono- or polymonomethyl alkanes. The blended jet fuel composition has a pour point not exceeding -47°C, qualifying it as a Jet A-1 grade. The blend composition has a T10 not exceeding 205°C, an FBP not exceeding 300°C, and a density of 775 kg / m³ to 840 kg / m³ at 15°C, conforming to ASTM D1655-18a for Jet A-1 premium fuels. The blended jet fuel contains at least 0.5% by weight of C17+ hydrocarbons, suitably at least 1% by weight of C17+ hydrocarbons, preferably at least 1.5% by weight of C17+ hydrocarbons, and more preferably at least 2% by weight of C17+ hydrocarbons. The blended jet fuel contains at least 0.5% by weight of C17+ alkanes, suitably at least 1% by weight of C17+ alkanes, preferably at least 1.5% by weight of C17+ alkanes, and more preferably at least 2% by weight of C17+ alkanes. The blended jet fuel has at least 0.5% by weight of C17+ isoparaffins, suitably at least 1% by weight of C17+ isoparaffins, preferably at least 1.5% by weight of C17+ isoparaffins, and more preferably at least 2% by weight of C17+ isoparaffins.

[0059] The methods and apparatus may include Figure 3 Low-polymer segment 230 and Figure 4 The hydrogenation section 300 in the middle. First, go to... Figure 3 The low-polymer segment 230 will come from Figure 2 The compressed olefin feed stream in line 210 is supplied to the oligomerization section 230. The olefin feed stream may contain a significant amount of ethylene. The fed olefin feed stream may primarily contain ethylene. In one aspect, the olefin feed stream may contain at least 95 mol% ethylene. The olefin feed stream in line 210 may be referred to as the ethylene feed stream. As previously mentioned, the olefin feed stream may be provided by dehydration of ethanol or from the MTO unit. The temperature of the olefin feed stream may be from 60°C (140°F) to 150°C (302°F), preferably from 80°C (176°F) to 100°C (212°F), and the pressure may be from 3.5 MPa (500 psig) to 8.4 MPa (1200 psig).

[0060] Go to Figure 1 The oligomerization section 10 feeds the initial vaporized olefin stream containing C2 olefins from line 1 to the initial separator 6. The initial vaporized olefin stream is separated in the initial separator to provide the vaporized olefin stream in line 3 and the liquefied olefin stream in line 4. The initial separator can operate at temperatures from 16°C (60℉) to 38°C (100℉) and pressures from 2.1 MPa (gauge pressure) (300 psig) to 2.8 MPa (gauge pressure) (400 psig). The vaporized olefin stream in line 3 can be compressed in compressor 8 to the oligomerization pressure in line 12.

[0061] A preliminary liquid olefin stream containing C3+ olefins is supplied to the oligomerization section 10 from line 2. This preliminary olefin stream may contain C3-C8 olefins. The preliminary olefin stream from line 2 may be combined with a liquefied olefin stream (which may contain C3-C8 olefins) from line 4 and fed into the liquid feed buffer tank 7. A liquid olefin stream containing C3-C8 liquid olefins from line 5 of the liquid feed buffer tank 7 may be combined with a compressed vaporized olefin stream from line 11 to supply the vaporized olefin stream in line 12.

[0062] The vapor and liquid olefin feed streams may contain significant amounts of ethylene and propylene. The vapor and liquid olefin feed streams may primarily contain ethylene and / or propylene. In one aspect, the vapor and liquid olefin feed streams may contain at least 95 mol% ethylene and / or propylene. The vapor and liquid olefin feed streams in lines 5 and 12 may be referred to as light olefin feed streams. Additional olefinic substances with a carbon number range of C4 to C8 are contemplated in the feed streams. The light olefin feed streams may be provided by dehydration of ethanol or from the MTO unit. The temperature of the light olefin feed streams may be from 20°C (68℉) to 150°C (302℉), and the pressure may be from 2.16 MPa (350 psig), preferably from 3.5 MPa (500 psig) to 8.4 MPa (1200 psig).

[0063] The light olefin feed stream can initially be contacted with a first-stage oligomerization catalyst to oligomerize ethylene and propylene into oligomers, and then contacted with a second oligomerization catalyst to oligomerize the unconverted ethylene and propylene from the first-stage oligomerization.

[0064] Oligopolymerization reactions generate a significant amount of exothermic heat. For example, the dimerization of ethylene can generate 612 kcal / kg (1100 BTU / lb) of heat. Therefore, this significant exothermic heat must be managed. Consequently, the light olefin feed streams in lines 5 and 12 can be separately diverted into multiple olefin feed streams. Figure 1 In this process, the light olefin feed stream is split into two separate streams. The compressed vapor olefin feed stream in line 12 can be split into a first vapor olefin feed stream in line 12a and a second vapor olefin feed stream in line 12b. The liquid olefin feed stream in line 5 is split into a first liquid olefin feed stream in line 5a and a second liquid olefin feed stream in line 5b. The first vapor olefin feed stream in line 12a is mixed with the first liquid olefin feed stream in line 5a to provide a first feed olefin feed stream in the first feed olefin line 13a. The second vapor olefin feed stream in line 12a is mixed with the second liquid olefin feed stream in line 5b to provide a second feed olefin feed stream in the second feed olefin line 13b. More or fewer independent olefin feed streams can be used. Up to six feed olefin feed streams are readily conceivable.

[0065] The compressed vapor olefin feed stream in line 12 can be split into multiple olefin feed streams of equal proportion in lines 12a and 12b. The liquid olefin feed stream in line 5 can be split into multiple olefin feed streams of equal proportion in lines 5a and 5b. Alternatively, the compressed vapor olefin feed stream in line 12 can be split into unequal streams. Similarly, the liquid olefin feed stream in line 5 can be split into unequal streams. For example, one or both of the vapor olefin feed stream or the liquid olefin feed stream can be split into streams with decreasing flow rates, wherein the feed olefin feed stream to the preceding reactor has a greater flow rate than the feed olefin feed stream to the following reactor. In one embodiment, both the vapor olefin feed stream and the liquid olefin feed stream can be split into two streams of equal flow rates, each stream containing 50% by volume of the feed olefin feed stream.

[0066] In another embodiment, the first feed olefin stream in the first feed olefin line 13a may account for 70% to 90% of the total feed olefin stream, and the second feed olefin stream in the second olefin line 13b may account for 10% to 30% of the total feed olefin stream. In another embodiment, each of the feed streams is split in different proportions. For example, the liquid olefin stream in line 5 may be split into two streams such that the first liquid stream in line 5a will account for 70% to 90% of the total liquid olefin stream in line 5, and the second liquid stream in line 5b will account for 10% to 30% of the total liquid olefin stream, while the compressed vapor stream in line 12 will be equally split into two streams such that the first vapor olefin stream in line 12a and the second olefin stream in line 12b each account for 50% of the total compressed vapor olefin stream.

[0067] To manage the exothermic reaction, the feed olefin stream can be diluted with a diluent stream to provide a diluted olefin stream to absorb the exothermic reaction. The diluent stream may contain the alkane stream from diluent line 14. The diluent stream from diluent line 14 may be added to the first feed olefin stream in the first feed olefin line 13a before it is fed into the first stage oligomerization reactor 22. Preferably, after the feed olefin streams in lines 5 and 12 are split into multiple olefin streams, the diluent stream is added to the first feed olefin stream in line 13a to provide the first diluted olefin feed stream in line 16a, thus allowing the diluent stream to pass through the entire first stage oligomerization reaction. Alternatively, the diluent stream may also be split into multiple streams, with each diluent stream added to one or more of the corresponding feed olefin streams. The mass flow rate of the diluent stream can be 2 to 8 times, and preferably 3 to 6 times, the combined mass flow rate of the first feed olefin stream in the first feed olefin line 13a and the second feed olefin stream in the second feed olefin line 13b.

[0068] The recycle olefin stream containing C4 to C8 olefins in recycle line 26 can be mixed with the feed olefin stream and oligomerized in the first-stage oligomerization reactor 22. In one embodiment, the recycle olefin stream in line 26 is split into multiple recycle olefin streams 26a-26d. The recycle olefin stream in the first recycle olefin line 26a can be mixed with the first feed olefin stream in line 13a and fed into the first-stage oligomerization reactor 22. In another embodiment, the first recycle olefin stream in the first recycle olefin line 26a is mixed with the first feed olefin stream in line 13a and the diluent stream in line 14 to provide the diluted first feed olefin stream in line 16a.

[0069] The first diluted olefin feed stream may contain no more than 50% by weight of olefins, suitably no more than 30% by weight of olefins, and preferably no more than 20% by weight of olefins. In one embodiment, the first diluted olefin feed stream contains 10% to 35% by weight of C2 to C8 olefins. The first diluted olefin feed stream may contain no more than 50% by weight of ethylene, suitably no more than 25% by weight of ethylene, and preferably no more than 20% by weight of ethylene. In one embodiment, the first diluted olefin feed stream contains 10% to 20% by weight of propylene. The first diluted olefin feed stream may contain no more than 50% by weight of propylene, suitably no more than 25% by weight of propylene, and preferably no more than 20% by weight of propylene. In one embodiment, the first diluted olefin feed stream contains 10% to 20% by weight of propylene.

[0070] The first-stage oligomerization reactor 22 may include a series of first-stage oligomerization catalyst beds 22a, 22b, 22c, and 22d, each for charging the olefin feed stream. The first-stage oligomerization reactor 22 preferably contains four fixed first-stage oligomerization catalyst beds 22a, 22b, 22c, and 22d. It is also envisioned that each of the first-stage oligomerization catalyst beds 22a, 22b, 22c, and 22d can be in a dedicated first-stage oligomerization reactor, or that multiple first-stage oligomerization catalyst beds can be in two or more independent first-stage oligomerization reaction vessels. Up to six first-stage oligomerization catalyst beds are readily conceivable. Figure 1 In this process, two first-stage oligomerization reaction vessels, 21a and 21b, are used.

[0071] When the first-stage oligomer reactor 22 has been passivated, it can be used in parallel with other first-stage oligomer reactors, during which time the first-stage oligomer reactor 22 is regenerated in situ by burning coke from the catalyst. In another embodiment, each first-stage oligomer reactor may include a pre-reactor, a lag reactor, and a backup reactor to facilitate regeneration. Figure 1 Only two reaction vessels, 21a and 21b, are shown in the image.

[0072] The diluted first feed olefin stream in line 16a can be cooled in a first feed cooler 18a to provide a cooled diluted first feed olefin stream in line 20a and to the first bed 22a of the first-stage oligomerization catalyst in the first first-stage oligomerization reaction vessel 21a of the first-stage oligomerization reactor 22. The cooled diluted first feed olefin stream in line 20a can be fed at a temperature of 180°C (356℉) to 260°C (500℉) and a pressure of 3.5 MPag (500 psig) to 8.4 MPag (1200 psig). Feed cooler 18a may include a steam generator.

[0073] The diluted first feed olefin stream is preferably fed into the first first-stage catalyst bed 22a in a downward flow operation. However, an upward flow operation may be suitable. The diluted first feed olefin stream is in a mixed vapor-liquid phase, wherein the vapor phase mainly comprises ethylene. Exothermic reactions occur when ethylene, propylene, and recycled olefins oligomerize in the first first-stage oligomerization catalyst bed 22a due to the highly exothermic nature of the olefin oligomerization reaction. Despite cooling and dilution, the oligomerization of the first feed olefin stream produces a first oligomer effluent stream in the first oligomer effluent line 24a at an elevated outlet temperature. The elevated outlet temperature is limited to between 150°C (302℉) and 260°C (500℉).

[0074] The second feed olefin stream in line 13b may be mixed with the second recycle olefin stream in the second recycle olefin line 26b and with the first oligomer effluent stream in the first oligomer effluent line 24a removed from the first first-stage oligomer catalyst bed 22a in the first first-stage reactor 21a to provide the mixed second feed olefin stream in line 16b. The first oligomer effluent stream in line 24a includes a diluent stream from diluent line 14, which is added to the first feed olefin stream in line 13a. The second feed olefin stream may contain no more than 35% by weight of C2 to C8 olefins, suitably no more than 25% by weight of C2 to C8 olefins, and preferably no more than 20% by weight of C2 to C8 olefins. The second diluted olefin stream may contain no more than 30% by weight of ethylene, suitably no more than 25% by weight of ethylene, and preferably no more than 20% by weight of ethylene. The second diluted olefin feed stream may contain no more than 30% by weight of propylene, suitably no more than 25% by weight of propylene, and preferably no more than 20% by weight of propylene. The second mixed feed olefin feed stream in line 16b may be cooled in a second feed cooler 18b located outside the first first-stage oligomerizing reactor 21a to provide a cooled second feed olefin feed stream in line 20b and to be loaded into the second bed 22b of the first-stage oligomerizing catalyst in the first first-stage oligomerizing reactor 21a. The feed cooler 18b may include a steam generator.

[0075] The second cooled feed olefin stream in line 20b can be charged at temperatures ranging from 180°C (356℉) to 230°C (446℉) and pressures ranging from 3.5 MPag (500 psig) to 8.4 MPag (1200 psig). The second cooled feed olefin stream will contain a diluent and olefins from the first oligomer stream. The diluted second feed olefin stream is in a mixed vapor-liquid phase, where the vapor phase primarily contains ethylene. The olefins from the first oligomer stream will oligomerize in the second first-stage catalyst bed 22b. The oligomerization of ethylene, propylene, recycled olefins, and oligomers in the second olefin stream in the second first-stage oligomer catalyst bed 22b produces a second oligomeric olefin effluent stream in the second oligomer effluent line 24b at an elevated outlet temperature. The elevated outlet temperature can be limited to between 30°C (54℉) and 50°C (90℉), which is higher than the inlet temperature to catalyst bed 22b.

[0076] The second oligomer effluent stream removed from the second first-stage oligomer catalyst bed 22b in the first first-stage reaction vessel 21a can be mixed with the third recycle olefin stream in the third recycle olefin line 26c to provide the first recycle olefin feed stream in line 16c. In one embodiment, neither the first feed olefin stream in line 13a nor the second feed olefin stream in line 13b is directly added to the first recycle olefin feed stream in line 16c. Alternatively, a portion of the feed olefin streams in lines 13a and 13b can be fed together with the second oligomer effluent stream and the first recycle olefin feed stream in line 16c. The second oligomer effluent stream in line 24b includes a diluent stream from diluent line 14, which is added to the first feed olefin stream in line 13a. The first recycle olefin feed stream in line 16c may contain no more than 30% by weight of ethylene, suitably no more than 25% by weight of ethylene, and preferably no more than 20% by weight of ethylene. The first recycle olefin feed stream may contain no more than 30% by weight of propylene, suitably no more than 25% by weight of propylene, and preferably no more than 20% by weight of propylene. The first recycle olefin feed stream in line 16c may contain no more than 30% by weight of C2 to C8 olefins, suitably no more than 25% by weight of C2 to C8 olefins, and preferably no more than 20% by weight of C2 to C8 olefins. The first recycle olefin feed stream in line 16c may be cooled in a third feed cooler 18c located outside the oligomerization reactor 22 to provide a cooled first recycle olefin feed stream in line 20c and fed into a third bed 22c of the first-stage oligomerization catalyst in the first-stage oligomerization reactor 22. In one embodiment, the third bed 22c of the first-stage oligomerization catalyst is disposed in a second first-stage oligomerization reaction vessel 21b. The feed cooler 18c may include a steam generator.

[0077] The cooled first recycle olefin feed stream in line 20c can be charged at temperatures from 180°C (356℉) to 230°C (446℉) and pressures from 3.5 MPag (500 psig) to 8.4 MPag (1200 psig). The first recycle olefin feed stream will include a diluent and olefins from the second and third recycle olefin feed streams. The olefins will be oligomerized in the third catalyst bed 22c. The oligomerization of ethylene and propylene in the third bed 22c of the first-stage oligomerizing catalyst, as well as the oligomerization of the oligomers in the first recycle olefin feed stream, produces a third oligomer effluent stream in the third oligomer effluent line 24c at an elevated outlet temperature. In one embodiment, the third oligomer effluent stream is the penultimate oligomerizing olefin feed stream, and the third oligomer effluent line 24c is the penultimate oligomer effluent line 24c. The increased outlet temperature is limited to between 30°C (54°F) and 50°C (90°F), which is higher than the inlet temperature to catalyst bed 22c.

[0078] The third oligomer effluent stream removed from the second first-stage oligomerization reactor 21b of the first-stage oligomerization reactor 22 in line 24c may be mixed with the fourth recycled olefin stream in line 26d to provide the second recycled olefin feed stream in line 16d. The third oligomer effluent stream in line 24c includes a diluent stream from diluent line 14, which is added to the first olefin stream in line 13a. The feed olefin streams in lines 13a and 13b are not directly added to the second recycled olefin feed stream in line 16d. In one embodiment, the third oligomer effluent stream in line 24c may also be mixed with and oligomerize with a portion of the feed olefin streams in lines 13a and 13b. The second recycled olefin feed stream may contain no more than 35% by weight of C2 to C8 olefins, suitably no more than 30% by weight of C2 to C8 olefins, and preferably no more than 25% by weight of C2 to C8 olefins. The second recycled olefin feed stream may contain no more than 30% by weight of ethylene, suitably no more than 25% by weight of ethylene, and preferably no more than 20% by weight of ethylene. The second recycled olefin feed stream may contain no more than 30% by weight of propylene, suitably no more than 25% by weight of propylene, and preferably no more than 20% by weight of propylene. The second recycled olefin feed stream in line 16d may be cooled in a fourth feed cooler 18d located outside the second vessel 21b of the first-stage oligomerizing reactor 22 to provide the cooled second recycled olefin feed stream in line 20d and to be fed into the fourth bed 22d of the first-stage oligomerizing catalyst in the second vessel of the first-stage oligomerizing reactor 22. The feed cooler 18d may include a steam generator.

[0079] The cooled second recycle olefin feed stream in line 20d can be charged at temperatures ranging from 180°C (356℉) to 230°C (446℉) and pressures ranging from 3.5 MPa (gauge pressure) (500 psig) to 8.4 MPa (gauge pressure) (1200 psig). The cooled second recycle olefin feed stream in line 20d will include a diluent and olefins from the third or penultimate oligomer effluent stream, as well as C4-C8 olefins from the fourth recycle olefin feed stream. The olefins will oligomerize on the fourth catalyst bed 22d. The oligomerization of ethylene and propylene in the second recycle olefin feed stream in the fourth bed 22d of the first-stage oligomer catalyst produces a fourth oligomer feed stream in the fourth oligomer effluent line 24d at an elevated outlet temperature. The elevated outlet temperature is limited to between 30°C (54℉) and 50°C (90℉), which is higher than the inlet temperature to catalyst bed 22d.

[0080] The fourth oligomer effluent stream in line 24d exits the second reaction vessel 21b of the first-stage oligomer reactor 22. In one embodiment, the fourth oligomer effluent stream in line 24d is the last oligomer effluent stream, and the fourth oligomer effluent line 24d is the last oligomer effluent line 24d.

[0081] The first-order oligomerization reaction mainly occurs in the liquid phase or in a mixture of liquid and gas phases based on olefins at a rate of 0.5 hr. -1 up to 10 hours -1 WHSV occurs. We have found that, throughout the first-stage oligomerization catalyst bed, typically 10% to 50% by weight of ethylene in the olefin stream is converted to higher olefins. Ethylene will initially dimerize to butene via the catalyst. The majority of propylene and butene in the olefin stream charged into the first-stage oligomerization catalyst bed are oligomerized. In one embodiment, at least 99 mol% of propylene and butene in the olefin stream are oligomerized.

[0082] The first-stage oligomerization catalyst may include a zeolite catalyst. The first-stage oligomerization catalyst can be considered a solid acid catalyst. The zeolite may comprise between 5% and 95% by weight of the catalyst, for example, between 5% and 85% by weight. Suitable zeolites include those with a structure having one of the following categories: MFI, MEL, ITH, IMF, TUN, FER, BEA, FAU, BPH, MEI, MSE, MWW, UZM-8, MOR, OFF, MTW, TON, MTT, AFO, ATO, and AEL. Three-letter codes indicating zeolite classes are as defined by the Structure Commission of the International Zeolite Association and are maintained at http: / / www.iza-structure.org / databases. UZM-8 is described in U.S. Patent No. 6,756,030. In a preferred aspect, the first-stage oligomerization catalyst may comprise a zeolite with a framework having a decacyclic porous structure. Examples of suitable zeolites with a decacyclic pore structure include TON, MTT, MFI, MEL, AFO, AEL, EUO, and FER. In a further preferred aspect, the first-stage oligomerizing catalyst comprising a zeolite with a decacyclic pore structure may comprise a one-dimensional pore structure. A one-dimensional pore structure indicates a zeolite containing non-intersecting pores substantially parallel to one of the crystal axes. The pores preferably extend through the zeolite crystal. Suitable examples of zeolites with a decacyclic one-dimensional pore structure may include MTT. In another aspect, the first-stage oligomerizing catalyst comprises MTT zeolite.

[0083] The first-stage oligomerizing catalyst can be formed by combining zeolite with a binder and then forming the catalyst into granules. The granules may optionally be treated with a phosphorus reagent to produce zeolite with a phosphorus component ranging from 0.5% to 15% by weight of the treated catalyst. The binder is used to impart hardness and strength to the catalyst. Binders include alumina, aluminum phosphate, silica, silica-alumina, zirconium oxide, titanium dioxide, and combinations of these metal oxides, as well as other refractory oxides and clays such as montmorillonite, kaolin, palygorskite, chlorite, and chlorite-palygorskite. Preferred binders are aluminum-based binders, such as alumina, aluminum phosphate, silica-alumina, and clay.

[0084] One component of the catalyst binder used in this invention is alumina. The alumina source can be any of various hydrated alumina oxides or alumina gels, such as α-alumina monohydrate with a boehmite or boehmite structure, α-alumina trihydrate with a trihydrate structure, β-alumina trihydrate with a bayonet structure, etc. Suitable alumina can be purchased under the trade name VERSAL from UOP LLC. Preferred alumina can be purchased under the trade name CATAPAL from Sasol North American Alumina Products Group. This material is a very high purity α-alumina monohydrate (boehmite), which, upon calcination at high temperatures, has shown the production of high purity γ-alumina.

[0085] A suitable first-stage oligomerizing catalyst is prepared by mixing zeolite and alumina in proportional volume to achieve a desired zeolite to alumina ratio. In one embodiment, the MTT content can be from 5% to 85% by weight, for example, 20% to 82% by weight of MTT zeolite, and the balance of alumina powder will provide a suitably supported catalyst. Silica supports are also considered.

[0086] A monobasic acid, such as nitric acid or formic acid, can be added to the mixture in the aqueous solution to dissolve the alumina gel in the binder. Additional water can be added to the mixture to provide sufficient moisture to form a dough of sufficient consistency for extrusion or spray drying. Extrusion aids, such as cellulose ether powder, can also be added. A preferred extrusion aid is available under the trade name Methocel from Dow Chemical Company.

[0087] The paste or dough can be prepared in granular form, preferably by extruding the dough through a die with openings of the desired size and shape, followed by crushing the extruded material into extrudates of the desired length and drying. Further calcination steps may be employed to impart increased strength to the extrudates. Typically, calcination is carried out in an air stream at temperatures ranging from 260°C (500℉) to 815°C (1500℉). The MTT catalyst is not selectively treated (e.g., with amines) to neutralize acid sites.

[0088] The extruded particles can have any suitable cross-sectional shape, i.e., symmetrical or asymmetrical, but most commonly have a symmetrical cross-sectional shape, preferably spherical, cylindrical, or multi-lobed. The cross-sectional diameter of the particles can be as small as 40 μm; however, it is typically from 0.635 mm (0.25 inch) to 12.7 mm (0.5 inch), preferably from 0.79 mm (1 / 32 inch) to 6.35 mm (0.25 inch), and most preferably from 0.06 mm (1 / 24 inch) to 4.23 mm (1 / 6 inch).

[0089] In one exemplary embodiment, an MTT-type zeolite catalyst is provided in a catalyst bed or more in the first-stage oligomer reactor 22, on a high-purity pseudoboehmite alumina substrate, at a ratio of 90 / 10 to 20 / 80, and preferably between 20 / 80 and 50 / 50.

[0090] The first-stage oligomerizing catalyst can be regenerated upon deactivation. Suitable regeneration conditions include subjecting the first-stage oligomerizing catalyst to hot air at, for example, 400°C to 500°C in situ. To facilitate regeneration without interrupting operation, a swirling bed arrangement can be used in conjunction with an alternative first-stage oligomerizing reactor. A regeneration gas stream can be introduced into the first-stage oligomerizing reactor 22 requiring regeneration. The regeneration gas may include air with an increased or decreased oxygen concentration. The activity and selectivity of the regenerated catalyst are comparable to those of the fresh catalyst.

[0091] Zeolite catalysts are advantageous as first-stage oligomerization catalysts. Zeolite catalysts exhibit relatively low sensitivity to oxygen contamination. Therefore, if produced via an ethanol dehydration process, the olefin feed in line 1 requires only a minimal degree of oxygen-containing compound removal.

[0092] Compared to the light olefin feed streams in lines 5 and 12, the last first-stage oligomer feed stream in the final first-stage oligomer effluent line 24d has an increased concentration of ethylene and propylene oligomers. The last first-stage oligomer feed stream in the final first-stage oligomer effluent line 24d is generated by steam in steam generator 18e or cooled by other heat exchange, and further cooled by heat exchange with the second-stage oligomer feed stream in line 34, and possibly further cooled by, for example, by a steam generator to provide feed first-stage oligomer feed stream, and fed into the second-stage oligomer reactor 32 in the second-stage oligomer feed line 28. To achieve the most desired olefin product, the second-stage oligomer reactor 32 is operated at a temperature of 80°C (176℉) to 200°C (392℉). The second-stage oligomerization reactor 32 operates at pressures ranging from 2.1 MPa (300 psig) to 7.6 MPa (1100 psig), and more preferably from 3.5 MPa (500 psig) to 6.9 MPa (1000 psig). The second-stage oligomerization feed stream oligomerizes in a mixed vapor-liquid phase, primarily as C4+ olefins.

[0093] The second-stage oligomerization reactor 32 may be downstream connected to the first-stage oligomerization reactor 22. The second-stage oligomerization reactor 32 is preferably operated in a downflow operation. However, an upflow operation may be suitable. The second-stage oligomerization feed stream contacts the second-stage oligomerization catalyst, causing the unconverted ethylene from the first-stage oligomerization reactor 22 to dimerize and trimerize, while higher olefins also dimerize, trimerize, and tetramerize to provide olefins with a distillate range. Regarding the second-stage oligomerization reactor 32, process conditions can be selected to produce a higher percentage of jet-range olefins, which, when hydrogenated in subsequent steps described below, produce the desired jet-range hydrocarbon products. The majority of the unconverted ethylene from the first-stage oligomerization reactor 22 is dimerized, trimerized, and tetramerized. In one embodiment, at least 99% by weight of the ethylene in the second-stage oligomerization feed stream is predominantly converted to butene.

[0094] The second-stage oligomer reactor 32 may include a first reaction vessel 31a and a second reaction vessel 31b. The first reaction vessel includes a first bed 32a of the second-stage oligomer catalyst, and the second reaction vessel includes a second bed 32b of the second-stage oligomer catalyst. A first second-stage oligomer feed stream exits from the first second-stage reaction vessel 31a, is cooled, and is loaded into the second second-stage reaction vessel 31b. A second-stage oligomer feed stream with an increased average carbon number in line 28 that is higher than the fed second-stage oligomer feed stream exits the second-stage oligomer reactor 32 in line 34.

[0095] The first-stage oligomer reactor 22 and the second-stage oligomer reactor 32 can utilize vapor-liquid distribution trays to mix and disperse ethylene vapor with liquid olefins and liquid alkanes to facilitate heat transfer and manage exothermic reactions.

[0096] The second-stage oligomerization catalyst is preferably an amorphous silica-alumina base material containing metals from Group VIII and / or Group VIB of the periodic table using Chemical Abstracts Service (CISA) symbols. In one aspect, the catalyst contains a Group VIII metal promoted by a Group VIB metal. Typically, silica and alumina are only present in the base material, thus the silica / alumina ratio of the catalyst is the same as that of the base material. The metal can be impregnated onto the silica-alumina base material or ion-exchanged with it. Co-milling is also considered. The catalyst used in this invention can have a low-temperature acidity ratio of at least 0.15, suitably 0.2, and preferably greater than 0.25, as determined by ammonia temperature-programmed desorption (ammonia TPD) as described below. Additionally, suitable catalysts will have a pH between 50 and 400 μm. 2 The surface area between / g, as determined by nitrogen BET.

[0097] The preferred second-stage oligomerizing catalyst comprises an amorphous silica-alumina support. One component of the catalyst support used in this invention is alumina. The alumina can be any of various hydrated aluminum oxides or alumina gels, such as α-alumina monohydrate with boehmite or boehmite structure, α-alumina trihydrate with trihydrate structure, β-alumina trihydrate with bayerite structure, etc. Particularly preferred alumina can be purchased under the trade name CATAPAL from Sasol North America Alumina Product Group. This material is a very high-purity α-alumina monohydrate (boehmite), which has shown to produce high-purity γ-alumina after calcination at high temperatures. The other component of the catalyst support is amorphous silica-alumina. Suitable silica-alumina with a silica-to-alumina ratio of 2.6 is purchased from CCIC (a subsidiary of JGC) in Japan.

[0098] Another component used in the preparation of the second-stage oligomer catalyst of this invention is a surfactant. Preferably, the surfactant is mixed with the aforementioned alumina and silica alumina powders. The resulting mixture of surfactant, alumina, and silica alumina is then formed as described below, dried, and calcined. Calcination effectively removes the organic components of the surfactant by combustion, but only after the surfactant has faithfully performed its function according to the invention. Any suitable surfactant can be used according to the invention. Preferred surfactants are those selected from a range of commercial surfactants sold by Solvay SA under the trade name "Antarox". "Antarox" surfactants are generally characterized as modified linear aliphatic polyethers and are low-foaming, biodegradable detergents and wetting agents.

[0099] A suitable silica-alumina mixture is prepared by mixing silica-alumina and alumina in proportional volume to achieve a desired silica to alumina ratio. In one embodiment, 75% to 99% by weight of amorphous silica-alumina and 10% to 20% by weight of alumina powder will provide a suitable carrier, wherein the silica to alumina ratio is 2.6. In one embodiment, other ratios of amorphous silica-alumina to alumina may be suitable.

[0100] The surfactant can be combined with the mixture of silica alumina and alumina using any convenient method. Preferably, the surfactant is mixed during the mixing and formation of the alumina and silica alumina. A preferred method is to mix an aqueous solution of the surfactant with the blend of alumina and silica alumina prior to the final formation of the carrier. Preferably, the surfactant is present in the paste or dough in an amount from 0.01% to 10% by weight, based on the weight of the alumina and silica alumina.

[0101] Monobasic acids such as nitric acid or formic acid can be added to the mixture in the aqueous solution to dissolve the alumina glue in the binder. Additional water can be added to the mixture to provide sufficient moisture, thus forming a dough with a sufficient consistency to be extruded or spray-dried.

[0102] The paste or dough can be prepared in granular form. A preferred method is to extrude a dough mixture of alumina, silica alumina, surfactant, and water through a die having openings of the desired size and shape, followed by breaking the extruded material into extrudates of the desired length and drying them. A further calcination step may be employed to impart increased strength to the extrudates. Typically, calcination is carried out in a dry air stream at temperatures ranging from 260°C (500℉) to 815°C (1500℉).

[0103] The extruded particles can have any suitable cross-sectional shape, i.e., symmetrical or asymmetrical, but most commonly have a symmetrical cross-sectional shape, preferably spherical, cylindrical, or multi-lobed. The cross-sectional diameter of the particles can be as small as 40 μm; however, it is typically from 0.635 mm (0.25 inch) to 12.7 mm (0.5 inch), preferably from 0.79 mm (1 / 32 inch) to 6.35 mm (0.25 inch), and most preferably from 0.06 mm (1 / 24 inch) to 4.23 mm (1 / 6 inch).

[0104] The amorphous silica-alumina supports used in this paper are typically characterized by large total pore volume, average pore size, and surface area, providing ample space and area for depositing active metal components. As measured by conventional mercury porosimetry, the total pore volume of the support is typically from 0.2 cc / g to 2.0 cc / g, preferably from 0.25 cc / g to 1.0 cc / g, and most preferably from 0.3 cc / g to 0.9 cc / g. Typically, the pore volume of the support in pores with a diameter greater than 100 angstroms is less than 0.1 cc / g, preferably less than 0.08 cc / g, and most preferably less than 0.05 cc / g. As measured by the BET method, the surface area is typically greater than 50 m². 2 / gram, for example, above 200m 2 / gram, preferably at least 250m 2 / gram, and the optimal value is 300m 2up to 400m 2 / gram.

[0105] To prepare the second-stage oligomer catalyst, a support material is compounded with one or more precursors of at least one metal component from Group VIII or VIB of the periodic table (e.g., by single or multiple impregnation of calcined amorphous refractory oxide support particles). The Group VIII metal (preferably nickel) should be present at a concentration of 0.5% to 15% by weight, and the Group VIB metal (preferably tungsten) should be present at a concentration of 0% to 12% by weight. Impregnation can be achieved by any method known in the art (e.g., by spray impregnation), wherein a solution containing the metal precursor in dissolved form is sprayed onto the support particles. Another method is a multi-dip procedure, in which the support material is repeatedly contacted with the impregnation solution, with or without intermittent drying. Other methods involve immersing the support in a large volume of impregnation solution or circulating the support therein, and yet another method is a pore volume or pore saturation technique, in which the support particles are introduced into an impregnation solution of a volume just sufficient to fill the pores of the support. Sometimes, the pore saturation technique can be modified to utilize an impregnation solution having a volume that is 10% to 10% smaller than the volume that just fills the pore.

[0106] If the active metal precursors are combined by impregnation, subsequent or second calcination at elevated temperatures (e.g., between 399°C (750°F) and 760°C (1400°F)) converts the metals into their respective oxide forms. In some cases, calcination can be performed after each impregnation of the individual active metals. Subsequent calcination produces a catalyst containing the active metals in their respective oxide forms.

[0107] The preferred second-stage oligomer catalyst of the present invention comprises an amorphous silica-alumina matrix impregnated with 0.5 wt% to 15 wt% nickel, the matrix being in the form of a 3.175 mm (0.125 inch) extrusion with a density of 0.45 g / ml to 0.65 g / ml. It is also envisioned that the metal can be incorporated onto the support by other methods, such as ion exchange and co-milling.

[0108] The second-stage oligomerization catalyst can be regenerated upon deactivation. Suitable regeneration conditions include subjecting the catalyst, for example, to hot air at 400°C to 500°C in situ. To facilitate regeneration without interrupting operation, a swirling bed arrangement can be employed in conjunction with an alternative second-stage oligomerization reactor. The regeneration gas may include air with an increased or decreased oxygen concentration. The activity and selectivity of the regenerated catalyst are comparable to those of the fresh catalyst.

[0109] The second-stage oligomerization reaction is also essentially exothermic. The final oligoolefin feed stream in line 24d includes a diluent feed stream from diluent line 14, which is added to the first feed olefin feed stream in the first feed olefin line 13a and carried through the first-stage oligomerization catalyst bed 22a to 22d. The diluent feed stream is then transported in line 28 to the second-stage oligomerization reactor 32 to absorb the exothermic reaction in the second-stage oligomerization reactor. A dedicated diluent line to the second-stage oligomerization reactor 32 is also envisioned for rapidly controlling the exothermic rise or cooling the second-stage oligomerization reactor 32.

[0110] When the oligomerization reaction is carried out according to the above process conditions, a C4 olefin conversion rate of ≥95% or ≥97% is achieved. The second-stage oligomer stream obtained in line 34 includes various olefin products, which are hydrocarbons in the distillate range.

[0111] Compared to the second-stage oligomer feed stream in line 28, the oligoolefin feed stream in line 34, with an increased C8+ olefin concentration, exchanges heat with the first-stage oligomeric feed stream in line 24d, resulting in a pressure reduction. It then exchanges heat with the bottom feed stream from the olefin splitter in line 30 and is fed into the dealkylation tower 40. The temperature of the oligoolefin feed stream in line 34 is between 160°C (320℉) and 190°C (374℉), and the pressure is between 3.9 MPa (gauge pressure) (550 psig) and 7 MPa (gauge pressure) (1000 psig).

[0112] We have discovered that light alkanes, such as ethane and / or propane, are produced in the first-stage oligomerization reactor 22 and / or the second-stage oligomerization reactor 32. These light alkanes must be removed from the second-stage oligomerization feed stream used for fuel production, particularly to facilitate the recycling of light olefins back to the first-stage oligomerization reactor 22. Light alkanes are inert and will accumulate in the recycling loop. Therefore, the second-stage oligomerization feed stream in line 34 is dealkane-de- ...

[0113] In dealkane column 40, light alkanes (such as C3 hydrocarbons and suitably C2 hydrocarbons) may be separated in the light alkane overhead stream in overhead line 42 from the dealkane bottom stream containing C4+ hydrocarbons and suitably C3+ hydrocarbons in bottom line 44. Olefins may be recycled from the dealkane overhead stream in overhead line 42 to the first-stage oligomerization reactor 22. Dealkane column 40 can be operated at a bottom temperature of 177°C (350℉) to 302°C (575℉) and a top pressure of 207 kPa (gauge pressure) (30 psig) to 690 kPa (gauge pressure) (100 psig) if operated as a deethanizer. The dealkation column 40 can operate at a bottom temperature of 194°C (381°F) to 333°C (630°F) and a top pressure of 207 kPa (gauge pressure) (30 psig) to 1.38 MPa (gauge pressure) (200 psig) if operated as a dealkation column.

[0114] The light alkane overhead stream in overhead line 42 can be cooled and separated in dealkane tower receiver 46 to provide a dealkane waste gas stream in waste gas line 47, where it can be cooled and fed to further processing, such as being taken out as fuel gas in line 48 along with the net vapor stream in receiver tower overhead line 68. The condensate from dealkane tower receiver 46 can be refluxed back to dealkane tower 40 in dealkane tower overhead liquid line 49. The dealkane waste gas stream can be used as fuel to provide heating load in method 10. In one embodiment, some condensate from dealkane tower receiver 44 in line 49 can be carried as olefin recycle in line 51 to the first-stage oligomer reactor in lines 72 and 26.

[0115] The dealkane feed stream in the bottom line 44 may be split between a reboiler stream in line 50 and a net bottom stream in line 54. The reboiler stream is reboiled by heat exchange with a first hot diesel stream in line 52, which may be taken from the jet fuel fractionator bottom heat exchange stream in jet fuel bottom heat exchange line 74. The net bottom stream may be fed directly to the olefins splitter column 60 without heating. The reboiled bottom stream in line 50 may be boiled back to the dealkane column 40 to provide heating. In another embodiment, the feed to the dealkane column 40 is not preheated by the olefins splitter bottom stream in line 30, but the feed to the olefins splitter column 60 in the net bottom line 54 will be preheated by heat exchange with the olefins splitter bottom stream in line 64.

[0116] The dealkane feed stream in the net bottom line 54 of the dealkane tower is fractionated in the olefin splitter tower 60 into a light olefin feed stream that may be in the top line 62 of the olefin splitter tower and a heavy olefin feed stream that may be in the bottom line 64 of the olefin splitter tower. Olefins can be recycled from the top line 62 of the olefin splitter tower to the first-stage oligomerization reactor 22. The top line of the olefin splitter tower can be cooled to 19°C (66℉) to 93°C (200℉), and a portion of the resulting condensate is refluxed back to the olefin splitter tower 60 from the receiver 66. The net vapor stream from the receiver 66 in the receiver tower top line 68 can be further processed, such as fuel gas in line 48 and waste gas stream in waste gas line 47. The light olefin condensate from the bottom of the olefin splitter receiver in line 70 can be split between the reflux stream returning to the tower in line 71 and the light olefin recirculation stream in recirculation line 72, which can be recirculated to the first-stage oligomerization reactor 22 or alternatively to the second-stage oligomerization reactor 32. The light olefin stream in line 72 can constitute 1% to 15% by weight, or possibly the majority, of the light olefin stream in line 70. The light olefin stream in line 72 can contain 40% to 80% by weight of C4-C8 olefins. In one embodiment, the light olefin feed stream in line 72 can be flashed in separator 75 to remove vapor from the light olefin vapor feed stream, which can be conveyed to the hydrogenation section in top line 77, and the liquid recirculated olefin oligomer feed stream in line 26 can be recycled to the first-stage oligomerization reactor 22 to oligomerize C4-C8 olefins.

[0117] The heavy olefin stream in the bottom line 64 of the splitter tower can be split between the reboiler stream in the reboiler reboiler line 65 and the reboiler stream reboiled by heat exchange with a second hot diesel stream in line 73, which may be taken from... Figure 2 The jet fuel fractionator bottom heat exchange feed stream in jet tower bottom heat exchange line 74 is fed back to olefins splitter tower 60. The cooled second hot diesel feed stream in line 112 returns... Figure 2 The hydrogenation section 110, after being reboiled, is returned to the jet fractionation tower 100. The heavy olefin stream in the net bottom line 30 is cooled by heat exchange with the second-stage oligomer stream in line 34, and then transported to the hydrogenation section (not shown). Purge material from the heavy olefin stream can be removed in line 33. The heavy olefin stream contains C9+ olefins, which, once cooled, can be transported to the hydrogenation section 110.

[0118] Go to Figure 2 Hydrogenation section 110 in the middle, from Figure 1The heavy olefin stream containing C9+ oligoolefins in the bottom line 30 of the net olefins splitter can be hydrogenated in hydrogenation reactor 80 to saturate the olefin bonds and provide fuel. This step is performed to ensure that the product fuel meets or exceeds the thermal oxidation requirements specified in ASTM D7566-20 for ethanol to jet-synthetic alkane kerosene (ATJ-SPK). Additionally, saturating the oligomeric heavy olefins will provide an alkane stream that can be used as a diluent stream in line 14. The heavy olefin stream in line 30 can be cooled to generate vapor and can be combined with... Figure 1 The light olefin liquid feed stream containing C2 to C8 olefins in line 77 is combined to produce the combined olefin feed stream in line 79. The combined olefin feed stream in line 79 can also be combined with the hydrogen feed stream in line 76 to provide the combined hydrogenation feed stream in line 81, which is cooled and charged into the hydrogenation reactor 80 at 125°C (257℉) to 204°C (400℉) and 2.8 MPa (400 psig) to 6.9 MPa (1000 psig). An excess of hydrogen, such as 1.5 to 5.0 stoichiometric amounts, can be used to ensure complete saturation.

[0119] Hydrogenation is typically carried out using conventional hydrogenation or hydrotreating catalysts and may include metal catalysts containing, for example, palladium, rhodium, nickel, ruthenium, platinum, rhenium, cobalt, molybdenum, or combinations thereof, and their supported forms. The catalyst support can be any solid, inert material, including but not limited to oxides such as silica, alumina, titanium dioxide, calcium carbonate, barium sulfate, and carbon. The catalyst support can be in the form of powder, granules, or pellets.

[0120] In an exemplary embodiment, hydrogenation is performed in a hydrogenation reactor 80, which comprises an alumina-supported platinum catalyst, for example, 0.1 wt% to 2 wt%, preferably 0.5 wt% to 0.9 wt%. In another embodiment, the hydrogenation catalyst comprises 5 wt% to 30 wt% of a nickel catalyst. The hydrogenation reactor 80 converts olefins into alkane products having the same carbon number distribution as the olefins, thereby forming distillate-range alkanes suitable for use as jet fuel and diesel fuel.

[0121] The hydrogenated heavy feed stream discharged from the hydrogenation reactor 80 in line 83 can be separated in a thermal separator 82 that provides hydrocarbon splitting. In the thermal separator 82, the hydrogenated heavy feed stream is separated into a hot hydrogenated vapor stream in the top line 84 and a hot hydrogenated liquid stream in the bottom line 86. The diluent stream is then recycled to… Figure 1Before the first-stage oligomer reactor 22, the hydrogenated heavy liquid stream in the bottom line 86 can be heated by heat exchange with the diluent stream in line 14. The heated hydrogenated heavy liquid stream in the bottom line 86 can be fed into the stripping tower 90. The thermal separator can operate at temperatures from 204°C (400℉) to 343°C (650℉) and pressures from 2.8 MPa (400 psig) to 6.9 MPa (1000 psig).

[0122] The hot hydrogenated vapor stream in the top line 84 of the hot tower can be cooled and fed to the cold separator 88. The cold separator separates the cooled hot hydrogenated vapor stream in the top line 84 into a cold vaporized hydrogenated stream in the top line 87 and a cold heavy hydrogenated liquid stream in the bottom line 89. The purge stream in the purge line 85 can be taken from the cold vaporized hydrogenated stream in the top line 87, and the remainder can be compressed and combined with the makeup hydrogen in line 88 to provide the hydrogen stream in line 76. The cold hydrogenated heavy liquid stream in the bottom line 89 can be fed into the stripping tower 90 at a feed position higher than the feed position of the hot hydrogenated heavy liquid stream in the bottom line 86 of the hot separator. The cold separator can operate at temperatures from 32°C (90°F) to 71°C (150°F) and pressures from 2.8 MPa (400 psig) to 4.5 MPa (650 psig).

[0123] Stripper 90 may be a flash stripper to remove light gases from the hot hydrogenated liquid stream in hot bottom line 86 and the cold hydrogenated liquid stream in cold bottom line 89. These two streams may be combined and charged into stripper 90, or they may be charged separately as shown. Stripper 90 removes residual light gases from the liquid hydrogenated stream to provide the stripper overhead stream in stripper overhead line 92 and the stripper bottom stream in stripper bottom line 94. The stripper overhead stream in stripper overhead line 92 is cooled and separated in stripper receiver 96 to provide the stripper exhaust stream in stripper receiver overhead line 97 and the condensate stream returned to the tower in line 98. The stripping tower 90 can operate at bottom temperatures of 232°C (450°F) to 316°C (600°F) and top pressures of 207 kPa (30 psig) to 689 kPa (100 psig).

[0124] After stripping to remove volatiles in stripper 90, the stripped fuel stream in stripper bottom line 94 can be fed into jet fractionation tower 100 without further heating. Alternatively, stripper 90 may be omitted upstream of jet fractionation tower 100. In jet fractionation tower 100, the stripped fuel stream can be separated into jet exhaust gas stream in top line 102, green jet stream in side line 104 from one side of jet fractionation tower 100, and green diesel stream in bottom line 106. Jet fractionation tower 100 can operate at a bottom temperature of 288°C (550℉) to 400°C (750℉) and a top pressure of 35 kPa (5 psig) to 350 kPa (50 psig).

[0125] The jet fractionation overhead stream in the overhead line 102 can be cooled, and a portion of the resulting condensate is refluxed back to the jet fractionation tower 100 from the jet fractionation receiver 108 in the jet fractionation overhead liquid line 109, while the remaining condensate is treated as naphtha-range product stream in line 119. Additionally, a net exhaust stream containing C4 to C8 hydrocarbons is drawn from the jet fractionation receiver 108 into the receiver tower overhead line 105. The net exhaust stream in the receiver tower overhead line 105 contains mostly lighter hydrocarbons and can be used to fuel the reboiler heater 116 of the jet fractionation tower 100.

[0126] The green jet stream extracted from side line 104 contains kerosene-range C8-C18 hydrocarbons and can be cooled and extracted as a jet fuel product meeting applicable SPK standards. In an alternative embodiment, instead of returning all condensate to the tower, the green jet stream can be extracted from the condensate stream in line 109 from jet fractionation receiver 108. This green jet stream extracted from line 109 must then be further stripped to remove light fractions. In such embodiments, side line 104 will not be extracted to recover the green jet stream.

[0127] The green diesel bottom feed stream in bottom pipeline 106 can be diverted into the reboiled diesel feed stream in pipeline 107 and the diesel product feed stream in pipeline 114. The reboiled diesel feed stream in pipeline 107 can be diverted into the jet tower bottom heat exchange feed stream 74 and the bypass bottom feed stream in bottom bypass pipeline 111. For example... Figure 1 As shown, the jet bottom heat exchange feed in jet tower bottom heat exchange line 74 can be split into a first hot diesel feed in line 52 and a second hot diesel feed in line 73 to provide reboiling heat to the dealkane tower 40 and olefin splitter 60, respectively. The bypass bottom feed in line 111, through its valve, or the cooled second hot diesel feed in line 112, or some of both, is taken out in jet reboiling line 113, reboiled in flame heater 116, and fed back to jet fractionation tower 100.

[0128] The diesel product stream in pipeline 114 is split into a diesel product stream in diesel product pipeline 118 and a diluent stream in pipeline 14. The diluent stream in pipeline 14 can be cooled by heat exchange with the hot hydrogenated heavy liquid stream in the bottom pipeline 86 of the thermal separator, and is then recycled back to the... Figure 1 The feed olefin streams (preferably the first feed olefin stream in line 13a) in the oligomerization section 10 are mixed to provide the first diluted olefin feed stream in line 16a to absorb the exothermic reaction in the oligomerization reactor 22. The green diesel stream in diluent line 14 is alkane-based and therefore inert to the oligomerization and hydrogenation reactions it may undergo. The diesel product stream in diesel product line 118 can be cooled and fed into the diesel pool.

[0129] Starting with ethylene and / or propylene, the disclosed method can efficiently produce SAF that meets applicable ASTM D7566-20a for use as jet fuel. It can be blended with Jet A fuel that meets ASTM D1655-18a requirements to provide Jet A-1 premium fuel.

[0130] Specific implementation plan

[0131] While the following description is presented in conjunction with specific embodiments, it should be understood that the description is intended to illustrate, and not limit, the scope of the foregoing description and the appended claims.

[0132] A first embodiment of this disclosure is a composition having at least 85% by weight of isoparaffins and at least 3% by weight of C17+ hydrocarbons, the C17+ hydrocarbons having a freezing point not exceeding -47°C, a final boiling point not exceeding 300°C, and a density of 730 kg / m³ to 772 kg / m³ at 15°C. An embodiment of this disclosure is one, any, or all of the embodiments described in the preceding embodiments to the first embodiment of this paragraph, further having at least 4% by weight of C17+ hydrocarbons. An embodiment of this disclosure is one, any, or all of the embodiments described in the preceding embodiments to the first embodiment of this paragraph, further having at least 5% by weight of C17+ hydrocarbons. An embodiment of this disclosure is one, any, or all of the embodiments described in the preceding embodiments to the first embodiment of this paragraph, further having a flash point of at least 38°C. An embodiment of this disclosure is one, any, or all of the embodiments described in the preceding embodiments to the first embodiment of this paragraph, further having a T10 not exceeding 205°C. One embodiment of this disclosure is one, any, or all of the embodiments described in the preceding embodiments to the first embodiment of this paragraph, wherein the embodiment further has a freezing point not exceeding -70°C. One embodiment of this disclosure is one, any, or all of the embodiments described in the preceding embodiments to the first embodiment of this paragraph, wherein the embodiment further has a T95 of at least 265°C. One embodiment of this disclosure is one, any, or all of the embodiments described in the preceding embodiments to the first embodiment of this paragraph, wherein the embodiment further has a ratio of at least 10 isoparaffins to n-paraffins. One embodiment of this disclosure is one, any, or all of the embodiments described in the preceding embodiments to the first embodiment of this paragraph, wherein the embodiment further has at least 85% by weight of alkanes, which are monomethyl isoparaffins and polymethyl isoparaffins. One embodiment of this disclosure is one, any, or all of the embodiments described in the preceding embodiments to the first embodiment of this paragraph, wherein the embodiment further has at least 3% by weight of C17 alkanes. One embodiment of this disclosure is one, any, or all of the embodiments described in the preceding embodiments of this paragraph to the first embodiment of this paragraph, wherein the embodiment further comprises at least 1% by weight of C18 alkanes. One embodiment of this disclosure is one, any, or all of the embodiments described in the preceding embodiments of this paragraph to the first embodiment of this paragraph, wherein the embodiment further comprises at least 0.3% by weight of C19 alkanes. One embodiment of this disclosure is one, any, or all of the embodiments described in the preceding embodiments of this paragraph to the first embodiment of this paragraph, wherein the embodiment further comprises no more than 5% by volume of aromatic compounds.One embodiment of this disclosure is one, any, or all of the embodiments described in the preceding embodiments to the first embodiment of this paragraph, wherein the embodiment is blended with conventional jet fuel having a pour point not exceeding -40°C, wherein the resulting blend comprises at least 50% by volume of conventional jet fuel and has a pour point not exceeding -47°C. Another embodiment of this disclosure is one, any, or all of the embodiments described in the preceding embodiments to the first embodiment of this paragraph, wherein the embodiment has a bulk modulus at least 2% lower than that of conventional Jet A aircraft turbine fuel.

[0133] A second embodiment of this disclosure is a composition having at least 5% by weight of an aromatic compound, at least 30% by weight of an isoparaffin and at least 2% by weight of a C17+ hydrocarbon having a freezing point not exceeding -47°C, a final boiling point not exceeding 300°C, and a density of 775 kg / m³ to 840 kg / m³ at 15°C.

[0134] A third embodiment of this disclosure is a method for blending jet fuel, comprising blending a first jet fuel component with a second jet fuel component, the first jet fuel component having a composition comprising at least 85% by weight of isoparaffins and at least 3% by weight of C17+ hydrocarbons, the C17+ hydrocarbons having a pour point not exceeding -47°C, a final boiling point not exceeding 300°C, and a density of 730 kg / m³ to 772 kg / m³ at 15°C, the second jet fuel component having a composition comprising a pour point exceeding -47°C, wherein the blend has a pour point not exceeding -47°C. One embodiment of this disclosure is one, any one, or all of the embodiments described above to the third embodiment in this paragraph, wherein the blend comprises at least 50% by volume of the second jet fuel component. One embodiment of this disclosure is one, any one, or all of the embodiments described above to the third embodiment in this paragraph, wherein the second jet fuel component has at least 5% by volume of aromatic compounds, but the first jet fuel composition has less than 5% by volume of aromatic compounds.

[0135] Although no further detailed description has been provided, it is believed that those skilled in the art can make full use of this disclosure by employing the foregoing description and can readily determine the essential characteristics of this disclosure without departing from the spirit and scope of the invention, and can make various changes and modifications to this disclosure to suit various uses and conditions. Therefore, the foregoing preferred specific embodiments should be understood as illustrative only and not as limiting the remainder of this disclosure in any way, and are intended to cover various modifications and equivalent arrangements included within the scope of the appended claims.

[0136] In the foregoing, all temperatures are expressed in degrees Celsius, and all portions and percentages are by weight unless otherwise specified.

Claims

1. A fuel composition comprising: at least about 85% by weight of isoparaffin and at least about 3% by weight of C17+ hydrocarbon, said C17+ hydrocarbon having a freezing point not exceeding about -47°C, a final boiling point not exceeding about 300°C, and a density of about 730 kg / m³ to about 772 kg / m³ at 15°C.

2. The composition according to claim 1, wherein the composition further comprises at least about 4% by weight of C17+ hydrocarbons.

3. The composition according to claim 1, wherein the composition further comprises at least about 5% by weight of C17+ hydrocarbons.

4. The composition according to claim 1, wherein the composition further has a flash point of at least about 38°C.

5. The composition according to claim 1, wherein the composition further has a T10 temperature not exceeding about 205°C.

6. The composition according to claim 1, wherein the composition further has a freezing point not exceeding about -70°C.

7. The composition according to claim 1, wherein the composition further has a T95 of at least 265°C.

8. The composition according to claim 1, wherein the composition further comprises at least a ratio of isoparaffin to n-paraffin of 10.

9. The composition according to claim 1, wherein the composition further comprises at least 85% by weight of an alkane, said alkane being a monomethyl isoalkane or a polymethyl isoalkane.

10. A method for blending jet fuel, the method comprising: A first jet fuel component is blended with a second jet fuel component, the first jet fuel component having a composition comprising at least about 85% by weight of isoparaffins and at least about 3% by weight of C17+ hydrocarbons having a pour point not exceeding about -47°C, a final boiling point not exceeding about 300°C, and a density of about 730 kg / m³ to about 772 kg / m³ at 15°C, the second jet fuel component having a composition having a pour point exceeding -47°C, wherein the blend has a pour point not exceeding -47°C.

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