Process for producing jet fuel with thermal integration

Abstract

The process allows for the separation of the liquid hydrocracking stream from the liquid hydrogen isomerization stream, and thus the preservation of heat in the hydrocracking stream. Preserving heat in the hydrocracking stream avoids the need to reheat the hydrocracking stream before product fractional distillation. In particular, kerosene in the hydrocracking stream is not cooled in the hydrogen isomerization stream and is then reheated in fractional distillation for distillation of kerosene-range hydrocarbons from diesel-range hydrocarbons.

Description

[Technical Field]

[0001] (Priority statement) This application claims priority to U.S. Provisional Patent Application No. 63 / 418,020, filed on October 20, 2022, which is incorporated in its entirety herein.

[0002] (Field of Invention) This field involves producing hydrocarbons useful as aviation fuels from hydrocarbon supply materials. In particular, this field may relate to producing hydrocarbons useful as aviation fuels from biorenewable supply materials such as triglycerides and free fatty acids found in materials such as plant and animal fats and oils. [Background technology]

[0003] As global demand for fuel increases, there is growing interest in producing fuel from sources other than crude oil and blending its components. These sources, often referred to as bio-renewable sources, include, but are not limited to, vegetable oils such as corn, rapeseed, canola, and soybean; microbial oils such as algal oil; animal fats such as non-edible animal fats; fish oil; and various waste flows such as yellow and brown grease and sewage sludge. A common characteristic of these sources is that they are composed of glycerides and free fatty acids (FFAs). Both triglycerides and FFAs contain aliphatic carbon chains with 8 to 24 carbon atoms. The aliphatic carbon chains in triglycerides or FFAs can be fully saturated, or monovalent, divalent, or polyunsaturated.

[0004] Hydroprocessing may involve the process of converting hydrocarbons into more valuable products in the presence of a hydroprocessing catalyst and hydrogen. Hydrotreating or hydrogenation is a process of contacting hydrocarbons with hydrogen in the presence of a hydroprocessing catalyst that is active in removing heteroatoms such as sulfur, nitrogen, oxygen, and metals from hydrocarbon feedstock. Hydroprocessing can saturate hydrocarbons with double and triple bonds, such as olefins.

[0005] The production of hydrocarbon products within the diesel boiling point range can be achieved by hydrogenating biorenewable feedstock. Hydrogenated hydrocarbons can be deoxygenated, decarboxylated, and / or decarbonylated by hydrogenating biorenewable feedstock. Decarboxylation and decarbonylation remove carbon from paraffin molecules, while deoxygenation does not. Hydrogen isomerization may be performed after hydrogenating to improve the low-temperature flow properties of the diesel and jet fuel products. Hydrogen isomerization, or dewaxing by hydrogenation, is a hydrogenation process that improves the low-temperature flow properties of hydrocarbons by increasing alkyl branching on the hydrocarbon skeleton in the presence of hydrogen and a hydrogen isomerization catalyst. Hydrogen isomerization, as used herein, includes dewaxing by hydrogenation.

[0006] Hydrocracking is a hydrogenation process in which hydrocarbons are broken down into lower molecular weight hydrocarbons in the presence of hydrogen and a hydrocracking catalyst. Depending on the desired yield, a hydrocracking unit may contain one or more beds of the same or different catalysts.

[0007] When producing jet fuel from triglycerides (also known as "fat"), some degree of hydrocracking and isomerization is required to meet the jet fuel standards outlined in ASTM D7566 Annex 2 and ASTM D1655. These key standards required for jet fuel in D7566 include a freezing point below -40°C (ASTM D5972, D7153, or D7154) and a fuel density of 772 kg / m³. 3The following are the requirements for jet fuel: density (ASTM D1298 or D4052), T10 below 205°C (ASTM D86), and final boiling point (FBP) below 300°C (ASTM D86). Larger molecules that do not meet these jet fuel standards are primarily hydrocracking to meet them, which in turn results in lower yields in the production process and undesirable low-energy-density fuel. Aviation fuel is valued for its high energy per unit volume.

[0008] Carbon intensity is a term that refers to the number of moles of carbon dioxide produced to produce one mole of fuel. Combustion of hydrocarbons, such as heating a hydrocarbon flow in a hydrogenation unit, increases carbon intensity. It is desirable to provide renewable fuels from processes that reduce carbon intensity by reducing heating requirements. [Overview of the project]

[0009] The inventors have found that preheating of all reactor charge flows can be achieved by heat exchange with the hydrogenated effluent flow from the hydrogenation reactor. As a result, combustion charge heaters can be omitted, leading to reduced capital and operating costs, as well as reduced carbon intensity due to the absence of combustion of hydrocarbons that provide enthalpy. [Brief explanation of the drawing]

[0010] [Figure 1] This is a schematic process flow diagram of this disclosure.

[0011] definition The term "communication" means that the flow of material between the listed components is operablely permitted.

[0012] The term "downstream communication" means that in downstream communication, at least a portion of the material flowing to the object can be operably flowed from the communicating object.

[0013] The term "upstream communication" means that in upstream communication, at least a portion of the material flowing from an object can flow operably into the communicating object.

[0014] The term "direct connection" means that the flow from the upstream component enters the downstream component without passing through a fractionation or transformation unit and undergoing compositional changes due to physical fractionation or chemical transformation.

[0015] The term "indirect communication" means that a flow from an upstream component enters a downstream component after passing through a fractionation or transformation unit and undergoing compositional changes due to physical fractionation or chemical transformation.

[0016] The term "bypass" means that an object is removed from downstream communication with the object it is bypassing, at least to the extent that it is bypassing it.

[0017] The term “column” refers to a distillation column (singular or plural) for separating one or more different volatile components. Unless otherwise specified, each column includes a condenser at the top of the column to condense and reflux a portion of the top flow returning to the top of the column, and a reboiler at the bottom of the column to vaporize a portion of the bottom flow and return it to the bottom of the column. The feedstock to the column may be preheated. The top pressure is the pressure of the top vapor at the vapor outlet of the column. The bottom temperature is the temperature at the liquid bottom outlet. The top line and bottom line refer to the net lines from column to column downstream of any reflux or reboil. A stripper column omits the reboiler at the bottom of the column and instead may provide heating requirements and separation propulsion from a fluidizing inert medium such as vapor. A stripping column typically feeds the feedstock to a top tray and removes the main product from the bottom.

[0018] As used herein, the term “component-rich flow” means that the rich flow exiting the container has a higher concentration of components than the feed into the container.

[0019] As used herein, the term "component lean stream" means that the lean stream exiting the vessel has a lower concentration of components than the feedstock supplied to the vessel.

[0020] As used herein, the term "boiling point temperature" means the atmospheric equivalent boiling point (AEBP) calculated from the observed boiling point and distillation pressure using the equations provided in ASTM D86 or ASTM D2887.

[0021] As used herein, the term "True Boiling Point" (TBP) means a test method for determining the boiling point of a substance. This test method is for producing liquefied gas, distillation fractions, and residual oil of standardized quality from which analytical data can be obtained, and for determining the yields of the above fractions by both mass and volume, where a graph of temperature versus mass percent distilled is generated using 15 theoretical stages in a column with a reflux ratio of 5:1, corresponding to ASTM D-2892.

[0022] As used herein, the terms "T5" or "T95" mean the temperature at which 5 mass percent or 95 mass percent, respectively, of a sample boils, using ASTM D-86 or TBP as appropriate.

[0023] As used herein, the term "initial boiling point" (IBP) means the temperature at which a sample begins to boil, using ASTM D2887, ASTM D-86, or TBP as appropriate.

[0024] As used herein, the term "final boiling point" (FBP) means the temperature at which a sample has completely evaporated, using ASTM D2887, ASTM D-86, or TBP as appropriate.

[0025] As used herein, the term “diesel boiling point range” means that a hydrocarbon boils within the range of “diesel cut points,” which include IBP at 125°C (257°F) to 175°C (347°F), or T5 at 150°C (302°F) to 200°C (392°F), and T95 at 343°C (650°F) to 399°C (750°F), using the TBP distillation method.

[0026] As used herein, the term "diesel conversion" means converting a feed that boils above the diesel cutoff point to a substance that boils below the diesel cutoff point, within the diesel boiling point range.

[0027] As used herein, the term “separator” means a vessel having an inlet and at least a top vapor outlet and a bottom liquid outlet, and which may also have an aqueous outlet from the boot. A flash drum is a type of separator that can be downstream-communicated with a separator that can operate at higher pressures.

[0028] As used herein, the terms “dominant” or “dominant” mean more than 50%, preferably more than 75%, and more preferably more than 90%.

[0029] When used herein, "C x The term should be understood to refer to a molecule having a number of carbon atoms represented by the subscript "x". Similarly, "C x The term "-" refers to a molecule containing x or fewer carbon atoms, preferably x and less. x The term "+" refers to a molecule having x or more (more than or equal to x), preferably x and more (x and more) carbon atoms.

[0030] As used herein, the term "carbon number" refers to the number of carbon atoms in a hydrocarbon molecule, typically a paraffin molecule. [Modes for carrying out the invention]

[0031] The inventors have found that all the heating requirements for charges in diesel and jet fuel production processes can be provided by hydrogenation effluent. In particular, hydrogenation deoxygenation of oxygenated hydrocarbons from glyceride molecules obtained from renewable bio-renewable feedstocks provides sufficient heating requirements for charge flows for hydrogenation reactors, hydrogen isomerization reactors, and even hydrocracking reactors (if used).

[0032] Figure 1 shows a process 10 for processing a hydrocarbon feedstock according to an exemplary embodiment. Preferably, the hydrocarbon feedstock is a biorenewable hydrocarbon feedstock. A feed line 12 transports a hydrocarbon stream of fresh, preferably biorenewable, feedstock to a feed surge drum 14. The biorenewable feedstock may be blended with a mineral feedstock, but preferably constitutes the majority or all of the feedstock. The mineral feedstock is a conventional feedstock derived from crude oil extracted from the ground. The biorenewable feedstock may contain nitrogen concentrations of 1 wppm to 2000 wppm. The biorenewable feedstock may contain high oxygen content, which may be up to 10% by weight or more. The biorenewable feedstock may also contain sulfur of 1 to 500 wppm, typically less than 200 wppm.

[0033] A variety of different bio-renewable feedstocks may be suitable for process 10. The term “biorenewable feedstocks” means feedstocks other than those obtained from crude oil. Bio-renewable feedstocks may include any of those feedstocks that contain at least one of glycerides and free fatty acids. Most glycerides are triglycerides, but monoglycerides and diglycerides are also present and can be processed similarly. Free fatty acids may be obtained from phospholipids that can supply phosphorus to the feedstock. Examples of these bio-renewable raw materials include, but are not limited to, camelina oil, canola oil, corn oil, soybean oil, rapeseed oil, soybean oil, rapeseed oil, tall oil, sunflower oil, hemp seed oil, olive oil, linseed oil, coconut oil, babassu oil, castor oil, peanut oil, palm oil, mustard oil, animal fat, yellow and brown grease, lard, whale oil, milk fat, fish oil, algal oil, and sewage sludge. Additional examples of bio-renewable raw materials include non-edible vegetable oils from the group including Jatropha curcas (ratanjo, wildcaster, janglierandi), Madhuca indica (mowwa), Pongamia pinnata (kalanji, honji), calophyllum inophyllum, moringa oleifera, and Azadirachta indica (neem). Typical vegetable or animal fat triglycerides and FFAs contain aliphatic hydrocarbon chains with 8 to 30 carbon atoms in their structure. Biorenewable feedstocks may also include biomass pyrolysis oils and Fischer-Tropsch waxes. As understood, biorenewable feedstocks may include mixtures of one or more of the examples described above. Biorenewable feedstocks may be pretreated to remove contaminants and filtered to remove solids.

[0034] The hydrocarbon flow in supply line 12, after being injected together with the sulfurizer in line 15, flows from the supply surge drum 14 via a charge pump and mixes with the recirculated hydrogen treatment hydrogen flow in hydrogen treatment hydrogen line 20 to provide the mixed hydrocarbon flow in line 24. The mixed hydrocarbon flow in line 24 is mixed with the hydrogen treatment recirculation flow in recirculation line 16 to provide the hydrogen treatment charge hydrocarbon flow in hydrogen treatment charge line 26. The recirculation to supply ratio may be 1:1 to 5:1. The hydrogen treatment charge flow in line 26 can be preheated in hydrogen treatment charge exchanger 22 by heat exchange with the twice-cooled hydrogen treatment flow in twice-cooled hydrogen treatment line 32b. The heated hydrogen treatment charge hydrocarbon flow in hydrogen treatment charge line 26 can then be charged into the hydrogen treatment reactor 25. The heat exchange with the twice-cooled hydrogen treatment flow in hydrogen treatment charge exchanger 22 provides all the preheating requirements for the hydrogen treatment reactor 25. Heat exchange with the hydrogenation flow is sufficient, and a combustion heater is not required to bring the hydrogenation charge flow in line 26 to the hydrogenation reaction temperature. The hydrogenation charge flow is heated only by indirect heat exchange with other unburned flows.

[0035] The hydrogenation reactor 25 may include a protective bed reactor 27. The protective bed reaction temperature may be in the range of 246°C (475°F) to 343°C (650°F), preferably 288°C (550°F) to 304°C (580°F). The protective bed reactor 27 is operated to be low enough to prevent polymerization of olefins in the FFA, but high enough to promote the saturation, hydrogenodemetallation, hydrogenodeoxygenation, hydrogenodesulfurization, and hydrogenodenitridation reactions of the olefins. The hydrogenodeoxygenation reaction preferably minimizes hydrogenodecarbonylation and hydrogenodecarboxylation reactions in order to preserve carbon atoms on the paraffin chain.

[0036] The protective bed reactor 27 may have 1 to 5 beds of protective catalyst. In Figure 1, the protective bed reactor 27 has 3 beds of protective catalyst. The protective bed catalyst may include a base metal catalyst on a support. Base metals usable in this process include non-precious metals, nickel, chromium, molybdenum, and tungsten. Other base metals that may be used include tin, indium, germanium, lead, cobalt, gallium, and zinc. The process may also use metal sulfides, the metal in the metal sulfide being selected from one or more of the listed base metals. The hydrogenation charge stream may be charged through the base metal catalyst at a pressure of 1379 kPa (abs) (200 psia) to 13790 kPa (abs) (2000 psia). In further embodiments, the protective bed catalyst may include a second metal, the second metal comprising one or more of the following metals: tin, indium, ruthenium, rhodium, rhenium, osmium, iridium, germanium, lead, cobalt, gallium, zinc, and thallium. Alumina-supported nickel molybdenum may be a suitable catalyst in the protective bed reactor 27. A suitable protective catalyst is BGB 300, available from UOP LLC in Des Plaines, Illinois. Although the protective bed reactor 27 is shown in Figure 1, one or more protective beds, such as two, three, or more, may be included in a single hydrogenation reactor. Hydrogen quench from the hydrogen manifold 18, taken from the recirculated hydrogen stream in line 19, may be injected into the inter-bed position to control the heat generation.

[0037] The contacted hydrocarbon stream is discharged from the protected floor reactor 27 into line 28. Within the protected floor reactor 27, most of the hydrogenation demetallation and hydrogenation deoxygenation reactions will occur, accompanied by some hydrogenation denitrification and hydrogenation desulfurization. The metals removed from the biorenewable feedstock include alkali metals, alkaline earth metals, and phosphorus. The contacted hydrocarbon stream is discharged from the protected floor reactor 27 into line 28, where it will receive a hydrogen quench from the hydrogen manifold 18 and enter the hydrogenation reactor 29.

[0038] Hydrogenation reactor 25 also includes hydrogenation reactor 29. In hydrogenation reactor 29, the contacted hydrocarbon stream is brought into contact with a hydrogenation catalyst in the presence of hydrogen under hydrogenation conditions to saturate the olefin or unsaturated portions of the n-paraffin chains in the feedstock. The hydrogenation catalyst also catalyzes hydrogenation and deoxygenation reactions, including hydrogenation and decarbonylation, to remove oxygenated functional groups from hydrocarbon molecules in the biorenewable feedstock, converting the biorenewable feedstock into water and carbon oxides. The hydrogenation catalyst also catalyzes the hydrogenation and denitrification of organic sulfur and organic nitrogen in the biorenewable feedstock. Essentially, the hydrogenation reactions saturate the olefins in the feedstock by removing heteroatoms from hydrocarbons.

[0039] The hydrogenation catalyst is supplied to one, two, or more beds in a single or multiple vessel, and an inter-bed hydrogen quench flow from the hydrogen quench flow may be used. The recirculated hydrogen quench flow collected from the recirculated hydrogen line 19 into the hydrogen manifold 18 may be supplied to the hydrogenation reactor 29 for inter-bed quenching. Two hydrogenation catalyst beds 29 are shown in Figure 1, but one or more are intended.

[0040] The hydrogenation catalyst may include nickel, nickel / molybdenum, or cobalt / molybdenum dispersed on a high-surface-area support such as alumina. Other catalysts include one or more noble metals dispersed on a high-surface-area support. Non-limiting examples of noble metals include platinum and / or palladium dispersed on an alumina support such as gamma-alumina. A preferred hydrogenation catalyst is BDO 200, BDO 300, or BDO 400, available from UOP LLC in Des Plaines, Illinois. The hydrogenation reaction temperature may be in the range of 271°C (520°F) to 427°C (800°F), preferably 304°C (580°F) to 400°C (752°F). Generally, the hydrogenation conditions include a pressure of 700 kPa (100 psig) to 21 MPa (3000 psig).

[0041] The hydrogenation flow is generated in the hydrogenation line 32 from the hydrogenation reactor 29 within the hydrogenation reactor 25 and contains a hydrocarbon fraction with a considerable n-paraffin concentration. The oxygenate concentration in the hydrocarbon fraction is essentially zero, while the olefin concentration is substantially reduced compared to the flow in contact. The organic sulfur concentration in the hydrocarbon fraction may be less than 500 wppm, and the organic nitrogen concentration in the hydrocarbon fraction may be less than 10 wppm.

[0042] Since the reaction occurring in the hydrogenation reactor 25 is highly exothermic, the enthalpy in the hydrogenation flow 32 leaving the hydrogenation reactor 25 is very large, providing a particularly good opportunity for heat transfer to the reactor charge flow.

[0043] The hydrogenation flow in the hydrogenation line 32 initially flows into the combined isomerized feed exchanger 34, heating the hydrogen isomerized charge flow in the hydrogen isomerized charge line 44, providing a heated hydrogen isomerized charge flow in the heated hydrogen isomerized charge line 46, and cooling the hydrogenation flow, thereby providing a once-cooled hydrogenation flow in line 32a through indirect heat exchange. Subsequently, the once-cooled hydrogenation flow in the hydrogenation line 32a undergoes heat exchange with the combined hydrogen crackling charge flow in the combined hydrogen crackling charge line 154 within the hydrogen crackling effluent charge heat exchanger 155, heating the combined hydrogen crackling charge flow in the combined hydrogen crackling charge line 154, providing a heated hydrogen crackling charge flow in the heated hydrogen crackling charge line 156, and further cooling the once-cooled hydrogenation flow in the hydrogenation line 32a, thereby providing a twice-cooled hydrogenation flow in line 32b. Next, the twice-cooled hydrogenation flow in hydrogenation line 32b exchanges heat with the combined hydrocarbon flow in line 26 in the combined feed heat exchanger 22, as previously described, to further cool the twice-cooled hydrogenation flow in hydrogenation line 32b, providing a third-cooled hydrogenation flow in hydrogenation line 32c, and heating the hydrogenation charge flow to provide a heated hydrogenation charge flow in hydrogenation line 31. Then, the third-cooled hydrogenation flow in hydrogenation line 32 may be further cooled before separation, possibly generating steam. Sufficient heat is present in the hydrogenation flow in line 32 to heat all of the reactor charge flows 44, 154, and 26. The reactor charge flows in lines 44, 154, and 26 may be further heated by heat exchange to increase their temperature or to help raise their temperature to the reaction temperature, but intense heating, such as in a combustion heater where hydrocarbons are burned, is not required to achieve the reaction temperature. The charge flows are heated only by indirect heat exchange with other unburned flows.

[0044] The hydrogen isomerization charge flow in line 44 is heated by heat exchange with the hydrogenation treatment flow in line 32 before the hydrogen crackling charge flow in line 154 is heated by heat exchange with the hydrogenation treatment flow in line 32a with respect to the flow direction of the hydrogenation treatment flow 32. Furthermore, the hydrogen crackling charge flow in line 154 is heated by heat exchange with the hydrogenation treatment flow in line 32a before the hydrogenation treatment charge flow is heated by heat exchange with the hydrogenation treatment flow in line 32b with respect to the flow direction of the hydrogenation treatment flow.

[0045] The cooled hydrogenation flow can be separated in a hydrogenation separator 36, which may include an enhanced hot separator (EHS), with the help of a stripping gas supplied to a stripping line 39 taken from the isomerization top line 58. The hydrogenation flow is separated to provide a hydrogenation liquid flow in the hydrogenation bottom line 40, which has a lower oxygen concentration than the hydrogenation vapor flow in the hydrogenation top line 38 and the hydrogenation charge flow in line 26. The hydrogenation separator 36 may be a high-pressure stripping column. In the hydrogenation separator 36, the hydrogenation flow from the hydrogenation line 32 flows down through the column, where, by contact with the stripping gas from the stripping line 39, hydrogen, carbon dioxide, carbon monoxide, water vapor, propane, hydrogen sulfide, and phosphine, which are potential hydrogen isomerization catalyst poisons, are partially stripped. The stripping gas may include the feed-off hydrogen gas that has passed through the isomerization reactor 48, the hydrocracking reactor 150, and the hydroprocessing separator 56, as described below.

[0046] The stripping gas in the stripping line 39 enters the hydrotreatment separator 36 below the inlet for the hydrotreatment flow in the hydrotreatment line 32. The hydrotreatment separator 36 may include internal structures, such as trays or packing materials, located between the inlet for the hydrotreatment flow in line 32 and the inlet for the stripping gas in the stripping line 39, in order to facilitate the stripping of the hydrotreatment flow. The stripping gas, containing the stripped gas, enters the hydrotreatment vapor flow in the hydrotreatment top line 38 extending from the top of the hydrotreatment separator 36, together with the hydrogen isomerization liquid flow in the hydrogen isomerization bottom line 60, mixes with the cryogenic aqueous flow in the cryogenic aqueous line 63 from the boot of the cryogenic separator 62, is cooled in the cooler 64, and enters the cryogenic separator 62.

[0047] The hydrogenation separator 36 operates at 177°C (350°F) to 371°C (700°F), preferably at 204°C (400°F) to 260°C (500°F). The hydrogenation separator 36 may operate at a slightly lower pressure than the hydrogenation reactor 25, taking into account the pressure drop due to the intervening equipment. The hydrogenation separator 36 may operate at a pressure of 3.4 MPa (gauge pressure) (493 psig) to 20.4 MPa (gauge pressure) (2959 psig). The hydrogenation vapor flow in the hydrogenation separator top line 38 may have a temperature corresponding to the operating temperature of the hydrogenation separator 36.

[0048] The hydrogenation liquid flow that can be stripped collects at the bottom of the hydrogenation separator 36 and flows through the hydrogenation bottom line 40. The liquid hydrogenation flow contains diesel-range substances with a high paraffin concentration when the hydrocarbon feedstock includes biorenewable feedstock. The liquid hydrogenation flow in the hydrogenation separator bottom line 40 may be split into two flows: a hydrogen isomerization charge flow that is collected in the hydrogen isomerization charge line 42, and a recirculating hydrogenation flow that is collected in the recirculation line 16, both of which are collected from the liquid hydrogenation flow in the hydrogenation bottom line 40. The recirculating hydrogenation flow in the recirculation line 16 is pumped and combined with the combined hydrocarbon flow in line 24 to provide the hydrogenation charge flow in the hydrogenation charge line 26, as described above.

[0049] Desired products, such as transport fuels, can be supplied in the hydrogenation tower bottom line 40 because the liquid hydrogenation flow contains a higher concentration of linear paraffins. However, these products will have poor low-temperature flow characteristics and high FBP, which would deprive them of the qualification to meet jet fuel standards. Therefore, the hydrogenation liquid flow can be hydrogen-isomerized to improve low-temperature flow characteristics and reduce FBP.

[0050] The replenishment hydrogen gas in replenishment line 41 can be compressed in replenishment gas compressor 45 to provide compressed replenishment gas in compressed replenishment gas header 47. The hydrogen isomerization replenishment gas flow is taken from replenishment gas header 47 in line 43 and mixed with the hydrogen isomerization charge flow in line 42 to provide a combined hydrogen isomerization charge flow in combined hydrogen isomerization charge line 44. The combined hydrogen isomerization charge flow in combined hydrogen isomerization charge line 44 can be heated in hydrogen isomerization feed exchanger 34 by heat exchange with the hydrogenation treatment flow in hydrogenation treatment line 32, as described above, before charging the combined hydrogen isomerization charge flow into the hydrogen isomerization reactor 48, thereby bringing the combined hydrogen isomerization charge flow to the hydrogen isomerization temperature. The heat exchange with the hydrogenation treatment flow is sufficient, and a combustion heater is not required to bring the hydrogen isomerization charge flow in line 44 to the hydrogen isomerization reaction temperature. However, heat exchange may be used to bring the combined hydrogen isomerization charge flow in the combined hydrogen isomerization charge line 44 to the reaction temperature upstream of the heat exchange with the hydrogenation treatment flow in line 32 in the combined hydrogen isomerization feed exchanger 34. For example, the combined hydrogen isomerization charge flow in the combined hydrogen isomerization charge line 44 can be heat-exchanged with the hydrogen isomerization flow in the hydrogen isomerization line 50 to preheat the combined hydrogen isomerization charge flow in the hydrogen isomerization charge line 44 upstream of the combined isomerization feed exchanger 34. The hydrogen isomerization charge flow is heated only by indirect heat exchange with other unburned flows.

[0051] In the hydrogen isomerization reactor 48, hydrogen isomerization (including hydrogen dewaxing) of straight-chain hydrocarbons can be achieved on one or more beds of hydrogen isomerization catalyst, and hydrogen isomerization can be operated in parallel flow mode. Both fixed-bed trickle bed downflow mode and fixed-bed liquid-filled upflow mode are preferred.

[0052] The hydrogen isomerization catalyst comprises a dehydrogenated metal, a molecular sieve, and a metal oxide binder. The hydrogen isomerization catalyst may include a dehydrogenated metal containing a Group VIII metal. The dehydrogenated metal may be selected from platinum, palladium, nickel, nickel-molybdenum sulfide, or nickel-tungsten sulfide. Preferably, the dehydrogenated metal is selected from platinum or nickel-tungsten sulfide. The concentration of the dehydrogenated metal on the hydrogen isomerization catalyst may be 0.05 to 5% by weight, based on the transition metal.

[0053] The dehydrogenated metal is distributed between the molecular sieve and the binder, with 40-65% by weight, preferably 45-60% by weight of the metal distributed on the molecular sieve, and 40-65% by weight, preferably 45-60% by weight of the metal distributed on the binder. A relevant advantage of the hydrogen isomerization catalyst is its high activity and selectivity for hydrogen isomerization. In a further embodiment, the hydrogen isomerization catalyst further contains less than 0.5% by weight of carbon, with the relevant benefit of high activity and selectivity for hydrogen isomerization.

[0054] In one embodiment, the hydrogen isomerization catalyst comprises one or more molecular sieves having a topology selected from AEI, AEL, AFO, AFX, ATO, BEA, CHA, FAU, FER, MEL, MFI, MOR, MRE, MTT, MWW, or TON, for example, EU-2, ZSM-11, ZSM-22, ZSM-23, ZSM-48, SAPO-5, SAPO-11, SAPO-31, SAPO-34, SAPO-41, SSZ-13, SSZ-16, SSZ-39, MCM-22, zeolite Y, ferrielite, mordenite, ZSM-5, or zeolite beta, wherein the molecular sieves have the related benefit of being active in the hydrogen isomerization of linear hydrocarbons.

[0055] The metal oxide binder can be obtained from the group including alumina, silica, silica-alumina, and titania, or mixtures thereof. Preferably, the metal oxide binder is alumina, and more preferably gamma-alumina.

[0056] Hydrogen isomerization catalysts typically include particles having a diameter of 1 to 5 millimeters. Catalyst production typically involves the formation of a stable porous support, followed by impregnation with an active metal. Stable porous supports typically include metal oxides and molecular sieves which can be zeolites. Stable supports are produced with high porosity to ensure a maximum surface area, and it is typically desirable to disperse the active metal over the entire internal and external surface area of the support. DI-200, available from UOP LLC of Des Plaines, Illinois, can be a suitable hydrogen isomerization catalyst.

[0057] Hydrogen isomerization conditions generally include a temperature of 150 °C (302 °F) to 450 °C (842 °F) and a pressure of 1724 kPa (abs) (250 psia) to 13.8 MPa (abs) (2000 psia). In another embodiment, the hydrogen isomerization conditions are a temperature of 300 °C (572 °F) to 388 °C (730 °F), a pressure of 3102 kPa (abs) (450 psia) to 13790 kPa (abs) (2000 psia), a LHSV of 0.5 to 3 hr -1 of, and a hydrogen rate of 337 Nm 3 / m 3 (2,000 scf / bbl) to 2527 Nm 3 / m 3 oil (15,000 scf / bbl). The hydrogen isomerization quench gas can be taken from the quench gas manifold 57 and provided to the hydrogen isomerization reactor 48 at an interbed location.

[0058] The hydrogen isomerization flow in the hydrogen isomerization line 50 from the hydrogen isomerization reactor 48 is a branched paraffin-rich flow. Preferably, the hydrogen isomerization flow is mainly a branched paraffin flow. The hydrogen isomerization effluent is expected to contain 80, 90, or 95% by mass of branched paraffins of the total paraffin content. The hydrogen isomerization conditions in the hydrogen isomerization reactor 48 are selected to avoid undesirable decomposition, and therefore, the main product in the hydrogen isomerization flow in the hydrogen isomerization line 50 is branched paraffins. By avoiding undesirable decomposition, the hydrogen isomerization flow in the hydrogen isomerization line 50 has a composition with respect to carbon number that is close to, and only slightly less than, the composition of the hydrogen isomerization supply flow in the hydrogen isomerization charge line 42. The optimal amount of residual linear paraffins in line 50 depends on the selectivity of the hydrogen isomerization catalyst, but is typically 1 to 7% by weight.

[0059] The hydrogen isomerized flow in the hydrogen isomerization line 50 from the hydrogen isomerization reactor 48 can be mixed with the hydrocracking flow in line 152 to provide a mixed hydrogenwork flow in line 54. The mixed hydrogenwork flow in line 54 can be further cooled in the cryogenic liquid exchanger 55 by heat exchange with the cryogenic separator bottom flow in line 70 and supplied to the hydrogenwork separator 56 for separation into a liquid hydrogenwork flow and a vapor hydrogenwork flow. Internal packing may be located at the top of the hydrogenwork separator 56 to ensure that the liquid component is prevented from flowing out into the hydrogenwork top line 58. The vapor hydrogenwork flow in the hydrogenwork top line 58 extending from the top of the hydrogenwork separator 56 can be cooled and supplied to a drum to remove condensate and compressed in a compressor 59 to provide stripping gas in the stripping line 39 for the hydrogenwork separator 36 and quench gas in the quench gas manifold 57.

[0060] In one embodiment, the liquid hydrogenation flow in the hydrogenation column bottom line 60 extending from the bottom of the hydrogenation separator 56 is pumped to the cryogenic separator 62, where it can be further separated together with the vapor hydrogenation flow in line 38 and the cryogenic aqueous flow in cryogenic aqueous line 63, which is pumped around from the boot of the cryogenic separator 62. The cryogenic aqueous flow in line 63 can be combined with the vapor hydrogenation flow in line 38 and the liquid hydrogenation flow in line 60 to provide a lower temperature hydrogenation flow in line 61. The lower temperature hydrogenation flow in line 61 can be cooled in a cooler 64 and supplied to the cryogenic separator 62. The cryogenic aqueous flow in cryogenic aqueous line 63 is added to the liquid hydrogenation flow and the vapor hydrogenation flow to dissolve any salts that may be present in the liquid hydrocarbons in the cryogenic separator 62.

[0061] In the cryogenic separator 62, various components in the hydrogenation liquid flow and the vapor hydrogenation flow will be separated and rise to provide a cryogenic vapor hydrogenation flow in the cryogenic top line 68, a liquid hydrogenation flow in the cryogenic bottom line 70, and a cryogenic aqueous flow taken from the boot into the cryogenic aqueous line 63. A portion of the cryogenic aqueous flow from the boot may be used for water treatment. The cryogenic vapor hydrogenation steam in the cryogenic top line 68 may be scrubbed in the scrubber 74 to remove acidic gases and provide a scrubbed hydrogen flow in line 72. The scrubbed hydrogen flow in line 72 may be divided into a recirculating hydrogen flow in line 19 and a purge gas flow. The recirculating hydrogen flow in line 19 is compressed in a recirculating gas compressor and recirculated in the manifold line 18 to the hydrogenation reactor 25 for inter-floor quenching and to the hydrogenation hydrogen line 20 for combination with the hydrocarbon flow in the feed line 12.

[0062] The liquid fuel components in the liquid hydrogenation flow and the vapor hydrogenation flow will exit the cryogenic separator and enter the cryogenic hydrogen isomerization column bottom line 70. The strippered liquid hydrogen isomerization flow in the cryogenic hydrogenation column bottom line 70 contains diesel and fuels in the jet boiling point range, as well as other hydrocarbons such as propane and naphtha.

[0063] In one embodiment, the cryogenic liquid hydrogenation flow in the cryogenic bottom line 70 can be stripped in the stripping column 86 to remove hydrogen sulfide and other gases. The stripped liquid hydrogenation flow in the cryogenic bottom line 70 can be heated by heat exchange in the cryogenic liquid exchanger 55 with the hydrogenation flow in the hydrogenation line 54 to cool the hydrogenation flow, heat the cryogenic liquid hydrogenation flow, and supply it to the stripping column 86.

[0064] To strip light gases from the stripper liquid hydrogen isomerization flow in line 70, an inert gas stripping medium, such as water vapor, may be used from the stripping medium line 89. The stripping column 86 delivers top stripping flows of naphtha, LPG, hydrogen, hydrogen sulfide, water vapor, and other gases into the stripper top line 87, and fractionator hydrogen isomerization flows into the stripped bottom line 90. The top stripping flows in the top line 87 can be condensed and separated by cooling in the stripping receiver 95. The net stripper top line 88 from the receiver 95 can transport the net stripper top flow to the sponge absorber 140. Unstabilized liquid naphtha from the bottom of the receiver 95 can be transported through the stripper receiver bottom line 96 to the debutane unit column 170 for naphtha and LPG recovery. The acidic water stream can be collected from the boot of the tower top receiver 95.

[0065] The stripping column 86 can be operated at a top pressure of 0.35 MPa (gauge pressure) (50 psig), preferably 0.70 MPa (gauge pressure) (100 psig) or more and 2.0 MPa (gauge pressure) (290 psig) or less. The temperature inside the top receiver 95 is in the range of 38°C (100°F) to 66°C (150°F), and the pressure is essentially the same as inside the top of the stripping column 86.

[0066] The stripped hydrogenation flow in the stripper column bottom line 90 can be heated and supplied to the product fractionation column 120 to provide fractionation products. The diesel flow in the bottom line 124 is collected from the bottom of the product fractionation column 120. The hydrocracking charge flow in line 126 can be collected from the diesel flow in the bottom line 124 from the product fractionation column 120. The product fractionation column 120 can be re-boiled in a combustion heater 121 by heat exchange with a suitable high-temperature flow to provide the heat required for distillation. Alternatively, the column can be heated using a stripping medium that is an inert gas, such as steam from the stripping medium. The re-boiled flow is sent to the combustion heater 121 and returned to the product fractionation column 120 while boiling. The diesel product flow can be collected in the diesel product line 125 to the diesel pool and may be green diesel. The diesel flow in the distillation column bottom line 124 may be a diesel flow having a T5 of 230°C (446°F) to 296°C (590°F) and a T90 of 343°C (650°F) to 399°C (750°F).

[0067] The product fractionation column 120 provides a top gas stream of naphtha at the top line 122. The fractionation top stream can be completely condensed and separated from water in the distillation receiver 130. Unstabilized liquid naphtha from the bottom of the receiver 130 in the fractionation top liquid line 132 can be combined with the naphtha stream in line 176. The acidic water stream can be collected from the boot of the distillation receiver 130.

[0068] The kerosene stream can be collected from the side of the product fractionation column 120 into the side line 134. The kerosene stream collected in the side line 134 can be stripped in the kerosene stripper column 136 to displace low-boiling substances, which are returned to the product fractionation column 120 at a higher position in the top kerosene line 135. The stripped bottom kerosene stream is generated in the bottom kerosene line 137 to provide the jet fuel product stream. The jet fuel product stream in line 137 meets the jet fuel standard according to ASTM D86 and may be the green jet fuel stream collected from the bottom of the kerosene stripper column 136. The jet fuel product stream in line 137 can be cooled and transported to the jet fuel pool.

[0069] Optionally, a light diesel flow may be taken into a second side line and stripped in a side diesel stripper (not shown).

[0070] The product fractionation column 120 may be operated at a bottom temperature of 149°C (300°F) to 288°C (550°F), preferably 260°C (500°F) or less, and a top pressure of 0.35 MPa (gauge pressure) (50 psig), preferably 0.70 MPa (gauge pressure) (100 psig) or more to 2.0 MPa (gauge pressure) (290 psig) or less. The temperature in the top receiver 130 is in the range of 38°C (100°F) to 66°C (150°F), and the pressure is essentially the same as in the top of the product fractionation column 120. The product fractionation column 120 may only provide the net top flow containing jet fuel in the fractionator top liquid line 132, and it is also conceivable that naphtha and lighter flows may be taken into the fractionator receiver net top line (not shown).

[0071] The top stripping streams of naphtha, LPG, hydrogen, hydrogen sulfide, water vapor, and other gases in the stripper net top line 88 can be selectively scrubbed to remove acidic gases and passed to the sponge absorption column 140 for hydrocarbon recovery.

[0072] The sponge absorber column 140 can receive the hydrocarbon-rich flow in the stripper net top line 88. The lean absorbent flow in the lean absorbent line 142 can be supplied to the sponge absorber column 140 through the absorbent inlet. The lean absorbent may include the naphtha flow in the lean absorbent line 142, possibly from the butane bottom flow in line 176. In the sponge absorber column 140, the lean absorbent flow and the scrubbed hydrocarbon-rich flow are in countercurrent contact. The sponge absorbent absorbs LPG hydrocarbons from the net stripper gas flow into the absorbent-rich flow.

[0073] The hydrocarbons absorbed by the sponge absorbent include some methane and ethane in the net stripper gas stream, as well as LPG, the majority of C3 and C4 hydrocarbons, and any C5 and C4 hydrocarbons. 6+ It contains light naphtha hydrocarbons. The sponge absorber column 140 can operate at temperatures from 34°C (93°F) to 60°C (140°F) and at pressures essentially the same as or lower than those of the off-gas scrubbing column 140, with low frictional losses. The sponge absorber off-gas flow, depleted of LPG hydrocarbons, is withdrawn from the top of the sponge absorber column 140 through the sponge absorber top line 144 at the top outlet. The sponge absorber off-gas flow in the sponge absorber top line 144 may be transported to a fuel gas header (not shown) to provide fuel gas requirements. The rich absorbent flow, rich in LPG hydrocarbons, is withdrawn from the bottom of the sponge absorber column 140 through the rich absorber bottom line 146 at the bottom outlet and may be supplied to the debutane unit column 170 via the stripper top liquid flow in the stripper receiver bottom line 96.

[0074] In one embodiment, the butane removal column 170 primarily uses C to remove the stripper liquid top flow in line 96 and the rich absorbent flow in the rich absorption device bottom line 146. 5+The butane column bottom flow containing hydrocarbons can be fractionated into a butane column top flow containing LPG hydrocarbons. The butane column top flow in the butane column top line 172 can provide recovery of LPG in the butane column top liquid flow. The butane column bottom flow can be withdrawn from the bottom of the butane column 170 into the butane column bottom line 176. The butane column bottom flow in line 176, containing naphtha, can be supplemented with liquid naphtha from the bottom of the receiver 130 in the fractionator column top liquid line 132 and divided into a lean absorbent flow in the lean absorbent line 142 and a product naphtha flow that is cooled and sent to the gasoline pool in line 178.

[0075] The bottom flow in the bottom line 124 of the fractionation column may contain hydrocarbons in the diesel boiling point range. In one embodiment, the jet fuel flow in line 137 and the diesel flow in line 125 may be collected in a once-through manner without recirculation. The cut point in the product fractionation column 120 between the diesel flow in the bottom line 124 and the jet fuel flow in the side line 134 may be adjusted to ensure that the jet fuel flow has a suitable composition that meets jet fuel standards, in particular, jet fuel density standards, at least after blending. However, larger paraffins are concentrated in the bottom flow of the fractionation column and are therefore very suitable for hydrocracking to kerosene-range hydrocarbons.

[0076] In an optional embodiment, the hydrocracking charge flow in the hydrocracking charge line 126 can be used to charge the hydrocracking reactor 150. The hydrocracking reactor 150 is located downstream of the hydrogen isomerization reactor 48 and the hydrogenation reactor 25. The hydrocracking charge flow can be mixed with the hydrocracking hydrogen flow in line 52, which is taken from the compressed feed hydrogen flow in the compressed feed gas header 47, to provide a combined hydrocracking charge flow in the combined hydrocracking charge line 154. The combined hydrocracking charge flow can be heated in the hydrocracking effluent charge exchanger 155 by heat exchange with the once-cooled hydrogenation flow in line 32a to provide a twice-cooled hydrogenation flow in line 32b and a heated hydrocracking charge flow in the heated hydrocracking charge line 156, which is used to charge the hydrocracking reactor 150. As explained above, the heat exchange in the hydrocracking effluent charge exchanger 155 is sufficient to bring the combined hydrocracking charge stream to the hydrocracking reaction temperature before charging the hydrocracking reactor 150. The heat exchange with the hydrotreatment stream is sufficient, and a combustion heater is not required to bring the hydrocracking charge stream in line 154 to the hydrocracking reaction temperature. However, heat exchange may be used to bring the combined hydrocracking charge stream in the combined hydrocracking charge line 154 to the reaction temperature upstream of the heat exchange with the previously cooled hydrotreatment stream in line 32a within the hydrocracking effluent charge exchanger 155. For example, the combined hydrocracking charge stream in the combined hydrocracking charge line 44 can be heat-exchanged with the hydrocracking stream in the hydrocracking line 152 to preheat the combined hydrocracking charge stream in the hydrocracking charge line 154 upstream of the hydrocracking effluent charge exchanger 155. The hydrocracking charge flow is heated only through indirect heat exchange with other unburned flows.

[0077] The hydrocracking reactor 150 may be a fixed-bed reactor comprising one or more vessels, one or more catalyst beds in each vessel, and various combinations of hydrocracking catalysts in one or more vessels. The hydrocracking reactor 150 may operate within a conventional open-gas, moving-bed, or fluidized-bed hydrocracking reactor.

[0078] The combined hydrocracking flow is hydrocracked on a hydrocracking catalyst in the hydrocracking reactor 150 in the presence of a hydrocracking hydrogen flow from the hydrocracking hydrogen line 52 to provide a hydrocracking flow. Quench gas collected from the quench gas manifold 57 can be supplied to the hydrocracking reactor 150 at a floor-level position.

[0079] The hydrocracking reactor can provide a total conversion rate of at least 20 vol%, typically over 60 vol%, of the hydrocracking charge stream in the heated hydrocracking charge line 156 to a product that boils below the heavy diesel range of 293°C (560°F) to 310°C (590°F). The hydrocracking reactor 150 can be operated with a partial conversion rate of over 30 vol%, or a complete conversion rate of at least 90 vol%, of the feed, based on the total conversion rate. The hydrocracking reactor 150 can be operated under mild hydrocracking conditions that provide a total conversion rate of 20–60 vol%, preferably 20–50 vol%, of the hydrocracking charge stream to a product that boils below the heavy diesel boiling point range.

[0080] Hydrocracking catalysts can selectively produce a balance between light diesel and jet fuel distillates by utilizing an amorphous silica-alumina base or zeolite base combined with one or more Group VIII or Group VIB metal hydrogenation components. In another embodiment, catalysts generally comprising any crystalline zeolite cracking base on which Group VIII metal hydrogenation components are deposited may be preferred. Additional hydrogenation components may be selected from Group VIB for incorporation into the zeolite base. Furthermore, the hydrogen isomerization catalyst from the hydrogen isomerization reactor 48 may be used as a hydrogencracking catalyst in the hydrogencracking reactor 150, but may be operated at the upper limit of the hydrogen isomerization temperature range.

[0081] Zeolite decomposition bases, sometimes referred to as molecular sieves in the art, typically consist of silica, alumina, and one or more interchangeable cations such as sodium, magnesium, calcium, and rare earth metals. They are further characterized by crystalline pores with relatively uniform diameters of 4 to 14 angstroms. It is preferable to use zeolites having a relatively high silica / alumina molar ratio of 3 to 12. Suitable naturally occurring zeolites include, for example, mordenite, stilbite, heulandite, ferrielite, dachyaldite, chabazite, erionite, and faujasite. Suitable synthetic zeolites include, for example, B, X, Y, and L crystal forms, such as synthetic faujasite and mordenite. Preferred zeolites have crystalline pore diameters of 8 to 12 angstroms and a silica / alumina molar ratio of 4 to 6. An example of a zeolite in the preferred group is synthetic Y molecular sieves.

[0082] Naturally occurring zeolites are typically found in sodium form, alkaline earth metal form, or mixed form. Synthetic zeolites are almost always prepared in sodium form. In any case, for use as a decomposition base, it is preferable to ion-exchange most or all of the original zeolite monovalent metal with polyvalent metal and / or ammonium salt, followed by heating to decompose the ammonium ions associated with the zeolite, leaving hydrogen ions and / or exchange sites at those sites, which are actually decationized by further removal of water. This type of hydrogen or "decationized" Y zeolite is described in more detail in U.S. Patent No. 3,100,006.

[0083] Mixed polyvalent metal-hydrogen zeolites can be prepared by ion exchange with an ammonium salt, then by partial reverse exchange with a polyvalent metal salt, and then by calcination. In some cases, as in the case of synthetic mordenite, the hydrogen form can be prepared by direct acid treatment of an alkali metal zeolite. In one embodiment, a preferred decomposition base is one that is deficient in at least 10% by weight, preferably at least 20% by weight, of metal cations, based on the initial ion exchange capacity. In another embodiment, a desirable and stable class of zeolite is one in which at least 20% by weight of the ion exchange capacity is filled with hydrogen ions.

[0084] The active metals used as hydrogenation components in the preferred hydrocracking catalysts of this disclosure are those of Group VIII, namely iron, cobalt, nickel, ruthenium, rhodium, palladium, osmium, iridium, and platinum. In addition to these metals, other accelerators, including Group VIB metals such as molybdenum and tungsten, may also be used. The amount of metal hydride in the catalyst can vary over a wide range. Generally, any amount between 0.05% and 30% by weight can be used. In the case of noble metals, it is usually preferable to use 0.05% to 2% by weight of the noble metal. Noble metals may deactivate the noble metal catalyst, but they may be preferred as metal hydride on the hydrocracking catalyst to provide selectivity for jet fuel because hydrogen sulfide and ammonia, which are removed upstream of the process, are not present.

[0085] A method for incorporating metal hydrides involves contacting a base material with an aqueous solution of a suitable compound of the desired metal, in which the metal exists in a cation form. After adding one or more selected metal hydrides, the resulting catalyst powder is filtered, dried, pelletized with the addition of lubricants, binders, etc. as needed, and calcined in air at temperatures such as 371°C (700°F) to 648°C (200°F) to activate the catalyst and decompose ammonium ions. Alternatively, the base component may be pelletized, followed by the addition of the hydride component and activation by calcination.

[0086] The catalysts described above may be used in an undiluted form, or the powder catalyst may be mixed and co-pelleted with other relatively less active catalysts, diluents, or binders in a proportion ranging from 5 to 90% by weight, such as alumina, silica gel, silica-alumina cogel, or activated clay. These diluents may be used as is, or may contain small amounts of added metal hydrides, such as Group VIB and / or Group VIII metals. Additional metal-accelerated hydrocracking catalysts, including, for example, aluminophosphate molecular sieves, crystalline chromosilicates, and other crystalline silicates, may also be used in the processes of this disclosure. Crystalline chromosilicates are described more fully in U.S. Patent No. 4,363,178.

[0087] According to one approach, the hydrocracking conditions are a temperature of 290°C (550°F) to 468°C (875°F), preferably 300°C (572°F) to 445°C (833°F), a pressure of 2.7 MPa (gauge pressure) (400 psig) to 20.7 MPa (gauge pressure) (3000 psig), and a duration of 0.4 to 2.5 hours. -1 Liquid hourly space velocity (LHSV) less than 337 Nm 3 / m 3 Oil (2,000scf / bbl)~2,527Nm 3 / m 3 This may include the hydrogen rate of oil (15,000 scf / bbl).

[0088] The hydrocracking flow can exit the hydrocracking reactor 150 and enter the hydrocracking line 152. The hydrocracking flow in line 152 is combined with the hydrogen isomerization flow in line 50 to provide the hydrogen processing flow in line 54, which is processed as described above.

[0089] In the process described above, the entire reactor charge flow is heated to the reaction temperature through heat exchange with the hydrogenation flow in line 32. As a result, a combustion heater is not required to heat the pressurized reactor charge flow. [Examples]

[0090] The inventors simulated a disclosed process that does not use a combustion heater to preheat the feed to any of the reactors. The loads for each heater or exchanger are provided in the table below. In the table, the reference numbers for the elements in the figure are provided in parentheses for each heater or exchanger in the figure.

[0091] [Table 1]

[0092] Simulations reveal that all the duty cycles for preheating the feeds to the isomerization reactor 48, the hydrogenation reactor 25, and the hydrocracking reactor 150 are provided by heat exchange with the hydrogenation flow in line 32. The disclosed process eliminates the need for combustion heaters typically required in conventional processes, such as hydrogenation feed heaters and isomerization feed heaters.

[0093] Specific Embodiments The following will be explained in conjunction with specific embodiments, but it should be understood that this explanation is intended to illustrate the scope of the preceding explanation and the attached claims, and is not intended to limit them.

[0094] A first embodiment of the present invention is a process for hydrogenating a hydrocarbon flow, comprising: heating a hydrogenation charge flow by heat exchange with a hydrogenation flow to provide a heated hydrogenation charge flow; hydrogenating the heated hydrogenation charge flow on a hydrogenation catalyst in the presence of hydrogen to provide a hydrogenation flow; heating a hydrogen isomerization charge flow by heat exchange with a hydrogenation flow to provide a heated hydrogen isomerization charge flow; and hydrogen isomerizing the heated hydrogen isomerization charge flow on a hydrogen isomerization catalyst in the presence of hydrogen to provide a hydrogen isomerization flow. Embodiments of the present invention are any or all of the prior embodiments in this paragraph leading up to the first embodiment in this paragraph, further comprising: heating a hydrocracking charge flow by heat exchange with a hydrogenation flow to provide a heated hydrocracking charge flow; and hydrocracking the heated hydrocracking charge flow in the presence of hydrogen to provide a hydrocracking flow. Embodiments of the present invention are any or all of the prior embodiments in this paragraph leading up to the first embodiment in this paragraph, further comprising heating the hydrogenation charge flow without heating the hydrogenation charge flow in a combustion heater. Embodiments of the present invention are any or all of the prior embodiments in this paragraph leading up to the first embodiment in this paragraph, further comprising heating the hydrogenation charge flow solely by heat exchange with the hydrogenation charge flow. Embodiments of the present invention are any or all of the prior embodiments in this paragraph leading up to the first embodiment in this paragraph, further comprising heating the hydrogen isomerization charge flow without heating the hydrogen isomerization charge flow in a combustion heater. Embodiments of the present invention are any or all of the prior embodiments in this paragraph leading up to the first embodiment in this paragraph, further comprising heating the hydrogen isomerization charge flow solely by heat exchange with another flow. Embodiments of the present invention are one, any, or all of the prior embodiments described in this paragraph, up to the first embodiment described herein, further comprising heating the hydrocracking charge flow without heating the hydrocracking charge flow in a combustion heater.Embodiments of the present invention are any or all of the prior embodiments in this paragraph leading up to the first embodiment in this paragraph, further comprising heating the hydrocracking charge flow only by heat exchange with another flow. Embodiments of the present invention are any or all of the prior embodiments in this paragraph leading up to the first embodiment in this paragraph, with respect to the hydrocracking flow, in which the hydrogen isomerization charge flow is heated by heat exchange with the hydrocracking flow before the hydrocracking charge flow is heated by heat exchange with the hydrocracking flow. Embodiments of the present invention are any or all of the prior embodiments in this paragraph leading up to the first embodiment in this paragraph, with respect to the hydrocracking flow, in which the hydrogen isomerization charge flow is heated by heat exchange with the hydrocracking flow before the hydrocracking charge flow is heated by heat exchange with the hydrocracking flow. Embodiments of the present invention are any or all of the prior embodiments in this paragraph leading up to the first embodiment in this paragraph, in which, with respect to a hydrogenation flow, the hydrogenocracked charge flow is heated by heat exchange with the hydrogenation flow before the hydrogenation charge flow is heated by heat exchange with the hydrogenation flow. Embodiments of the present invention are any or all of the prior embodiments in this paragraph leading up to the first embodiment in this paragraph, in which the hydrogenation of the heated hydrogenation charge flow includes hydrogenation of the hydrogenation charge flow in a protective reactor and in a hydrogenation reactor. Embodiments of the present invention are any or all of the prior embodiments in this paragraph leading up to the first embodiment in this paragraph, in which the hydrogenation charge flow is a fresh hydrocarbon flow. Embodiments of the present invention are any or all of the prior embodiments in this paragraph leading up to the first embodiment in this paragraph, in which the hydrogen isomerized charge flow is taken from the hydrogenation flow. Embodiments of the present invention are any or all of the prior embodiments described in this paragraph, up to the first embodiment described herein, in which the hydrocracking charge stream is taken from the hydrogen isomerization stream.

[0095] A second embodiment of the present invention is a process for hydrogenating a hydrocarbon flow, comprising: heating a hydrogenation charge flow by heat exchange with a hydrogenation flow to provide a heated hydrogenation charge flow; hydrogenating the heated hydrogenation charge flow on a hydrogenation catalyst in the presence of hydrogen to provide a hydrogenation flow; heating a hydrogen isomerization charge flow by heat exchange with a hydrogenation flow to provide a heated hydrogen isomerization charge flow; hydrogen isomerizing the heated hydrogen isomerization charge flow on a hydrogen isomerization catalyst in the presence of hydrogen to provide a hydrogen isomerization flow; heating a hydrocracking charge flow by heat exchange with a hydrogenation flow to provide a heated hydrocracking charge flow; and hydrocracking the heated hydrocracking charge flow in the presence of hydrogen to provide a hydrocracking flow. Embodiments of the present invention are one, any, or all of the prior embodiments described in this paragraph, up to the second embodiment described in this paragraph, further comprising heating the hydrogenation charge flow, the hydrogen isomerization charge flow, and the hydrocracking charge flow solely by heat exchange with the hydrogenation charge flow.

[0096] A third embodiment of the present invention is a process for hydrogenating a hydrocarbon flow, comprising: heating a hydrogenation charge flow by heat exchange with a hydrogenation flow to provide a heated hydrogenation charge flow; hydrogenating the heated hydrogenation charge flow on a hydrogenation catalyst in the presence of hydrogen to provide a hydrogenation flow; collecting a hydrogen isomerized charge flow from the hydrogenation flow; heating the hydrogen isomerized charge flow by heat exchange with a hydrogenation flow to provide a heated hydrogen isomerized charge flow; and hydrogen isomerizing the heated hydrogen isomerized charge flow on a hydrogen isomerization catalyst in the presence of hydrogen to provide a hydrogen isomerized flow. Embodiments of the present invention are any or all of the prior embodiments in this paragraph, up to the third embodiment in this paragraph, further comprising: taking a hydrocracking charge flow from a hydrogen isomerization flow; heating the hydrocracking charge flow by heat exchange with a hydrogenation treatment flow to provide a heated hydrocracking charge flow; and hydrocracking the heated hydrocracking charge flow in the presence of hydrogen to provide a hydrocracking flow. Embodiments of the present invention are any or all of the prior embodiments in this paragraph, up to the third embodiment in this paragraph, with respect to the hydrogenation treatment flow, in which the hydrogen isomerization charge flow is heated by heat exchange with the hydrogenation treatment flow before the hydrocracking charge flow is heated by heat exchange with the hydrogenation treatment flow, and the hydrocracking charge flow is heated by heat exchange with the hydrogenation treatment flow before the hydrogenation treatment charge flow is heated by heat exchange with the hydrogenation treatment flow.

[0097] Without further detail, it is expected that those skilled in the art will be able to utilize the Disclosure to the fullest extent without departing from the spirit and scope of the Disclosure, readily identify its essential characteristics, and make various changes and modifications to adapt it to various uses and conditions. Accordingly, the prior preferred specific embodiments should be interpreted as merely illustrative examples and not in any way limiting the remainder of the Disclosure, but are intended to cover various modifications and equivalent configurations that fall within the scope of the appended claims.

[0098] In the above, all temperatures are given in degrees Celsius, and all parts and percentages are based on weight unless otherwise indicated.

Claims

1. A process for hydrogenating hydrocarbon flows, The hydrogenation treatment charge flow is heated by heat exchange with the hydrogenation treatment flow, thereby providing a heated hydrogenation treatment charge flow. The heated hydrogenation treatment charge flow is hydrogenated on a hydrogenation treatment catalyst in the presence of hydrogen to provide the hydrogenation treatment flow. The hydrogen isomerization charge flow is heated by heat exchange with the aforementioned hydrogenation treatment flow, thereby providing a heated hydrogen isomerization charge flow. A process comprising hydrogen isomerizing the heated hydrogen isomerized charge flow on a hydrogen isomerization catalyst in the presence of hydrogen to provide a hydrogen isomerized flow.

2. The process according to claim 1, further comprising: heating a hydrocracking charge flow by heat exchange with the hydrogenation treatment flow to provide a heated hydrocracking charge flow; and hydrocracking the heated hydrocracking charge flow in the presence of hydrogen to provide a hydrocracking flow.

3. The process according to claim 1, further comprising heating the hydrogenation charge flow without heating the hydrogenation charge flow in a combustion heater.

4. The process according to claim 1, further comprising heating the hydrogenation charge flow solely by heat exchange with the hydrogenation flow.

5. The process according to claim 1, further comprising heating the hydrogen isomerized charge flow without heating the hydrogen isomerized charge flow in a combustion heater.

6. The process according to claim 1, further comprising heating the hydrogen isomerized charge flow solely by heat exchange with another flow.

7. The process according to claim 1, further comprising heating the hydrocracking charge flow without heating the hydrocracking charge flow in a combustion heater.

8. The process according to claim 1, further comprising heating the hydrocracking charge flow solely by heat exchange with another flow.

9. The process according to claim 1, wherein, with respect to the hydrogenation treatment flow, the hydrogen isomerization charge flow is heated by heat exchange with the hydrogenation treatment flow before the hydrogenation treatment charge flow is heated by heat exchange with the hydrogenation treatment flow.

10. The process according to claim 2, wherein, with respect to the hydrogenation treatment flow, the hydrogen isomerization charge flow is heated by heat exchange with the hydrogenation treatment flow before the hydrogen cracking charge flow is heated by heat exchange with the hydrogenation treatment flow.

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